https://processdesign.mccormick.northwestern.edu/api.php?action=feedcontributions&user=WFu&feedformat=atomprocessdesign - User contributions [en]2024-03-29T14:59:31ZUser contributionsMediaWiki 1.39.2https://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5124Desalination - Team G2016-03-11T18:16:36Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> ASPEN+ simulation <div>]]<br />
<div align=left><br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A: Block Flow Diagram==<br />
[[File:Watson appendixA.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix B: Process flow diagram==<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix C: Aspen Model===<br />
[[File:Watson_aspen.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix D: Equipment and Utility Costs===<br />
[[File:Watson_utility_appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
[[File:Watson equip appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix E: Seawater Composition===<br />
[[File:Watson_seawater_appE.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix F: Calculations for Projected Production===<br />
[[File:Watson_appF.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix G: Stream Table===<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Stream Table based on ASPEN model'''<br />
! Date:<br />
!Temperature (C) <br />
!Pressure (bar) <br />
!Vapor Frac <br />
!Solid Frac <br />
!Mole Flow (kmol/hr) <br />
!Mass Flow (kg/hr) <br />
!Volume Flow (cum/hr) <br />
!Enthalpy (Gcal/hr) <br />
!Mass Flow H2O (kg/hr) <br />
!Mass Flow NACL (kg/hr) <br />
!Mole Flow H2O (kmol/hr) <br />
!Mole Flow NACL (kmol/hr)<br />
|-<br />
| BR1<br />
| 94.9<br />
| 0.796<br />
| 0<br />
| 0<br />
| 947915.803<br />
| 19242000<br />
| 24980.85<br />
| -64524.976<br />
| 16112200<br />
| 3129820<br />
| 894361.972<br />
| 53553.831<br />
|-<br />
| BR2<br />
| 91.9<br />
| 0.71<br />
| 0<br />
| 0<br />
| 942373.503<br />
| 19142200<br />
| 24834.029<br />
| -64207.75<br />
| 16012300<br />
| 3129820<br />
| 888819.672<br />
| 53553.831<br />
|-<br />
| BR3<br />
| 88.8<br />
| 0.632<br />
| 0<br />
| 0<br />
| 936866.23<br />
| 19042900<br />
| 24689.08<br />
| -63892.397<br />
| 15913100<br />
| 3129820<br />
| 883312.4<br />
| 53553.831<br />
|-<br />
| BR4<br />
| 85.8<br />
| 0.562<br />
| 0<br />
| 0<br />
| 931429.146<br />
| 18945000<br />
| 24546.917<br />
| -63580.935<br />
| 15815200<br />
| 3129820<br />
| 877875.315<br />
| 53553.831<br />
|-<br />
| BR5<br />
| 82.7<br />
| 0.498<br />
| 0<br />
| 0<br />
| 926025.475<br />
| 18847600<br />
| 24406.569<br />
| -63271.259<br />
| 15717800<br />
| 3129820<br />
| 872471.644<br />
| 53553.831<br />
|-<br />
| BR6<br />
| 79.7<br />
| 0.441<br />
| 0<br />
| 0<br />
| 920689.773<br />
| 18751500<br />
| 24268.92<br />
| -62965.351<br />
| 15621700<br />
| 3129820<br />
| 867135.942<br />
| 53553.831<br />
|-<br />
| BR7<br />
| 76.6<br />
| 0.389<br />
| 0<br />
| 0<br />
| 915386.001<br />
| 18656000<br />
| 24133.036<br />
| -62661.146<br />
| 15526200<br />
| 3129820<br />
| 861832.17<br />
| 53553.831<br />
|-<br />
| BR8<br />
| 73.6<br />
| 0.342<br />
| 0<br />
| 0<br />
| 910148.127<br />
| 18561600<br />
| 23999.778<br />
| -62360.597<br />
| 15431800<br />
| 3129820<br />
| 856594.296<br />
| 53553.831<br />
|-<br />
| BRINE<br />
| 70.8<br />
| 0.304<br />
| 0<br />
| 0<br />
| 905441.542<br />
| 18476800<br />
| 23880.836<br />
| -62090.445<br />
| 15347000<br />
| 3129820<br />
| 851887.711<br />
| 53553.831<br />
|-<br />
| D1<br />
| 94.9<br />
| 0.796<br />
| 1<br />
| 0<br />
| 3306.022<br />
| 59558.904<br />
| 126413.546<br />
| -189.255<br />
| 59558.904<br />
| 0<br />
| 3306.022<br />
| 0<br />
|-<br />
| D2<br />
| 91.9<br />
| 0.71<br />
| 1<br />
| 0<br />
| 5542.299<br />
| 99846.076<br />
| 235561.239<br />
| -317.403<br />
| 99846.076<br />
| 0<br />
| 5542.299<br />
| 0<br />
|-<br />
| D3<br />
| 88.8<br />
| 0.632<br />
| 1<br />
| 0<br />
| 5507.273<br />
| 99215.065<br />
| 260816.407<br />
| -315.529<br />
| 99215.065<br />
| 0<br />
| 5507.273<br />
| 0<br />
|-<br />
| D4 <br />
| 85.8<br />
| 0.562<br />
| 1<br />
| 0<br />
| 5437.085<br />
| 97950.605<br />
| 287426.924<br />
| -311.637<br />
| 97950.605<br />
| 0<br />
| 5437.085<br />
| 0<br />
|-<br />
| D5<br />
| 82.7<br />
| 0.498<br />
| 1<br />
| 0<br />
| 5403.671<br />
| 97348.64<br />
| 319693.064<br />
| -309.852<br />
| 97348.64<br />
| 0<br />
| 5403.671<br />
| 0<br />
|-<br />
| D6<br />
| 79.7<br />
| 0.441<br />
| 1<br />
| 0<br />
| 5335.702<br />
| 96124.173<br />
| 353955.3<br />
| -306.082<br />
| 96124.173<br />
| 0<br />
| 5335.702<br />
| 0<br />
|-<br />
| D7<br />
| 76.6<br />
| 0.389<br />
| 1<br />
| 0<br />
| 5303.771<br />
| 95548.926<br />
| 395584.927<br />
| -304.379<br />
| 95548.926<br />
| 0<br />
| 5303.771<br />
| 0<br />
|-<br />
| D8<br />
| 73.6<br />
| 0.342<br />
| 1<br />
| 0<br />
| 5237.874<br />
| 94361.766<br />
| 440141.36<br />
| -300.723<br />
| 94361.766<br />
| 0<br />
| 5237.874<br />
| 0<br />
|-<br />
| FEED<br />
| 11<br />
| 1.013<br />
| 0<br />
| 0<br />
| 56107.895<br />
| 1032950<br />
| 1083.597<br />
| -3857.463<br />
| 1000930<br />
| 32021.357<br />
| 55559.983<br />
| 547.912<br />
|-<br />
| HOTFEED<br />
| 57.7<br />
| 1.013<br />
| 0<br />
| 0<br />
| 56107.895<br />
| 1032950<br />
| 1100.165<br />
| -3810.03<br />
| 1000930<br />
| 32021.357<br />
| 55559.983<br />
| 547.912<br />
|-<br />
| INPUT<br />
| 98<br />
| 1.013<br />
| 0<br />
| 0<br />
| 951221.824<br />
| 19301600<br />
| 25084.412<br />
| -64714.056<br />
| 16171700<br />
| 3129820<br />
| 897667.993<br />
| 53553.831<br />
|-<br />
| M1<br />
| 70.7<br />
| 0.304<br />
| 0.004<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 402879.11<br />
| -65216.53<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| M2<br />
| 70.7<br />
| 0.304<br />
| 0.01<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 897392.279<br />
| -65163.387<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| M3<br />
| 70.7<br />
| 0.304<br />
| 0.015<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 1393370<br />
| -65110.087<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| M4<br />
| 70.8<br />
| 0.304<br />
| 0.021<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 1894150<br />
| -65056.274<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| M5<br />
| 70.8<br />
| 0.304<br />
| 0.027<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 2396470<br />
| -65002.297<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| M6<br />
| 70.8<br />
| 0.304<br />
| 0.032<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 2903710<br />
| -64947.793<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| M7<br />
| 70.8<br />
| 0.304<br />
| 0.038<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 3412590<br />
| -64893.114<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| MIXED<br />
| 70.1<br />
| 0.304<br />
| 0<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 24740.57<br />
| -65269.173<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| P1<br />
| 93.4<br />
| 0.796<br />
| 0<br />
| 0<br />
| 3306.022<br />
| 59558.904<br />
| 61.859<br />
| -221.768<br />
| 59558.904<br />
| 0<br />
| 3306.022<br />
| 0<br />
|-<br />
| P2<br />
| 90.3<br />
| 0.71<br />
| 0<br />
| 0<br />
| 5542.299<br />
| 99846.076<br />
| 103.478<br />
| -372.082<br />
| 99846.076<br />
| 0<br />
| 5542.299<br />
| 0<br />
|-<br />
| P3<br />
| 87.3<br />
| 0.632<br />
| 0<br />
| 0<br />
| 5507.273<br />
| 99215.065<br />
| 102.608<br />
| -370.033<br />
| 99215.065<br />
| 0<br />
| 5507.273<br />
| 0<br />
|-<br />
| P4<br />
| 84.3<br />
| 0.562<br />
| 0<br />
| 0<br />
| 5437.085<br />
| 97950.605<br />
| 101.092<br />
| -365.615<br />
| 97950.605<br />
| 0<br />
| 5437.085<br />
| 0<br />
|-<br />
| P5<br />
| 81.2<br />
| 0.498<br />
| 0<br />
| 0<br />
| 5403.671<br />
| 97348.64<br />
| 100.269<br />
| -363.664<br />
| 97348.64<br />
| 0<br />
| 5403.671<br />
| 0<br />
|-<br />
| P6<br />
| 78.2<br />
| 0.441<br />
| 0<br />
| 0<br />
| 5335.702<br />
| 96124.173<br />
| 98.815<br />
| -359.382<br />
| 96124.173<br />
| 0<br />
| 5335.702<br />
| 0<br />
|-<br />
| P7<br />
| 75.2<br />
| 0.389<br />
| 0<br />
| 0<br />
| 5303.771<br />
| 95548.926<br />
| 98.036<br />
| -357.522<br />
| 95548.926<br />
| 0<br />
| 5303.771<br />
| 0<br />
|-<br />
| P8<br />
| 72.1<br />
| 0.342<br />
| 0<br />
| 0<br />
| 5237.874<br />
| 94361.766<br />
| 96.639<br />
| -353.366<br />
| 94361.766<br />
| 0<br />
| 5237.874<br />
| 0<br />
|-<br />
| P9<br />
| 69.4<br />
| 0.304<br />
| 0<br />
| 0<br />
| 4706.586<br />
| 84790.457<br />
| 86.695<br />
| -317.757<br />
| 84790.457<br />
| 0<br />
| 4706.586<br />
| 0<br />
|-<br />
| PREHEAT<br />
| 70.8<br />
| 0.304<br />
| 0.041<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 3716880<br />
| -64860.601<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
|-<br />
| PURE<br />
| 72.1<br />
| 0.342<br />
| 0.019<br />
| 0<br />
| 35737.995<br />
| 643829.982<br />
| 58727.526<br />
| -2404.05<br />
| 643829.982<br />
| 0<br />
| 35737.995<br />
| 0<br />
|-<br />
| TANK<br />
| 70.1<br />
| 0.304<br />
| 0<br />
| 0<br />
| 961549.437<br />
| 19509800<br />
| 24979.868<br />
| -65900.475<br />
| 16347900<br />
| 3161840<br />
| 907447.693<br />
| 54101.743<br />
|-<br />
| WASTE<br />
| 70.1<br />
| 0.304<br />
| 0<br />
| 0<br />
| 9211.272<br />
| 186896<br />
| 239.297<br />
| -631.301<br />
| 156606.799<br />
| 30289.201<br />
| 8692.998<br />
| 518.274<br />
|}<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5123Desalination - Team G2016-03-11T18:14:19Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> ASPEN+ simulation <div>]]<br />
<div align=left><br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A: Block Flow Diagram==<br />
[[File:Watson appendixA.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix B: Process flow diagram==<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix C: Aspen Model===<br />
[[File:Watson_aspen.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix D: Equipment and Utility Costs===<br />
[[File:Watson_utility_appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
[[File:Watson equip appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix E: Seawater Composition===<br />
[[File:Watson_seawater_appE.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix F: Calculations for Projected Production===<br />
[[File:Watson_appF.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix G: Stream Table===<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Stream Table based on ASPEN model'''<br />
! Date:<br />
!Temperature (C) <br />
!Pressure (bar) <br />
!Vapor Frac <br />
!Solid Frac <br />
!Mole Flow (kmol/hr) <br />
!Mass Flow (kg/hr) <br />
!Volume Flow (cum/hr) <br />
!Enthalpy (Gcal/hr) <br />
!Mass Flow H2O (kg/hr) <br />
!Mass Flow NACL (kg/hr) <br />
!Mole Flow H2O (kmol/hr) <br />
!Mole Flow NACL (kmol/hr)<br />
|-<br />
| BR1<br />
| 94.9<br />
| 0.796<br />
| 0<br />
| 0<br />
| 947915.803<br />
| 19242000<br />
| 24980.85<br />
| -64524.976<br />
| 16112200<br />
| 3129820<br />
| 894361.972<br />
| 53553.831<br />
|-<br />
| BR2<br />
| 91.9<br />
| 0.71<br />
| 0<br />
| 0<br />
| 942373.503<br />
| 19142200<br />
| 24834.029<br />
| -64207.75<br />
| 16012300<br />
| 3129820<br />
| 888819.672<br />
| 53553.831<br />
| -<br />
|-<br />
| BR3<br />
| 88.8<br />
| 0.632<br />
| 0<br />
| 0<br />
| 936866.23<br />
| 19042900<br />
| 24689.08<br />
| -63892.397<br />
| 15913100<br />
| 3129820<br />
| 883312.4<br />
| 53553.831<br />
| –<br />
|-<br />
| BR4<br />
| 85.8<br />
| 0.562<br />
| 0<br />
| 0<br />
| 931429.146<br />
| 18945000<br />
| 24546.917<br />
| -63580.935<br />
| 15815200<br />
| 3129820<br />
| 877875.315<br />
| 53553.831<br />
| –<br />
|-<br />
| BR5<br />
| 82.7<br />
| 0.498<br />
| 0<br />
| 0<br />
| 926025.475<br />
| 18847600<br />
| 24406.569<br />
| -63271.259<br />
| 15717800<br />
| 3129820<br />
| 872471.644<br />
| 53553.831<br />
| –<br />
|-<br />
| BR6<br />
| 79.7<br />
| 0.441<br />
| 0<br />
| 0<br />
| 920689.773<br />
| 18751500<br />
| 24268.92<br />
| -62965.351<br />
| 15621700<br />
| 3129820<br />
| 867135.942<br />
| 53553.831<br />
| –<br />
|-<br />
| BR7<br />
| 76.6<br />
| 0.389<br />
| 0<br />
| 0<br />
| 915386.001<br />
| 18656000<br />
| 24133.036<br />
| -62661.146<br />
| 15526200<br />
| 3129820<br />
| 861832.17<br />
| 53553.831<br />
| –<br />
|-<br />
| BR8<br />
| 73.6<br />
| 0.342<br />
| 0<br />
| 0<br />
| 910148.127<br />
| 18561600<br />
| 23999.778<br />
| -62360.597<br />
| 15431800<br />
| 3129820<br />
| 856594.296<br />
| 53553.831<br />
| –<br />
|-<br />
| BRINE<br />
| 70.8<br />
| 0.304<br />
| 0<br />
| 0<br />
| 905441.542<br />
| 18476800<br />
| 23880.836<br />
| -62090.445<br />
| 15347000<br />
| 3129820<br />
| 851887.711<br />
| 53553.831<br />
| –<br />
|-<br />
| D1<br />
| 94.9<br />
| 0.796<br />
| 1<br />
| 0<br />
| 3306.022<br />
| 59558.904<br />
| 126413.546<br />
| -189.255<br />
| 59558.904<br />
| 0<br />
| 3306.022<br />
| 0<br />
| –<br />
|-<br />
| D2<br />
| 91.9<br />
| 0.71<br />
| 1<br />
| 0<br />
| 5542.299<br />
| 99846.076<br />
| 235561.239<br />
| -317.403<br />
| 99846.076<br />
| 0<br />
| 5542.299<br />
| 0<br />
| –<br />
|-<br />
| D3<br />
| 88.8<br />
| 0.632<br />
| 1<br />
| 0<br />
| 5507.273<br />
| 99215.065<br />
| 260816.407<br />
| -315.529<br />
| 99215.065<br />
| 0<br />
| 5507.273<br />
| 0<br />
| –<br />
|-<br />
| D4 <br />
| 85.8<br />
| 0.562<br />
| 1<br />
| 0<br />
| 5437.085<br />
| 97950.605<br />
| 287426.924<br />
| -311.637<br />
| 97950.605<br />
| 0<br />
| 5437.085<br />
| 0<br />
| –<br />
|-<br />
| D5<br />
| 82.7<br />
| 0.498<br />
| 1<br />
| 0<br />
| 5403.671<br />
| 97348.64<br />
| 319693.064<br />
| -309.852<br />
| 97348.64<br />
| 0<br />
| 5403.671<br />
| 0<br />
| –<br />
|-<br />
| D6<br />
| 79.7<br />
| 0.441<br />
| 1<br />
| 0<br />
| 5335.702<br />
| 96124.173<br />
| 353955.3<br />
| -306.082<br />
| 96124.173<br />
| 0<br />
| 5335.702<br />
| 0<br />
| –<br />
|-<br />
| D7<br />
| 76.6<br />
| 0.389<br />
| 1<br />
| 0<br />
| 5303.771<br />
| 95548.926<br />
| 395584.927<br />
| -304.379<br />
| 95548.926<br />
| 0<br />
| 5303.771<br />
| 0<br />
| –<br />
|-<br />
| D8<br />
| 73.6<br />
| 0.342<br />
| 1<br />
| 0<br />
| 5237.874<br />
| 94361.766<br />
| 440141.36<br />
| -300.723<br />
| 94361.766<br />
| 0<br />
| 5237.874<br />
| 0<br />
| –<br />
|-<br />
| FEED<br />
| 11<br />
| 1.013<br />
| 0<br />
| 0<br />
| 56107.895<br />
| 1032950<br />
| 1083.597<br />
| -3857.463<br />
| 1000930<br />
| 32021.357<br />
| 55559.983<br />
| 547.912<br />
| –<br />
|-<br />
| HOTFEED<br />
| 57.7<br />
| 1.013<br />
| 0<br />
| 0<br />
| 56107.895<br />
| 1032950<br />
| 1100.165<br />
| -3810.03<br />
| 1000930<br />
| 32021.357<br />
| 55559.983<br />
| 547.912<br />
| –<br />
|-<br />
| INPUT<br />
| 98<br />
| 1.013<br />
| 0<br />
| 0<br />
| 951221.824<br />
| 19301600<br />
| 25084.412<br />
| -64714.056<br />
| 16171700<br />
| 3129820<br />
| 897667.993<br />
| 53553.831<br />
| –<br />
|-<br />
| M1<br />
| 70.7<br />
| 0.304<br />
| 0.004<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 402879.11<br />
| -65216.53<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| M2<br />
| 70.7<br />
| 0.304<br />
| 0.01<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 897392.279<br />
| -65163.387<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| M3<br />
| 70.7<br />
| 0.304<br />
| 0.015<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 1393370<br />
| -65110.087<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| M4<br />
| 70.8<br />
| 0.304<br />
| 0.021<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 1894150<br />
| -65056.274<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| M5<br />
| 70.8<br />
| 0.304<br />
| 0.027<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 2396470<br />
| -65002.297<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| M6<br />
| 70.8<br />
| 0.304<br />
| 0.032<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 2903710<br />
| -64947.793<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| M7<br />
| 70.8<br />
| 0.304<br />
| 0.038<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 3412590<br />
| -64893.114<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| MIXED<br />
| 70.1<br />
| 0.304<br />
| 0<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 24740.57<br />
| -65269.173<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| P1<br />
| 93.4<br />
| 0.796<br />
| 0<br />
| 0<br />
| 3306.022<br />
| 59558.904<br />
| 61.859<br />
| -221.768<br />
| 59558.904<br />
| 0<br />
| 3306.022<br />
| 0<br />
| –<br />
|-<br />
| P2<br />
| 90.3<br />
| 0.71<br />
| 0<br />
| 0<br />
| 5542.299<br />
| 99846.076<br />
| 103.478<br />
| -372.082<br />
| 99846.076<br />
| 0<br />
| 5542.299<br />
| 0<br />
| –<br />
|-<br />
| P3<br />
| 87.3<br />
| 0.632<br />
| 0<br />
| 0<br />
| 5507.273<br />
| 99215.065<br />
| 102.608<br />
| -370.033<br />
| 99215.065<br />
| 0<br />
| 5507.273<br />
| 0<br />
| –<br />
|-<br />
| P4<br />
| 84.3<br />
| 0.562<br />
| 0<br />
| 0<br />
| 5437.085<br />
| 97950.605<br />
| 101.092<br />
| -365.615<br />
| 97950.605<br />
| 0<br />
| 5437.085<br />
| 0<br />
| –<br />
|-<br />
| P5<br />
| 81.2<br />
| 0.498<br />
| 0<br />
| 0<br />
| 5403.671<br />
| 97348.64<br />
| 100.269<br />
| -363.664<br />
| 97348.64<br />
| 0<br />
| 5403.671<br />
| 0<br />
| –<br />
|-<br />
| P6<br />
| 78.2<br />
| 0.441<br />
| 0<br />
| 0<br />
| 5335.702<br />
| 96124.173<br />
| 98.815<br />
| -359.382<br />
| 96124.173<br />
| 0<br />
| 5335.702<br />
| 0<br />
| –<br />
|-<br />
| P7<br />
| 75.2<br />
| 0.389<br />
| 0<br />
| 0<br />
| 5303.771<br />
| 95548.926<br />
| 98.036<br />
| -357.522<br />
| 95548.926<br />
| 0<br />
| 5303.771<br />
| 0<br />
| –<br />
|-<br />
| P8<br />
| 72.1<br />
| 0.342<br />
| 0<br />
| 0<br />
| 5237.874<br />
| 94361.766<br />
| 96.639<br />
| -353.366<br />
| 94361.766<br />
| 0<br />
| 5237.874<br />
| 0<br />
| –<br />
|-<br />
| P9<br />
| 69.4<br />
| 0.304<br />
| 0<br />
| 0<br />
| 4706.586<br />
| 84790.457<br />
| 86.695<br />
| -317.757<br />
| 84790.457<br />
| 0<br />
| 4706.586<br />
| 0<br />
| –<br />
|-<br />
| PREHEAT<br />
| 70.8<br />
| 0.304<br />
| 0.041<br />
| 0<br />
| 952338.165<br />
| 19322900<br />
| 3716880<br />
| -64860.601<br />
| 16191300<br />
| 3131550<br />
| 898754.695<br />
| 53583.47<br />
| –<br />
|-<br />
| PURE<br />
| 72.1<br />
| 0.342<br />
| 0.019<br />
| 0<br />
| 35737.995<br />
| 643829.982<br />
| 58727.526<br />
| -2404.05<br />
| 643829.982<br />
| 0<br />
| 35737.995<br />
| 0<br />
| –<br />
|-<br />
| TANK<br />
| 70.1<br />
| 0.304<br />
| 0<br />
| 0<br />
| 961549.437<br />
| 19509800<br />
| 24979.868<br />
| -65900.475<br />
| 16347900<br />
| 3161840<br />
| 907447.693<br />
| 54101.743<br />
| –<br />
|-<br />
| WASTE<br />
| 70.1<br />
| 0.304<br />
| 0<br />
| 0<br />
| 9211.272<br />
| 186896<br />
| 239.297<br />
| -631.301<br />
| 156606.799<br />
| 30289.201<br />
| 8692.998<br />
| 518.274<br />
| –<br />
|}<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5122Desalination - Team G2016-03-11T18:04:50Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> ASPEN+ simulation <div>]]<br />
<div align=left><br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A: Block Flow Diagram==<br />
[[File:Watson appendixA.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix B: Process flow diagram==<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix C: Aspen Model===<br />
[[File:Watson_aspen.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix D: Equipment and Utility Costs===<br />
[[File:Watson_utility_appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
[[File:Watson equip appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix E: Seawater Composition===<br />
[[File:Watson_seawater_appE.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix F: Calculations for Projected Production===<br />
[[File:Watson_appF.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix G: Stream Table===<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Stream Table based on ASPEN model'''<br />
! Date:<br />
!Temperature (C) <br />
!Pressure (bar) <br />
!Vapor Frac <br />
!Solid Frac <br />
!Mole Flow (kmol/hr) <br />
!Mass Flow (kg/hr) <br />
!Volume Flow (cum/hr) <br />
!Enthalpy (Gcal/hr) <br />
!Mass Flow H2O (kg/hr) <br />
!Mass Flow NACL (kg/hr) <br />
!Mole Flow H2O (kmol/hr) <br />
!Mole Flow NACL (kmol/hr)<br />
|-<br />
| BR1<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| BR2<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| BR3<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| BR4<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| BR5<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| BR6<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| BR7<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| BR8<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| BRINE<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| D1<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| D2<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| D3<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| D4 <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
| –<br />
|-<br />
| D5<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| D6<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| D7<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| D8<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| FEED<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| HOTFEED<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| INPUT<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| M1<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| M2<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| M3<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| M4<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| M5<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| M6<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| M7<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| MIXED<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P1<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P2<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P3<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P4<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P5<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P6<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P7<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P8<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| P9<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| PREHEAT<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| PURE<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| TANK<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| WASTE<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|}<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5100Desalination - Team G2016-03-11T17:52:56Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> ASPEN+ simulation <div>]]<br />
<div align=left><br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A: Block Flow Diagram==<br />
[[File:Watson appendixA.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix B: Process flow diagram==<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix C: Aspen Model===<br />
[[File:Watson_aspen.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix D: Equipment and Utility Costs===<br />
[[File:Watson_utility_appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
[[File:Watson equip appD.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix E: Seawater Composition===<br />
[[File:Watson_seawater_appE.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix F: Calculations for Projected Production===<br />
[[File:Watson_appF.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
===Appendix G: Stream Table===<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=File:Watson_appF.JPG&diff=5099File:Watson appF.JPG2016-03-11T17:52:50Z<p>WFu: </p>
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<div></div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=File:Watson_seawater_appE.JPG&diff=5098File:Watson seawater appE.JPG2016-03-11T17:51:50Z<p>WFu: </p>
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<div></div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=File:Watson_equip_appD.JPG&diff=5097File:Watson equip appD.JPG2016-03-11T17:50:50Z<p>WFu: </p>
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<div></div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5095Desalination - Team G2016-03-11T17:45:55Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> ASPEN+ simulation <div>]]<br />
<div align=left><br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A: Block Flow Diagram==<br />
[[File:Watson appendixA.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
<br />
==Appendix B: Process flow diagram==<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5094Desalination - Team G2016-03-11T17:44:39Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> ASPEN+ simulation <div>]]<br />
<div align=left><br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A: Block Flow Diagram==<br />
[[File:Watson appendixA.JPG|frame|center|border|<div align=center> <div>]]<br />
<div align=left><br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=File:Watson_appendixA.JPG&diff=5093File:Watson appendixA.JPG2016-03-11T17:43:54Z<p>WFu: </p>
<hr />
<div></div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5092Desalination - Team G2016-03-11T17:43:04Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
[[File:Watson_PFD_appendixB.JPG|frame|center|border|<div align=center> ASPEN+ simulation <div>]]<br />
<div align=left><br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=File:Watson_PFD_appendixB.JPG&diff=5091File:Watson PFD appendixB.JPG2016-03-11T17:42:14Z<p>WFu: </p>
<hr />
<div></div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5090Desalination - Team G2016-03-11T17:40:24Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
[[File:Watson f3.JPG|frame|center|border|<div align=center> Figure 3. Sensitivity analysis based on change in the 20 year projection. <div>]]<br />
<div align=left><br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5089Desalination - Team G2016-03-11T17:38:14Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
[[File:Watson f2.JPG|frame|center|border|<div align=center> Figure 2. a) Final equilibrium temperature and pressure as a function of the number of flash drums. b) Cost of increasing number of drums. <div>]]<br />
<div align=left><br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
FIGURE 3<br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5088Desalination - Team G2016-03-11T17:37:29Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.JPG|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left>FIGURE 1<br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
FIGURE 2<br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
FIGURE 3<br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5087Desalination - Team G2016-03-11T17:36:34Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:Watson f1.jpg|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left>FIGURE 1<br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
FIGURE 2<br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
FIGURE 3<br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
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"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
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4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
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"Government study of Oregon’s energy and water needs."<br />
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7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
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8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
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"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
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10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
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"Public mandates regarding maximum contaminant levels for potable water."<br />
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11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
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"Cost-benefit analysis of desalination and other forms of water recycling."<br />
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12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
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"Summary of the current state of forward osmosis technology."<br />
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13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
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"Summary of the current state of reverse osmosis technology."<br />
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14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
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"Oregon drinking water maximum contamination limits."<br />
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15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
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"Fouling of membranes in reverse osmosis."<br />
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16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
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"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
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18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
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19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5086Desalination - Team G2016-03-11T17:35:10Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:watson_f1.jpg|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left>FIGURE 1<br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
FIGURE 2<br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
FIGURE 3<br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5085Desalination - Team G2016-03-11T17:34:38Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:watson f1.jpg|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left>FIGURE 1<br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
FIGURE 2<br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
FIGURE 3<br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Desalination_-_Team_G&diff=5084Desalination - Team G2016-03-11T17:33:39Z<p>WFu: </p>
<hr />
<div>Team G: Newport Desalination Plant<br />
<br />
Authors: KC Anderson, Neil Dalvie, Watson Fu, Helen Wu<br />
<br />
Instructors: Fengqi You, David Wegerer<br />
<br />
March 11, 2016<br />
<br />
=Executive Summary=<br />
<br />
This reports outlines the design and evaluation of a multi-stage flash distillation plant located in Newport, Oregon. As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Oregon state recently announced the end of drought conditions from the last few years. This plant is designed as preemptive action to reduce the effects of future droughts by meeting ~10% of the expected increase in water demand in the mid-coastal region of Oregon. However, prohibitive costs in the current design may inhibit preemptive investment. <br />
<br />
The choice of an MSF process over membrane technologies was twofold. First, Oregon has strict water purity limits that can be more easily met with a robust thermal separation. Second, Newport is a hub of renewable energy research, and we believe that in the future this process can be paired with cheap, renewable thermal energy. The process consists of a major 9 stage flash vacuum unit, and a large surge tank. Feed and recycle are mixed in the surge tank for heat capture, and flows are optimized for a 62% seawater yield. The flash unit consists of 9 conjoined drums with condensing equipment and collection trays in the upper portions. Seawater feed is used as the cooling medium for condensation. <br />
<br />
Optimization for high yield of distillate has the unwanted effect of creating large flows through the flash unit recycle loop. These flows mandate nearly unrealistic heat transfer requirements, resulting in large equipment and prohibitively high utility requirements. In future iterations of this design, we recommend parallel processes with a reduced distillate yield. Flash situations with lower flow rates and energies are essential for an affordable process. <br />
<br />
An economic analysis revealed a loss of several billion dollars in a 20 year prediction. This was largely due to a high utility requirement. The predicted net present value at 20 years is most sensitive to heat and power requirements. With a refined design and committed investment and government subsidies, this process may be feasible, while expensive. However, given the current water climate and price, an investment of this magnitude may be difficult to accomplish in preemptive non-drought conditions. Therefore, we recommend a refined process design, and reevaluation of the water market and availability in Oregon in 3-5 years.<br />
<br />
__TOC__<br />
<br />
=Introduction=<br />
<br />
As the human population increases, so does the need for new sources of freshwater. Global water use is currently at 9 trillion cubic meters per year and is expected to increase by about 60 billion cubic meters per year. In addition, droughts in the west coast of the United States have made the need for new sources of freshwater much more urgent in recent years. Growing urban populations in developed countries also have high requirements for water, and 39% of global population lives within 100 kilometers of an ocean coast.<sup>1</sup> This means that a large percentage of people do not have access to fresh water sources. Because of these concerns, a large market exists for desalinated water. As of 2013, desalination plants produced 78.4 million cubic meters of water per day and this number is expected to increase.<sup>2</sup><br />
<br />
There are two main categories of methods of desalination used in industry. The first category is thermal-based separation. Multi-Stage Flash Distillation (MSF) has been widely utilized and involves heating and pressurizing impure water to separate water vapor. MSF is the most popular thermal separation method because of the high purity that can be obtained.<sup>3</sup> The second category is membrane-based separation. Reverse Osmosis (RO) is becoming the preferred method in industry. RO uses a pressure gradient to drive water through a membrane. Compared to most other methods, RO has low energy requirements and higher yield.<sup>4</sup><br />
<br />
The purpose of this report is to examine the potential implementation of a MSF desalination plant and evaluate the economic feasibility of the design. The remainder of the report outlines the process design, economics of the design, and important recommendations to further optimize the design and increase economic feasibility.<br />
<br />
==Design Basis==<br />
<br />
===Location===<br />
<br />
This desalination plant will be located in Newport, Oregon to provide water to the mid-coastal region of Oregon. Oregon has recently suffered a major drought, and 23 out of 36 counties implemented agricultural water regulation and applied for federal assistance.<sup>5</sup> As 2016 arrives, Oregon has ended its state drought emergency, but many new water regulations and conservation efforts appear to be permanent going forward. Oregon also offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup> The town of Newport boasts proximity to free coastline, and separation from major wildlife and forest reserves. In 2013, Oregon State University selected Newport as the location for its Pacific Marine Energy Center, a large scale trial of renewable wave energy.<sup>7</sup> Independent of the efficacy of wave energy, this project indicates the overall availability and commitment to renewable energy in Newport. Finally, while it is slightly farther from the drought stricken southern Oregon counties that are most affected by the California shortage, Newport lies in proximity to the agriculture-rich and highly water-dependent agriculture in the Willamette valley.<sup>8</sup><br />
<br />
<br />
===Process Requirements===<br />
<br />
This plant will produce 15,500 cubic meters of desalinated water per day, aiming to offset on the order of ~10% of predicted increase in out-of-stream water demand in the mid-coastal region in coming years. The feed for this process is only seawater, sourced from the Newport coastal water with a salinity of 32 PSU (g/kg of seawater).<sup>9</sup> The process will produce 99.2% desalinated, potable water with a maximum chloride concentration of 250 mg/L and total dissolved solids of 500 mg/L. As waste, the process will release diluted, cooled, brine from a surge tank. Composition of the feed seawater can be found in Appendix E. <br />
<br />
=Technical Approach=<br />
<br />
We decided to use Multi-Stage Flash Distillation (MSF) for the desalination process in our plant. Principally, MSF allows us to achieve the purity required for Oregon regulations. Oregon water regulations include an upper limit of salt concentration at 250 mg/L for potable water.<sup>10</sup> Because of this, MSF provides a more reliable high purity product than does reverse osmosis, the main alternative. In addition, thermal methods like MSF achieve the desired purity with less dependence on input conditions. While we expect seawater concentrations to remain largely constant, a robust process is desirable. While membranes require significant pretreatment of feeds, thermal methods can process raw seawater and do not run the risk of microbial contamination.<sup>4</sup> Despite this advantage, MSF typically sees considerably lower yield, and higher thermal energy costs than reverse osmosis.<sup>11</sup> This decision was made after considering a number of options, described in this section. Design alternatives are based on a simple separation block diagram, shown in Appendix A.<br />
<br />
==Process Alternatives==<br />
<br />
===Pressure Control Design Options===<br />
<br />
An important aspect of most desalination processes is establishing a pressure gradient. In membrane technologies, the pressure gradient is a driving force for separation against a concentration gradient. Forward osmosis holds a major advantage in this section of the process, as little to no gauge pressure is required to drive osmosis.<sup>12</sup> In comparison, reverse osmosis requires high levels of pressure to achieve separation.<sup>13</sup> The magnitude of the pressures increases capital costs and utilities costs tremendously, which is a significant disadvantage. For thermal separation technology, low pressure works in accordance with the thermal changes to remove steam from the concentrated brine, as the water vapor saturation temperature changes with changing pressure. The two main methods of vacuum creation are seawater eductors and vacuum pumps. An eductor is convenient when high energy flows are accessible within the process. In the absence of extra flows, we decided to utilize a simple vacuum pump. While energy intensive, this pump achieves low pressures easily. <br />
<br />
===Pretreatment Design Options===<br />
<br />
Membrane technologies, including forward and reverse osmosis, are limited by the size and selectivity of the membrane. This presents an issue, as Oregon mandates strict upper limits on organic contaminants.<sup>14</sup> One solution to this issue is to source water from either several hundred meter depth or from beach wells, where water has already passed through sediment.<sup>15</sup> In addition to feed requirements, reverse osmosis methods require several pretreatment steps to avoid severe membrane fouling.<sup>13</sup> Forward osmosis processes require the addition of a draw solution on the permeate side of the membrane to create an osmotic pressure driving force.<sup>12</sup> Thermal desalination relies on the heating of seawater to obtain a pure distillate. In early design stages, we considered the implementation of a refrigeration loop. Unfortunately, the purchase of refrigerants are prohibitively expensive,<sup>16</sup> and a refrigeration loop is beneficial when heat needs to be transferred from one area of the process to another. With the implementation of a vacuum pump, there is nothing in the process that needs to be cooled. For this reason and cost, we decided to heat our process stream using a condensing steam heat exchanger.<br />
<br />
===Separation Design Options===<br />
<br />
One of the main separation methods for desalination is membrane separation. Forward osmosis relies on a membrane to allow transfer of water under purely osmotic forces. However, continuous flow is difficult to arrange spatially since the concentrated draw solution must be recycled back through the system. Very little literature exists on practical uses of forward osmosis membranes for desalination, so we have chosen to avoid this option. Reverse osmosis uses hydraulic pressure to force osmosis, rather than a draw solution and concentration gradient. Reverse osmosis can generally achieve only 98% salt removal, requiring multiple passes.<sup>12</sup> A vast majority of MSF processes are centered around a series of flash chambers with descending pressure and temperature. Vaporized water is collected in a tray as the pure distillate, with increasingly concentrated brine flowing into the next flash chamber. In order to maintain the pressure gradient needed, a vacuum pump is used. By aligning the flash chambers into one unit, only one pump would be needed to create the pressure gradient, reducing both capital and operating costs.<sup>17</sup> Therefore, we have decided to move forward with MSF with the use of flash chambers connected into one unit for our separation.<br />
<br />
===Waste Treatment Design Options===<br />
<br />
Reverse osmosis typically requires additional steps to return the water product to an acceptable pH after the initial acidification before release, in addition to dilution.<sup>13</sup> Forward osmosis technology requires separation of pure water from the draw solution through heating. This adds significantly to the otherwise minimal energy requirement of a forward osmosis process.<sup>12</sup> One technology that could improve waste treatment for an MSF process is adding a brine recycle. Two methods of concentrated brine recycle are prevalent. In one method, a portion of concentrated brine is recycled into the seawater feed, with the rest of the brine sent to dilution and waste.<sup>18</sup> Alternatively, concentrated bring can be rerouted to a surge tank. This tank is controlled to maintain a concentration acceptably diluted for waste, serving as the seawater feed and the waste “purge”, with the two having the same composition.<sup>17</sup> Traditionally in chemical processes, recycle systems require more energy to carry out the process. Because desalination is itself a separation, recycle may be advantageous because of the retained heat energy. In the second recycle method, the surge tank serves not only to cool the diluted waste to an acceptable release temperature, but also to preheat the process feed. In this setup, where no heat is rejected into the waste, thermal efficiency may actually increase, decreasing utility costs.<sup>18</sup> For these reasons, we have decided to implement a surge tank recycle stream.<br />
<br />
=Results=<br />
<br />
==Design Tradeoffs and Process Optimization==<br />
<br />
Once the overall design equipment and strategy was selected, mass and energy balances were calculated and optimized for yield and cost. To determine these values, temperatures, and flow rates, we made a number of assumptions and set points in our process. The feed and waste concentrations were held constants, at the composition of Oregon sea water, and the maximum allowable waste concentration. The flash inlet was held at 1 atm and 98°C, in order to maximize energy carried by the stream without premature boiling. The distillate flow rate was held constant in line with our initial problem statement and project goals. Finally, phase data was obtained from Aspen+. While true seawater will contain other contaminants, these have small effects on thermodynamic properties. Pretreatment and material selection will take additional contaminants into consideration, but they are neglected in mass and energy calculations.<br />
<br />
===Pressure Considerations and Yield===<br />
<br />
Aspen+ phase data revealed that because the energy used to vaporize the water is carried in the inlet stream, the amount of water flashed depends almost completely on the pressure in the last flash stage, or the lowest pressure in the process. Because of this, the mass balances over the entire process are largely dependent on the equilibrium conditions in the last drum. Therefore, for overall balances, we treated the connected series of flash drums as one unit. This assumption is based on the adiabatic nature of the drums, and the assumption that the brine reaches phase equilibrium before leaving the unit. This yields a simplified block diagram for the purpose of calculating overall mass balances, as shown in Appendix A. Figure 1a shows conditions at a range of vacuum pressures. As pressure is decreased, the yield of vaporization increases, which corresponds to an increase in the outlet concentration of NaCl for recycle. Temperature decreases with pressure to maintain vapor-liquid saturation conditions. The temperature profile is critical in designing the multistage flash unit, as higher temperatures through the pressure gradient will release hot distillate that can be captured in preheating.<br />
<br />
[[File:watson_f1.jpg|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left>FIGURE 1<br />
<br />
At first inspection, it appears advantageous to operate at the lowest possible pressure to obtain the highest vaporization yield. However, dilution for waste proved to be a more significant factor in overall process yield than the yield over the flash drum unit. Operating at the lowest possible pressure maximizes vapor yield, but creates a more concentrated recycle stream. This higher concentration requires more process feed to dilute to waste conditions, lowering the overall process yield. For this reason, it is desirable to produce a recycle stream as close to waste concentration as possible, minimizing the amount of process feed needed to dilute to waste conditions. Figure 1b shows the effect of flash pressure on overall process metrics. It becomes clear that the overall yield increases with pressure as an asymptote. Above a certain pressure, the recycle stream becomes too dilute to create a waste concentration of 40 g/kg, creating a negative feed requirement for this calculation. Because we would like to release waste of 40 g/kg, we focus on the feasible solutions below 0.4 atm. Figure 1b also shows the small effect on heating requirements as the pressure is changed. Because the amount of water vaporized is held constant, this energy is largely representative of the energy needed to vaporize that amount of water. Figure 1a shows that at higher operating pressures and lower vaporization yields, the brine recycle will remain hot. Therefore, despite increased recycle rates, the higher temperature keeps the energy requirement nearly constant. With these considerations, we will operate at a flash pressure that limits the vaporization yield, keeping the recycle stream near waste concentrations. When operating at a pressure of 0.3 atm, an overall yield of approximately 62% can be achieved. This higher pressure will also provide energy savings in vacuum creation.<br />
<br />
===Flash Stage Optimization and Sizing===<br />
<br />
The flash unit, where all flash stages occur, and makes up the bulk of the process. It consists of 9 vertical flash drums connected in series, each with a condenser in the upper portion. The drums are held at low pressure, allowing the volume to fill with saturated water vapor. This vapor condenses on heat exchange pipes in the top of the drum, and condenses, falling onto a collection tray. Once overall mass balances were calculated, detailed mass and energy balances on the major flash unit were analyzed. First, the equilibrium in each stage was characterized. Connected equilibrium stages exhibit linearly decreasing temperature.<sup>19</sup> Optimization of mass balances called for a pressure of 0.3 atm in the last drum to achieve the highest yield. This produces the following equilibrium conditions across all 9 drums (the number of drums eventually selected).<br />
<br />
FIGURE 2<br />
<br />
Interestingly, this equilibrium profile creates nearly constant vapor flow rates in each drum. These flow rates, along with the density of saturated vapor in each drum, were used to calculate the minimum chamber diameter to allow enough vapor-liquid interface. The minimum diameter for the lowest pressure drum was approximately 6 m, and the minimum for preceding drums varied minimally. For manufacturing simplicity, we have decided to build all 9 drums using a diameter and height of 6 m. As explained previously, drums will be constructed of stainless steel to prevent corrosion. Because each drum will be operated in vacuum conditions, the pressure on drum walls will never exceed 1 atm. We therefore calculated the thickness of material based on structural feasibility, requiring a thickness of 3 cm. Pricing of the drums was approximated using the required weight of stainless steel. In addition to drum material, the cost of each condensing unit was approximated using the required heat exchange area to condense the distillate. This calculation was performed under the assumption that the surge tank outlet (See Appendix L) will be used as the cooling stream on the tube side. Approximate cost of each condensing system was derived heuristically from the heat transfer area. All of these equilibrium and costing calculations were repeated for a growing number of stages. Figure 2b indicates the total cost of the flash unit for different numbers of stages. Using this, we selected 9 tanks, within the region of diminishing returns. Once this number was selected, we began specific design of each condensing region of the drums. Notably, the last three drums require excessive heat transfer area that exceeds 5000 square meters. The last drum, which requires nearly 19000 square meters, can be remedied by using the process seawater feed to cool, which is at a significantly lower temperature than the surge tank. Unfortunately, the process feed does not have the heat capacity to cool the 7th and 8th drums. For this initial design and economic analysis, these drums retain an unrealistic heat transfer area. In implementation, additional cooling water would be required. The cost of demisters and trays were also considered. The approximate price for a 6-meter diameter demister is $500. Both the demister and condensing tray price were considered negligible compared to the overall price of the flash chambers.<br />
<br />
==Process Overview==<br />
<br />
The final design process flow diagram is shown in Appendix B. Feed seawater is pumped into the plant, and immediately used as a condensing sink in the last drum. The warmed seawater is then sent for mixing in the surge tank. The surge tank outlet is used as the condensing heat sink for the other 8 tanks in series, before being delivered to the heat exchanger. In the exchanger, the flash feed is heated to 98°C before entering the first drum. The brine then passes through all drums, reaching phase equilibrium in each one as the pressure is reduced. From the last drum, the concentrated brine is pumped out and back into the surge tank. The surge tank includes a waste purge back out to the ocean. Stream tables are included with compositions and conditions for each stream. Notably, the concentration in the surge tank is 40 g/kg, the maximum allowable waste concentration. This also serves as the flash feed. In addition, there is a large amount of fluid in recirculation through the recycle loop, with relatively small process feed and waste. This has many implication, positive in the large increase in overall yield, and negative in the energy costs and large equipment sizes required. Economic implications of these large flows are addressed in later sections. For reference, the Aspen+ file used for phase and mass balance calculations is shown in Appendix C.<br />
<br />
==Equipment Sizing==<br />
<br />
===Surge Tank===<br />
<br />
To account for corrosion, the surge mixing tank will be constructed of stainless steel. The size of the surge tank was based on two criteria. First, a residence time of 30 minutes was specified to ensure full mixing of process feed and brine recycle. Second, the tank is designed to hold a large percentage of the brine in recycle circulation, in case the process needs experiences a sudden shutdown. These criteria resulted in surge tank dimensions of a diameter and height of 24.3 m.<br />
<br />
===Heater===<br />
<br />
The heater E-101 is the primary energy input for the process. It uses condensing steam to heat the flash feed to 98 C. This heat exchange will be very large, measuring 25 ft. in length and 7 ft. in diameter. It requires high amount of high pressure steam, and results in a pressure drop of 25 bar, creating much of the pumping requirement for the process. Detailed design of the heater can be found in Appendix M. <br />
<br />
===Pumps===<br />
<br />
Based on sizing estimations given in Towler<sup>20</sup>, Ch. 7, P-101 will be $323,151, P-102 will be $12,446, and P-103 will be $311,369. The utilities can be calculated using the brake hp of the pumps. P-101 has a brake hp of 166,870 kW, P-102 has a brake hp of 4.1 kW, and P-103 has a brake hp of 457.37 kW. The utility cost for P-101 is large because of the pressure drop in the heater, E-101. Detailed design of pumps can be found in Appendices N-P.<br />
<br />
==Safety, Control, and Environmental Considerations==<br />
<br />
===Controls===<br />
<br />
This process exhibits four major control loops. First, two control loops exist within the surge tank, comprising most of the process control. Liquid level is controlled by manipulating the waste flow rate, and composition is manipulated by controlling the feed flow rate. Pressure in the flash unit is controlled by manipulating the vacuum pump power. Finally, the brine inlet temperature is controlled by altering the steam delivered to the heat exchanger. <br />
<br />
===Environmental Considerations===<br />
<br />
Our process does not include a significant post treatment, and produces de-ionized water because our plant must meet demands for both agriculture and human consumption. Because Oregon has recently left drought conditions, we believe that local treatment centers currently have additional capacity available. Oregon has strict regulations on the salinity of wastewater for marine health. Our process reaches but does not exceed the maximum salt concentration of 40g/L in wastewater. We plan to utilize clean energy for our process. Oregon offers an abundance of and commitment to clean energy not seen in other states. In California, desalination in reaction to extreme drought has forced plants to rely on destructive fossil fuel energy sources, thereby offsetting the environmental effect of desalination. A desalination process in Oregon would be posed to take advantage of a high and growing grid of renewable power. So far, 73% of Oregon’s electricity generation is renewable.<sup>6</sup><br />
<br />
===Scaling and Corrosion===<br />
<br />
In the brine recycle stream, salt levels climb as high as 41 g/kg. While not extreme conditions, this salinity along with additional contaminants brings attention to the prevention of corrosion and control of scaling. We have decided to build our desalination plant with 316 stainless steel because of its excellent resistance to general and localized corrosion.<sup>21</sup> Although 316 stainless steel is three times as expensive as carbon steel it will maintain efficient operation with less failures due to corrosion damage.20 We have decided to add an antiscalant to our process instead of acid addition.<sup>22</sup> Acidification, although effective in preventing the precipitation of calcium carbonate, is relatively ineffective in preventing other types of scale and also less cost effective.<sup>23</sup> Based on performance studies, we have decided to use a polyphosphate acid inhibitor as our antiscalant, at a dosing rate of 1.5ppm due to the relatively low maximum operating temperature of 98°C.22<br />
<br />
==Economic Evaluation and Sensitivity Analysis==<br />
<br />
The ISBL capital costs were estimated to be 14.5 MM$, while OSBL costs were estimated to be 40% of ISBL costs. Individual equipment costs can be found in Appendix D. Since the plant is located on the West Coast, a location factor of 1.07 was applied. The variable cost of production for the plant has three main sources: raw materials, consumables, and utilities. The main raw materials costs are from the antiscalant, since we will not have to pay for the seawater feed; the antiscalant will cost $542,000/year for the flow of 100,000 m<sup>3</sup> of seawater per day. Utility costs mainly consisted of electricity for the pumps and steam for the heat exchanger. Detailed equipment and utility costs are shown in Appendix D. The total fixed capital cost was calculated to be 129.7 MM$. Major assumptions include having three shifts of five operators earning $50,000 salaries, maintenance of the plant at 5% of ISBL costs, and direct labor overhead being 25% of operator costs.<br />
<br />
Profitability of the plant was assessed by calculating the net present value (NPV) and internal rate of return (IRR). The price of purified water was assumed to be $3.00 per 1000 gallons from average water costs in Newport, Oregon and assuming we will have to sell our water at a lower price to treatment facilities before being sold for public use.<sup>24</sup> Assuming a 5-year MACRS depreciation schedule, this plant is not profitable by a large margin, losing up to 7 billion dollar over a 20 year window. Main sources of the imbalance are high utility costs ($150 MM for high pressure steam). If profit is the goal of this process, we would not recommend pursuing this project. However, if there a pressing need and investment for purified water, and a potential source of local energy from renewable sources, this process may be a feasible project, though expensive. The full economic analysis can be found in Appendix J. A sensitivity analysis revealed that lowering high energy requirements is imperative to drive down prohibitive costs. In addition, subsidies would be necessary to execute this process. Sensitivity to several process parameters is represented in Figure 3.<br />
<br />
FIGURE 3<br />
<br />
=Conclusion=<br />
<br />
In this report we outline a proposed multi-stage flash distillation process designed to meet ~10% of the expected increase in mid-coastal Oregon water demand. The plant uses a 9 flash drum vacuum unit for distillation, and a large surge tank for heat capture and process control. Most major concerns with the current proposed plant involve the massive flow rates through the recycle loop. This creates large energy requirements and unrealistic heat exchange at several locations. In a refined design, we recommend sacrificing process yield to reduce the recycle ratio. A lower flash yield with several identical processes in parallel allows for reasonable heat exchange and equipment design. The abundance of cool seawater should be used more fully in the process. <br />
<br />
An economic evaluation revealed significant losses in a 20 year prediction. With improved process design, this plant could be feasible with committed investment and government subsidies. However, this plant is designed to meet future needs, which are not pressing right now. Due to the immense energy intensive cost of this project, it may be difficult to secure support for a preemptive desalination strategy, and merits reevaluation in 3-5 years. <br />
<br />
=Appendices=<br />
==Appendix A==<br />
<br />
=References=<br />
1. Global Agenda Council on Water. World Economic Forum website. http://www.weforum.org/communities/global-agenda-council-on-water. Accessed January 14, 2016.<br />
<br />
"Calculated demand for water and energy on basis of population growth."<br />
<br />
2. Desalination industry enjoys growth spurt as scarcity starts to bite. Global Water Intelligence website. https://www.globalwaterintel.com/desalination-industry-enjoys-growth-spurt-scarcity-starts-bite/. Accessed January 14, 2015.<br />
<br />
"Study about desalination plants and their effectiveness in different regions."<br />
<br />
3. Sieder, Everett N, inventor; Us Interior, assignee. Multistage flash distillation with scale removal. US patent 3,476,654. November 4, 1969.<br />
<br />
"Patent on multistage flash distillation."<br />
<br />
4. Lee, KP, Arnot, TC, Mattia, D. A review of reverse osmosis membrane materials for desalination - Development to date and future potential. Journal of Membrane Science. 2011, 370: 1-22.<br />
<br />
"Article discussing efficacy of materials in reverse osmosis membranes."<br />
<br />
5. House, Kelley. Oregon drought forces cities to impose water use cutbacks. The Oregonian. http://www.oregonlive.com/environment/index.ssf/2015/08/oregon_drought_forces_cities_t.htm. Published August 1, 2015. Accessed January 13, 2016.<br />
<br />
"News article about severity of Oregon droughts."<br />
<br />
6. Oregon State Profile and Energy Estimates. U.S. Energy Information Administration. http://www.eia.gov/state/?sid=OR. Updated October 15, 2015. Accessed January 13, 2016. <br />
<br />
"Government study of Oregon’s energy and water needs."<br />
<br />
7. Batten, Belinda. Newport selected as home of Pacific Marine Energy Center. Oregon State University. http://oregonstate.edu/ua/ncs/archives/2013/jan/newport-selected-home-pacific-marine-energy-center. Published January 14, 2013. Accessed January 13, 2016.<br />
<br />
"Wave energy discussion by Oregon State professors."<br />
<br />
8. Oregon Agricultural Regions. State of Oregon Department of Agriculture. http://www.oregon.gov/ODA/shared/Documents/Publications/Administration/ORGrowingRegions.pdf. Accessed January 13, 2016.<br />
<br />
"Agricultural regions of Oregon."<br />
<br />
9. Salinity Distribution at the Ocean Surface. Centre Aval de Traitment des Données SMOS. http://www.salinityremotesensing.ifremer.fr/sea-surface-salinity/salinity-distribution-at-the-ocean-surface. Accessed January 14, 2016.<br />
<br />
"Tabulated data about salinity of ocean water in different regions."<br />
<br />
10. Maximum Contaminant Levels and Action Levels. Oregon Public Health Division. https://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Published May 8, 2014. Accessed January 14, 2016.<br />
<br />
"Public mandates regarding maximum contaminant levels for potable water."<br />
<br />
11. Desalination and Water Recycling. Terrascope. http://12.000.scripts.mit.edu/mission2017/desalination-and-water-recycling/. Accessed January 13, 2016.<br />
<br />
"Cost-benefit analysis of desalination and other forms of water recycling."<br />
<br />
12. Cath, T.Y., Childress, A.E., Elimelech, M. Forward osmosis: Principles, applications, and recent developments. Journal of Membrane Science. 2006. 281: 70-87. <br />
<br />
"Summary of the current state of forward osmosis technology."<br />
<br />
13. Greenlee, L.F., Lawler, D.F., Freeman, B.D., Moulin, P. Reverse osmosis desalination: Water sources, technology, and today’s challenges. Water Research. 2009. 43: 2317-2348.<br />
<br />
"Summary of the current state of reverse osmosis technology."<br />
<br />
14. Maximum Contaminant Levels and Action Levels. Oregon Health Authority website. http://public.health.oregon.gov/HealthyEnvironments/DrinkingWater/Rules/Documents/61-0030.pdf. Accessed January 28, 2016. <br />
<br />
"Oregon drinking water maximum contamination limits."<br />
<br />
15. Winters, H. Twenty years experience in sea water reverse osmosis and how chemicals in pretreatment affect fouling of membranes. Desalination. 1997. 110: 93-96. <br />
<br />
"Fouling of membranes in reverse osmosis."<br />
<br />
16. Refrigeration Cycles. Oklahoma University. http://www.ou.edu/class/che-design/che5480-07/Refrigeration%20Basics%20and%20LNG.pdf. Accessed January 27, 2016. <br />
<br />
Details on selection of appropriate refrigerant.<br />
<br />
17. Williamson, William R, inventor; American Mach & Foundry, assignee. Multistage flash distillation apparatus. U.S. patent 3,399,118. August 27, 1968. <br />
<br />
"Basis for our design. MSF with connected chambers, one eductor, and a complete brine dilution recycle."<br />
<br />
18. El-Dessouky, H.T., Ettouney, H.M., Al-Roumi, Y. Multi-stage flash desalination: present and future outlook. Chemical Engineering Journal. 1999, 73: 173-190.<br />
<br />
"Summary of MSF processes both traditional, and a new recycle method."<br />
<br />
19. Kaghazchi, Tahereh, et al. "A mathematical modeling of two industrial seawater desalination plants in the Persian Gulf region." Desalination 252.1 (2010): 135-142. Accessed February 29, 2016.<br />
<br />
"Temperature and pressure profiles through multiple connected flash chambers."<br />
<br />
20. Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. <br />
<br />
"Price comparison between carbon steel and stainless steel; equipment sizing"<br />
<br />
21. Malik, A. U., Al-Fozan, S. A. Corrosion and materials selection in MSF desalination plants. Corrosion Reviews. 2011: 29: 153-175.<br />
<br />
"Material performance for MSF in the presence of saltwater"<br />
<br />
22. Ghani, S., Al-Deffeeri, N. S. Impacts of different antiscalant dosing rates and their thermal performances in Multi Stage Flash (MSF) distiller in Kuwait. Desalination. 2010: 250: 463-472.<br />
<br />
"Background information on scaling"<br />
<br />
23. Scaling and Antiscalants. Lenntech Water Treatment Solutions website. http://www.lenntech.com/antiscalants.htm. Accessed February 28, 2016. <br />
<br />
"Scalant information and pricing"<br />
<br />
24. Utility Bill Calculator. City of Newport, Oregon website. http://www.thecityofnewport.net/dept/pwk/billcalc.asp. Accessed February 25, 2016. <br />
<br />
"Price of water in Newport, Oregon"<br />
<br />
25. Statewide Water Needs Assessment Oregon Water Supply and Conservation Initiative. Oregon Water Resources Department. http://www.oregon.gov/owrd/law/docs/owsci/owrd_demand_assessment_report_final_september_2008.pdf. Accessed January 14, 2016.<br />
<br />
"Optimization study about increasing water supply and demand chain problems."</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_production_cost_and_revenue&diff=5083Estimation of production cost and revenue2016-03-11T17:33:18Z<p>WFu: Undo revision 5082 by WFu (talk)</p>
<hr />
<div>Authors: Nick Pinkerton, Karen Schmidt, James Xamplas (ChE 352 in Winter 2014), Reed Kolbe (ChE 352, Winter 2016)<br />
<br />
Steward: David Chen, Fengqi You <br />
<br />
==Variable Cost of Production==<br />
Variable costs of production are dependent primarily on plant output and rate of production. There are many variables to consider when costing a plant. <br />
# Raw materials consumed<br />
# Utilities-steam, electricity, cooling water, fuel, etc.<br />
# Consumables - acids, bases, solvents, catalysts, etc.<br />
# Disposal<br />
# Shipping<br />
The majority of the variable costs for a production plant are the raw materials and utilities costs. Variable costs can be greatly cut through optimization techniques and intelligent plant design (Towler and Sinnott, 2013).<br />
===Raw Materials Cost===<br />
Calculating the annual cost of a raw material is calculated by simply multiplying the feed rate of the process by the appropriate price per volume or mass. These are the costs of chemical feed stocks required by the process. Feed stocks flow rates are obtained from PFD (Turton et al., 2013).There are several ways to optimize this cost to ensure that a process is not costing more than it should. First one should assess the actual consumption of a plant to see if it is significantly different from what should be expected based on process stoichiometry and selectivities (Towler and Sinnott, 2013). Finding may prove that a process is less efficient than it originally claimed.<br />
It is smart to benchmark a new plant design against an existing plant or pilot plant. Raw materials are typically the largest contributor to overall variable costs. For bulk chemicals and petrochemicals, raw materials represent 80-90% of the total cash cost of production (CCOP). <br />
===Utilities Cost===<br />
These are the costs of the various utilities streams required by the process. The flowrates for the utilities streams are located on the PFD (Turton et al., 2013). This includes:<br />
*Fuel gas, oil, or coal<br />
*Electric power<br />
*Steam<br />
*Cooling water<br />
*Process water<br />
*Boiler feed water<br />
*Air<br />
*Inert gas<br />
*Refrigeration<br />
<br />
Utility streams are excellent ways to streamline a process and are often indicative of how efficient of a process the project is. Process methods such as steam generation and pinch analysis can be used to greatly reduce utility costs across a plant. Further analysis of pinch analysis techniques and optimizing heat exchanger networks can be found in plant design texts such as first reference from Gavin Towler. The determination of process utility costs is often more difficult than the determination of raw material costs; however, the utilities are typically between 5-10% of CCOP (Towler and Sinnott, 2013). The cost of heating a process can be reduced by using process waste streams as fuel which consequently also reduces the need for waste disposal.<br />
<br />
===Waste Disposal Costs===<br />
These are defined as the cost of waste treatment to protect the environment (Turton et al., 2013). These are materials that cannot be recycled or sold off as by-products. Often times these streams require additives or additional treatment to meet governmental regulations. <br />
Hydrocarbon waste can often be incinerated directly to the atmosphere or used as process fuel to heat other streams in the system. Using the stream as process fuel allows the fuel value of the stream to be recovered into the system. The substituted value can be calculated by multiplying the conventional fuel price by the heat of combustion of the waste stream. <br />
<br />
<math>P_{WFV} = P_F * \Delta H_C^o</math><br />
<br />
where <br />
<math>P_{WFV}</math> = waste value of fuel ($/lb or $/kg)<br />
<br />
<math>P_F</math> = price of fuel ($/MMBtu or $/GJ)<br />
<br />
<math>\Delta H_C^o</math> = heat of combustion (MMBtu/lb or GJ/kg)<br />
<br />
Dilute aqueous streams must be sent to wastewater treatment typically prior to purging from the plant. Acidic or basic wastes are neutralized prior to treatment by salting out the acid or base. The cost of wastewater treatment is typically about $6 per 1000 gal but this is only an estimate that doesn't account for regional charges (Towler and Sinnott, 2013).<br />
<br />
Solid waste treatment can typically be sent to a landfill at a cost of approximately $50/ton (Towler and Sinnott, 2013). <br />
<br />
Hazardous wastes arise from the production of concentrated liquid streams that cannot be incinerated. Hazardous wastes should be avoided if possible, but that is not always feasible for some processes. The cost of hazardous waste disposal is strongly dependent on the location of the plant, the plants proximity to waste disposal plants and the degree of hazard of the waste.<br />
<br />
==Fixed Cost of Production==<br />
Fixed costs are those whose amounts are independent of production rates. Much of these costs are personnel salaries, taxes, insurance, and legal payments.<br />
===Labor Costs===<br />
These are the costs attributed to the personnel required to operate the process plant (Turton et al., 2013).<br />
<br />
The number of operators required per shift, <math>N_{OL}</math> can be estimated by<br />
<br />
<math>N_{OL}=(6.29+31.7P^2+0.23N_{np})^{0.5}</math><br />
<br />
where <math>P</math> is the number of processing steps involving particulate solids and <math>N_{np}</math> is the number of other processing steps (Turton et al., 2013). For each of the <math>N_{OL}</math> operators per 8-hour shift, approximately 4.5 operators must be hired for a plant that runs 24 hours per day, to account for the 3 shifts per day and the 3 weeks of leave typically taken by each operator per year (Turton et al., 2013). The salary for a chemical plant operator varies by location, and the estimator should look up the average value for the area.<br />
<br />
===Maintenance Costs===<br />
These are the costs associated with labor and materials necessary to maintain plant production. An estimate of these are 6% of the fixed capital investment (Turton et al., 2013).<br />
<br />
===Research and Development===<br />
These are the costs of research done in developing the process and/or products. This includes salaries for researchers as well as funds for research related equipment and supplies. An estimate of these costs are 5% of the total manufacturing cost (Turton et al., 2013).<br />
<br />
===Taxes and Insurance===<br />
Taxes vary by location, but a first estimate of property taxes and liability insurance is 3% of the fixed capital investment (Turton et al., 2013).<br />
<br />
===Plant Overhead===<br />
Overhead costs are the miscellaneous but necessary costs of running a business, including payroll, employee benefits, and janitorial services. This may be estimated as 70% of the operating labor costs, added to 4% of the fixed capital costs (Turton et al., 2013).<br />
<br />
===Licensing and Royalties===<br />
The costs of paying for the use of intellectual property clearly varies, but an estimate that may be used is 3% of the total manufacturing cost (Turton et al., 2013).<br />
<br />
==Revenues==<br />
The revenues of a process are the income earned form sales of the main products and the by-products. Revenue can be impacted by market fluctuations and production rates.<br />
===By-Product Revenues===<br />
Besides selling the main product from a process, by-products from separations and reactions can also be valuable in the market. Often it is more difficult to decide which by-products to recover and purify than it is to make decisions on the main product. <br />
<br />
By-products made in stoichiometric ratios from reactions must be either sold off or managed through waste disposal. Other by-products are sometimes produced through feed impurities or by nonselective reactions. There are several potential valuable by-products from a process: <br />
# Materials produced in stoichiometric quantities by the reactions that create the main product. If they are not recovered then the waste disposal expenses will be large.<br />
# Components that are produced in high yield by side reactions. <br />
# Components formed in high yield from feed impurities. Many sulfurs are produced as a by-product of fuels manufacture.<br />
# Components that are produced in low yield but have high value. An example includes acetophenone which is recovered as a by-product of phenol manufacture.<br />
# Degraded consumables (e.g. solvents, etc.) that have reuse value.<br />
<br />
A rule of thumb that can be used for preliminary screening of by-products for large plants is that for by-product recovery to be economically feasible the net benefit must be greater than $200,000 a year. A net benefit can be calculated by adding the possible resale value of the by-product and the avoided waste disposal cost (Towler and Sinnott, 2013). <br />
<br />
===Margin===<br />
The gross margin of a process is defined as the sum of product and by-product revenues minus the raw material cost. <br />
<br />
Gross margin = Revenues - Raw materials costs<br />
<br />
Because raw materials are most often the most expensive variable cost of a process, the gross margin is a good gauge as to what the total profitability of a process will be. Raw materials and product pricing are often subject to high degrees of variability which can be difficult to forecast. The size of margins are highly versatile depending on the <br />
industry. For many petrochemical industries the margin may be only 10%; however, for industries such as food additives and pharmaceuticals the margins are generally much higher (Towler and Sinnott, 2013).<br />
<br />
===Profits===<br />
There are several standards for calculating company profits. The cash cost of production (CCOP) is the sum of the fixed and variable production costs. <br />
<br />
<math>CCOP = VCOP + FCOP</math><br />
<br />
where <math>VCOP</math> is the variable cost of production and <math>FCOP</math> is the fixed cost of production. <br />
<br />
Gross profit, which should not be confused with gross margin, is then calculated by the following equation,<br />
<br />
<math>Gross\ profit = Main\ product\ revenues - CCOP</math><br />
<br />
Finally profit can be calculated by subtracting the income taxes that the plant would be subject to depending on the tax code of the county the plant is located in. <br />
<br />
<math>Net\ profit = gross\ profit - taxes</math><br />
<br />
==Pricing Products and Raw Materials==<br />
The revenues and costs of a project are vital to determining its economic feasibility. To calculate these values one needs to multiply the respective product and feed streams by their respective prices. The major difficulty of this process is determining the prices that should be used in this formula. When analyzing a plant, not only do the current prices need to be acknowledged but also the stability of the market to forecast future fluctuations and deviations. <br />
===Pricing Fundamentals===<br />
The pricing of a substance is determined by the fundamental economic principles of supply and demand. A supply curve and demand curve can be graphed and added to determine the market equilibrium price and projected market size. There are many ways a company can combat if the market equilibrium pricing is not suitable for a process. One of these ways is changing the market that the company is selling to. Instead of selling industrial grade product there may be markets for pharmaceutical grade or food grade that would allow for a company to sell their product at higher margins. Another avenue to look into is changing the geographic market being sold to. Rarely is there a global synchronous market, but rather a variation depending on where in the world the product is being sold. It is possible that a company could make more money by dedicating their sales to the Asian market as opposed to the US or vise versa. <br />
===Price Data Sources===<br />
There are many resources when trying to determine the price of a chemical or utility. This are important for looking at current pricing information as well as historical data that can be used for forecasting purposes. <br />
====Internal Company Forecasts====<br />
Large companies will often have the marketing or development departments develop a forecasting database that can be used internally in the company. Forecasts of this magnitude will often have multiple scenarios and projects that are evaluated under the given parameters. Companies may even license these forecasts to other companies for high fees if they desire. <br />
<br />
[[File:Capture.JPG]]<br />
<br />
Table 1 provides common industry acronyms that are used to indicate certain key words when determining pricing information.<br />
<br />
====Trade Journals====<br />
There are also many publications that report pricing data weekly. ''ICIS Chemical Business Americas'' used to publish the prices for hundreds of chemicals but have more recently changed their data to an online database that requires a subscription. This service is very expensive, but necessary for many companies. ''Oil and Gas Journal'' publishes the market prices of many crude oils and other petrochemicals using data from several continents. This journal also provides margin data for many refineries and plants on a monthly basis. ''Chemical Week'' provides the spot and contract prices for 22 chemicals in the US and European markets.<br />
<br />
====Consultants====<br />
If trade journals are not adequate for the information needed, some companies will contract consultants to do deep research into the subject. Consultants are excellent resources for providing economic and marketing information but come at a large price. There are several companies that provide this type of service but some of the larger firms include: ''Purvin and Gertz'', ''Cambidge Energy Research Associates'', ''Chemical Markets Associates Inc.'', and ''SRI: The Chemical Economics Handbook''<br />
<br />
====Online Brokers and Suppliers====<br />
Often time price data can be supplied by the supplier themselves and using online directories. Restraint should be used when quoting these prices however because they are often spot prices that are much higher than what would be expected from bulk contract supplying.<br />
<br />
==Example Case: Estimating Cost of Production==<br />
<br />
Use the following information to estimate the manufacturing cost of a plant producing 120*10^6 lb/year with a product price of $0.20/lb. <br />
<br />
:Fixed Capital: $15,000,000<br />
:Working Capital: $3,000,000<br />
:'''Fixed and Working Capital = FC + WC = $18,000,000'''<br />
:Raw Material Cost: $9,600,000/yr<br />
:Utilities: $1,440,000/yr<br />
:Labor: $1,800,000/yr<br />
:Maintenance (6% yr f.c.): $900,000/yr<br />
:Supplies (2% yr f.c.): $300,000/yr<br />
:Depreciation (8%/yr): $1,200,000/yr<br />
:Taxes, insurance (3%/yr): $450,000/yr<br />
:'''Total Manufacturing Cost = RMC + U + L + M + S + D + T = $15,690,000/yr'''<br />
:'''Gross Sales = Production * Product price = $24,000,000/yr'''<br />
:'''Gross Profit = Gross Sales - Manufacturing Cost = $8,310,000/yr'''<br />
<br />
==Market Effects on Process Design==<br />
Process design, like most other things, is inherently dependent on global markets, and the economy as a whole. Process design can depend on locational markets. For example, if an expected long run exchange rate between two specific currencies makes production of a specific chemical more profitable in a colder climate such as Russia as opposed to a warmer climate such as Mexico, this may have an effect on how the process to produce this specific chemical should be designed. Relative inflation rates of different countries can have similar effects on process design. Additionally, markets for key production inputs can have an effect on process design. If a company was designing a large scale production plant which required massive amounts of steel, the global market for steel (or to a lesser extent, iron mines) would impact how/where this company would choose to design and build their process. Also worth noting is the fact that the state of the economy as a whole may impact process design as well. If the economy is in a recession, funds are likely to be more tightly managed within a company. Thusly, it may make sense for a process engineer to design a process that operates on a smaller scale. Not only will this decrease the cost of running the process (lower utilities costs, lower costs for process inputs), but since less product will be produced, this will also reduce the risk of having the company running an inventory. This is a positive, as running an inventory can lead to losing even more money during a recession. Conversely, a healthy economy may encourage the design of facilities with larger production capabilities. This section aims to look at several recent examples of macroeconomic markets affecting process design, similar to the hypothetical situations described above.<br />
<br />
Back in 2009, when the USA, UK, and many other first world countries were facing a harsh recession, both the profitability and scale of many production processes diminished. The Royal Society of Chemistry found that many chemical companies with plants located in England, when faced with the recession, elected to shut down production entirely rather than simply slow down production (RSC, 2009). Many of these plant closings were accompanied with the construction of new plants in areas of the world less affected by the recession, including the Middle East. Many notable international companies, such as Dow, were involved in this migration from countries entrenched in recession to countries where production would be more economical. While this may seem obvious, it is still worth noting that a process engineer needs to take into account the state of the economy in the location where a production facility is looking to be built.<br />
<br />
Building off of what the Royal Society of Chemistry found, a 2010 KPMG case study aimed to further analyze the effects of the recession on chemical production plants. Though the KPMG case study was based on United States production, the economic climate in the UK was similar to that of the United States, so similar conclusions can be drawn from both studies. What KPMG found in the United States in 2010 backed up what the Royal Society of Chemistry found in England in 2009. The United States was seeing a similar outsourcing of both old production facilities that were being shut down and new production facilities that, though initially proposed for construction in the United States, were being outsourced. However, KPMG went further in depth to analyze how specific markets in the United States were affecting chemical production plants. They found that the domestic auto industry and the construction sector heavily influenced the number and scale of chemical plants being proposed for construction in the USA. This makes sense, as these are two of the largest markets for chemicals in the USA. By 2010, the domestic auto industry was rebounding, and this was found to be largely correlated with an increase in demand for US chemicals (KPMG, 2010). This highlights that a process engineer must consider what sectors use the output of his process (in this scenario, various chemicals) as an input to their process. Analyzing these specific markets in several competitive locations could provide the final input into determining where to build a production facility.<br />
<br />
Another case study, conducted by CNN, analyzed the world market for rare earth minerals. Rare earth minerals are vital to the preparation of catalysts, which impacts a large portion of the chemical engineering industry. Additionally, rare earth minerals are used extensively in a wide variety of consumer products, including but not limited to hybrid cars and smartphone chips. Any process engineer looking to design a plant where production utilizes rare earth minerals should analyze this global market in order to influence his decisions that go into designing the process. The case study conducted by CNN highlights the fact that China is the dominant player in the global market for rare earth minerals (Yan, 2015). Per National Center for Policy Analysis, with whom CNN consulted during their case study, China controls about 95% of global rare earth mineral production, and holds half of the world’s resources of these metals. A specific subset of rare earth minerals, rare earth oxides, are vital components of many catalysts. Figure 1 below shows just how strong of a monopoly China has on these rare earth oxides in the global market. The supply crunch brought on by China forced the lone United States producer of rare earth metals, Molycorp, into bankruptcy. China having a near monopoly on the global market for rare earth metals means that they can exert their dominance on the market in several fashions. One such example came in 2010, when Beijing abruptly reducing their export quota for rare earth minerals lead to skyrocketing prices. A process engineer who’s proposed design includes the use of any catalyst using these rare earth metals must take all of these possibilities into account. In this example, does the influence China has on the global rare earth metals market make it more sensible to build a plant in China? Or maybe the volatility of the prices is too concerning, which could lead to the process engineer being forced to redesign his process without the use of a catalyst. Though this would certainly reduce product yield, it could be the case that the markets for the metals that make the necessary catalyst render the process without the catalyst more profitable on a per unit basis. This rare earth metal example is just one of many; the overall lesson is that a process engineer must evaluate the global markets for all important inputs to his process, as the profitability of different designs will be heavily influenced by these markets.<br />
<br />
[[File:reo.jpg|frame|center|border|<div align=center> Figure 1: Global rare earth oxide production trends from 1956-2010 (Tse, 2011) <div>]]<br />
<div align=left><br />
<br />
<br />
A third case study conducted by the Economic Policy Institute analyzed broad world market effects in the form of trade agreements. A trade agreement is a tax, tariff and trade treaty between two or more nations that often includes investment guarantees. One prominent example of a trade agreement is the North American Free Trade Agreement (NAFTA), which was signed in 1993. NAFTA is a trilateral rules based trade bloc between the United States, Canada and Mexico. One of the biggest effects of NAFTA was the movement of chemical and manufacturing plants from the United States to Canada and Mexico. In the first decade of the act (1993-2002), approximately 880,000 U.S. jobs were lost to Canada and Mexico (Scott, 2003). As a result of the influx of jobs, real wages dropped in Mexico, causing the operation of plants and processing facilities in Mexico to be more profitable than in the United States. This in turn lead to more factories moving to Mexico, and the cycle perpetuated. NAFTA is just one example of many trade agreements which had this effect. Globally, trade agreements between developed and developing countries leads to a migration of jobs (and the accompanying lowering of wages) to the developing countries. A process engineer needs to analyze trade agreements between the country in which he works and any countries with which his country has trade agreements. Not only will he likely find cheaper labor in other countries, but depending on the trade agreement, there may be tax breaks, import deals, or a variety of other money saving clauses written in the trade agreement. Overall, regardless of whatever the process may be, a process engineer can save his company millions of dollars by simply analyzing trade agreements and optimizing plant location.<br />
<br />
====Example====<br />
<br />
As a final example to sum up some of these global market impacts on a process engineer’s decisions, imagine the following hypothetical process: the production of a new biofuel using ethanol as an input, a catalyzed packed bed reactor to produce the biofuel with carbon dioxide being one of the byproducts, and a distillation column required to purify the product. Let us say that the process engineer is deciding between China and the United States as a location for the plant.<br />
<br />
First of all, seeing as how ethanol is the main feed, if is obviously necessary to look into the global market for ethanol. As can be seen below in Figure 2, the United States and Brazil are the two dominant players in terms of production. This would mean that, despite the relative volatility of ethanol prices, it is reasonable to expect that ethanol would be the cheapest in either of these two countries, since acquiring ethanol would not include shipping costs and/or import tariffs. This favors the United States as the location for this plantAs a quick tangent, since a large amount of ethanol is derived from corn, a process engineer should look into future expected trends for corn globally.<br />
<br />
Next comes the catalyzed reactor. As highlighted previously, China dwarfs the United States in terms of control of the rare earth metals that go into producing many catalysts. Due to the dominance of China in the rare earth oxide catalysts market, this would favor China as a location for production. Based on the fact that the two important inputs for this process (ethanol and the catalyst) aren’t both cheaper in either the U.S. or China, a process engineer could alter the process to favor one location. For example, if plant production in the U.S. was desirable, perhaps a lesser quality catalyst without as many rare earth metals could be utilized, which would require a larger amount of ethanol in the feed to meet a desired product quantity. This would cut costs on the catalyst (which is much more expensive in the U.S.) while the increase in costs for ethanol (due to increased quantity) would not outweigh the savings on catalyst.<br />
<br />
The two main products exiting this hypothetical reactor are product biofuel and byproduct carbon dioxide. China recently announced that they will implement a national cap-and-trade system for greenhouse gas emissions in 2017, whereas the United States currently has no policy for taxing greenhouse gas emissions. However, it is highly likely that the United States will soon implement a carbon tax policy similar to Canada’s successful carbon tax policy introduced in 2008. Let us suppose for the sake of this example that a carbon tax will be implemented in the United States in 2017, the same time as China’s policy will come into effect. The cap-and-trade system in China and carbon tax system in the U.S. are very different policies. A carbon tax imposes a flat tax on each unit of carbon dioxide emitted, whereas a cap-and-trade system sets a maximum level of pollution and sells emissions permits to companies allowing them to emit up to the maximum level. Depending on the scale of the process and the quantity of carbon dioxide being released, one policy might be much more favorable, and thusly a process engineer must look into this further.<br />
<br />
A final aspect to consider is simply the costs associated with any plant. Though there are many costs associated with building and running a plant, two large costs will be analyzed here. The cost of steel required to build the plant, and the cost of labor required to operate the plant. Even though the price of steel is relatively volatile, the per unit costs in China and the U.S. have remained quite close to one another. As for labor, costs are generally lower in China overall, but a process engineer must look into specific salaries for each position at the production facility. Even though China as a whole has less expensive labor, a plant operator for biofuels production may be a much higher paying position in China than it is in the U.S. Overall, taking all of these factors into account, it may make sense for a process engineer to either favor production in one country over another, or alter the process in order to favor one country if he is not in control of where the production facility is being built.<br />
<br />
The above analyses cover only a fraction of what a process engineer must consider. The analyses conducted highlight some of the major components, but there are many others that must be considered as well. Global market effects are endless, and play a large part in helping a good process engineer determine how to design a process and where to build a production facility.<br />
<br />
[[File:Ethanol.JPG|frame|center|border|<div align=center> Figure 2: Global ethanol production from 2007-2011 (evsroll.com) <div>]]<br />
<br />
<br />
[[File:Steel.JPG|frame|center|border|<div align=center> Figure 3: Global steel prices from 2006-2012 (Gue, 2012) <div>]]<br />
<div align=left><br />
<br />
==Conclusion==<br />
Thermodynamics and kinetics are essential to designing an operational plant, but at the end of the day profits and margins are what make plants go from the engineering paper pad to operating continuously. Before any ground is broken, estimation of production costs and revenues are absolutely necessary to assure CEO's and shareholders that this process is a profitable and worth while venture. There are many avenues to achieve these answers with some being more accurate than others. The best indicator of these answers will be in pilot plant design which will provide appropriate estimations for scaled up processes.<br />
<br />
==References==<br />
<br />
Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997. <br />
<br />
Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.<br />
<br />
Seider WD, Seader JD, Lewin DR. Process Design Principles: Synthesis, Analysis, and Evaluation. 3rd ed. New York: Wiley; 2004.<br />
<br />
Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.<br />
<br />
Turton R, Bailie RC, Whiting WB, Shaewitz JA, Bhattacharyya D. Analysis, Synthesis, and Design of Chemical Processes. 4th ed. Upper Saddle River: Prentice-Hall; 2012.<br />
<br />
Chemicals Sector Struggles in Recession. Royal Society of Chemistry, rsc.org. July 29, 2009<br />
<br />
The Outlook for the US Chemical Industry. KPMG, kpmg.com. 2010.<br />
<br />
Yan S. China is About to Tighten its Grip on Rare Earth Minerals. CNN, money.cnn.com. June 5, 2015<br />
<br />
Tse P. China’s Rare-Earth Industry. USGS. pubs.usgs.gov. 2011. Accessed February 21, 2016<br />
<br />
Scott R.E. The High Price of “Free” Trade: NAFTA’s Failure has Cost the United States Jobs Across the Nation. Economic Policy Institute, www.epi.org. November 17, 2003<br />
<br />
Renewable Energy Ethanol. EVs Rock. http://evsroll.com/Renewable_Energy_Ethanol.html. Accessed February 21, 2016<br />
<br />
Gue E. Where Steel Prices are Headed. Investing.com, www.investing.com. April 3, 2012. Accessed February 21, 2016</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_production_cost_and_revenue&diff=5082Estimation of production cost and revenue2016-03-11T17:32:43Z<p>WFu: </p>
<hr />
<div>Authors: Nick Pinkerton, Karen Schmidt, James Xamplas (ChE 352 in Winter 2014), Reed Kolbe (ChE 352, Winter 2016)<br />
<br />
Steward: David Chen, Fengqi You <br />
<br />
==Variable Cost of Production==<br />
Variable costs of production are dependent primarily on plant output and rate of production. There are many variables to consider when costing a plant. <br />
# Raw materials consumed<br />
# Utilities-steam, electricity, cooling water, fuel, etc.<br />
# Consumables - acids, bases, solvents, catalysts, etc.<br />
# Disposal<br />
# Shipping<br />
The majority of the variable costs for a production plant are the raw materials and utilities costs. Variable costs can be greatly cut through optimization techniques and intelligent plant design (Towler and Sinnott, 2013).<br />
===Raw Materials Cost===<br />
Calculating the annual cost of a raw material is calculated by simply multiplying the feed rate of the process by the appropriate price per volume or mass. These are the costs of chemical feed stocks required by the process. Feed stocks flow rates are obtained from PFD (Turton et al., 2013).There are several ways to optimize this cost to ensure that a process is not costing more than it should. First one should assess the actual consumption of a plant to see if it is significantly different from what should be expected based on process stoichiometry and selectivities (Towler and Sinnott, 2013). Finding may prove that a process is less efficient than it originally claimed.<br />
It is smart to benchmark a new plant design against an existing plant or pilot plant. Raw materials are typically the largest contributor to overall variable costs. For bulk chemicals and petrochemicals, raw materials represent 80-90% of the total cash cost of production (CCOP). <br />
===Utilities Cost===<br />
These are the costs of the various utilities streams required by the process. The flowrates for the utilities streams are located on the PFD (Turton et al., 2013). This includes:<br />
*Fuel gas, oil, or coal<br />
*Electric power<br />
*Steam<br />
*Cooling water<br />
*Process water<br />
*Boiler feed water<br />
*Air<br />
*Inert gas<br />
*Refrigeration<br />
<br />
Utility streams are excellent ways to streamline a process and are often indicative of how efficient of a process the project is. Process methods such as steam generation and pinch analysis can be used to greatly reduce utility costs across a plant. Further analysis of pinch analysis techniques and optimizing heat exchanger networks can be found in plant design texts such as first reference from Gavin Towler. The determination of process utility costs is often more difficult than the determination of raw material costs; however, the utilities are typically between 5-10% of CCOP (Towler and Sinnott, 2013). The cost of heating a process can be reduced by using process waste streams as fuel which consequently also reduces the need for waste disposal.<br />
<br />
===Waste Disposal Costs===<br />
These are defined as the cost of waste treatment to protect the environment (Turton et al., 2013). These are materials that cannot be recycled or sold off as by-products. Often times these streams require additives or additional treatment to meet governmental regulations. <br />
Hydrocarbon waste can often be incinerated directly to the atmosphere or used as process fuel to heat other streams in the system. Using the stream as process fuel allows the fuel value of the stream to be recovered into the system. The substituted value can be calculated by multiplying the conventional fuel price by the heat of combustion of the waste stream. <br />
<br />
<math>P_{WFV} = P_F * \Delta H_C^o</math><br />
<br />
where <br />
<math>P_{WFV}</math> = waste value of fuel ($/lb or $/kg)<br />
<br />
<math>P_F</math> = price of fuel ($/MMBtu or $/GJ)<br />
<br />
<math>\Delta H_C^o</math> = heat of combustion (MMBtu/lb or GJ/kg)<br />
<br />
Dilute aqueous streams must be sent to wastewater treatment typically prior to purging from the plant. Acidic or basic wastes are neutralized prior to treatment by salting out the acid or base. The cost of wastewater treatment is typically about $6 per 1000 gal but this is only an estimate that doesn't account for regional charges (Towler and Sinnott, 2013).<br />
<br />
Solid waste treatment can typically be sent to a landfill at a cost of approximately $50/ton (Towler and Sinnott, 2013). <br />
<br />
Hazardous wastes arise from the production of concentrated liquid streams that cannot be incinerated. Hazardous wastes should be avoided if possible, but that is not always feasible for some processes. The cost of hazardous waste disposal is strongly dependent on the location of the plant, the plants proximity to waste disposal plants and the degree of hazard of the waste.<br />
<br />
==Fixed Cost of Production==<br />
Fixed costs are those whose amounts are independent of production rates. Much of these costs are personnel salaries, taxes, insurance, and legal payments.<br />
===Labor Costs===<br />
These are the costs attributed to the personnel required to operate the process plant (Turton et al., 2013).<br />
<br />
The number of operators required per shift, <math>N_{OL}</math> can be estimated by<br />
<br />
<math>N_{OL}=(6.29+31.7P^2+0.23N_{np})^{0.5}</math><br />
<br />
where <math>P</math> is the number of processing steps involving particulate solids and <math>N_{np}</math> is the number of other processing steps (Turton et al., 2013). For each of the <math>N_{OL}</math> operators per 8-hour shift, approximately 4.5 operators must be hired for a plant that runs 24 hours per day, to account for the 3 shifts per day and the 3 weeks of leave typically taken by each operator per year (Turton et al., 2013). The salary for a chemical plant operator varies by location, and the estimator should look up the average value for the area.<br />
<br />
===Maintenance Costs===<br />
These are the costs associated with labor and materials necessary to maintain plant production. An estimate of these are 6% of the fixed capital investment (Turton et al., 2013).<br />
<br />
===Research and Development===<br />
These are the costs of research done in developing the process and/or products. This includes salaries for researchers as well as funds for research related equipment and supplies. An estimate of these costs are 5% of the total manufacturing cost (Turton et al., 2013).<br />
<br />
===Taxes and Insurance===<br />
Taxes vary by location, but a first estimate of property taxes and liability insurance is 3% of the fixed capital investment (Turton et al., 2013).<br />
<br />
===Plant Overhead===<br />
Overhead costs are the miscellaneous but necessary costs of running a business, including payroll, employee benefits, and janitorial services. This may be estimated as 70% of the operating labor costs, added to 4% of the fixed capital costs (Turton et al., 2013).<br />
<br />
===Licensing and Royalties===<br />
The costs of paying for the use of intellectual property clearly varies, but an estimate that may be used is 3% of the total manufacturing cost (Turton et al., 2013).<br />
<br />
==Revenues==<br />
The revenues of a process are the income earned form sales of the main products and the by-products. Revenue can be impacted by market fluctuations and production rates.<br />
===By-Product Revenues===<br />
Besides selling the main product from a process, by-products from separations and reactions can also be valuable in the market. Often it is more difficult to decide which by-products to recover and purify than it is to make decisions on the main product. <br />
<br />
By-products made in stoichiometric ratios from reactions must be either sold off or managed through waste disposal. Other by-products are sometimes produced through feed impurities or by nonselective reactions. There are several potential valuable by-products from a process: <br />
# Materials produced in stoichiometric quantities by the reactions that create the main product. If they are not recovered then the waste disposal expenses will be large.<br />
# Components that are produced in high yield by side reactions. <br />
# Components formed in high yield from feed impurities. Many sulfurs are produced as a by-product of fuels manufacture.<br />
# Components that are produced in low yield but have high value. An example includes acetophenone which is recovered as a by-product of phenol manufacture.<br />
# Degraded consumables (e.g. solvents, etc.) that have reuse value.<br />
<br />
A rule of thumb that can be used for preliminary screening of by-products for large plants is that for by-product recovery to be economically feasible the net benefit must be greater than $200,000 a year. A net benefit can be calculated by adding the possible resale value of the by-product and the avoided waste disposal cost (Towler and Sinnott, 2013). <br />
<br />
===Margin===<br />
The gross margin of a process is defined as the sum of product and by-product revenues minus the raw material cost. <br />
<br />
Gross margin = Revenues - Raw materials costs<br />
<br />
Because raw materials are most often the most expensive variable cost of a process, the gross margin is a good gauge as to what the total profitability of a process will be. Raw materials and product pricing are often subject to high degrees of variability which can be difficult to forecast. The size of margins are highly versatile depending on the <br />
industry. For many petrochemical industries the margin may be only 10%; however, for industries such as food additives and pharmaceuticals the margins are generally much higher (Towler and Sinnott, 2013).<br />
<br />
===Profits===<br />
There are several standards for calculating company profits. The cash cost of production (CCOP) is the sum of the fixed and variable production costs. <br />
<br />
<math>CCOP = VCOP + FCOP</math><br />
<br />
where <math>VCOP</math> is the variable cost of production and <math>FCOP</math> is the fixed cost of production. <br />
<br />
Gross profit, which should not be confused with gross margin, is then calculated by the following equation,<br />
<br />
<math>Gross\ profit = Main\ product\ revenues - CCOP</math><br />
<br />
Finally profit can be calculated by subtracting the income taxes that the plant would be subject to depending on the tax code of the county the plant is located in. <br />
<br />
<math>Net\ profit = gross\ profit - taxes</math><br />
<br />
==Pricing Products and Raw Materials==<br />
The revenues and costs of a project are vital to determining its economic feasibility. To calculate these values one needs to multiply the respective product and feed streams by their respective prices. The major difficulty of this process is determining the prices that should be used in this formula. When analyzing a plant, not only do the current prices need to be acknowledged but also the stability of the market to forecast future fluctuations and deviations. <br />
===Pricing Fundamentals===<br />
The pricing of a substance is determined by the fundamental economic principles of supply and demand. A supply curve and demand curve can be graphed and added to determine the market equilibrium price and projected market size. There are many ways a company can combat if the market equilibrium pricing is not suitable for a process. One of these ways is changing the market that the company is selling to. Instead of selling industrial grade product there may be markets for pharmaceutical grade or food grade that would allow for a company to sell their product at higher margins. Another avenue to look into is changing the geographic market being sold to. Rarely is there a global synchronous market, but rather a variation depending on where in the world the product is being sold. It is possible that a company could make more money by dedicating their sales to the Asian market as opposed to the US or vise versa. <br />
===Price Data Sources===<br />
There are many resources when trying to determine the price of a chemical or utility. This are important for looking at current pricing information as well as historical data that can be used for forecasting purposes. <br />
====Internal Company Forecasts====<br />
Large companies will often have the marketing or development departments develop a forecasting database that can be used internally in the company. Forecasts of this magnitude will often have multiple scenarios and projects that are evaluated under the given parameters. Companies may even license these forecasts to other companies for high fees if they desire. <br />
<br />
[[File:Capture.JPG]]<br />
<br />
Table 1 provides common industry acronyms that are used to indicate certain key words when determining pricing information.<br />
<br />
====Trade Journals====<br />
There are also many publications that report pricing data weekly. ''ICIS Chemical Business Americas'' used to publish the prices for hundreds of chemicals but have more recently changed their data to an online database that requires a subscription. This service is very expensive, but necessary for many companies. ''Oil and Gas Journal'' publishes the market prices of many crude oils and other petrochemicals using data from several continents. This journal also provides margin data for many refineries and plants on a monthly basis. ''Chemical Week'' provides the spot and contract prices for 22 chemicals in the US and European markets.<br />
<br />
====Consultants====<br />
If trade journals are not adequate for the information needed, some companies will contract consultants to do deep research into the subject. Consultants are excellent resources for providing economic and marketing information but come at a large price. There are several companies that provide this type of service but some of the larger firms include: ''Purvin and Gertz'', ''Cambidge Energy Research Associates'', ''Chemical Markets Associates Inc.'', and ''SRI: The Chemical Economics Handbook''<br />
<br />
====Online Brokers and Suppliers====<br />
Often time price data can be supplied by the supplier themselves and using online directories. Restraint should be used when quoting these prices however because they are often spot prices that are much higher than what would be expected from bulk contract supplying.<br />
<br />
==Example Case: Estimating Cost of Production==<br />
<br />
Use the following information to estimate the manufacturing cost of a plant producing 120*10^6 lb/year with a product price of $0.20/lb. <br />
<br />
:Fixed Capital: $15,000,000<br />
:Working Capital: $3,000,000<br />
:'''Fixed and Working Capital = FC + WC = $18,000,000'''<br />
:Raw Material Cost: $9,600,000/yr<br />
:Utilities: $1,440,000/yr<br />
:Labor: $1,800,000/yr<br />
:Maintenance (6% yr f.c.): $900,000/yr<br />
:Supplies (2% yr f.c.): $300,000/yr<br />
:Depreciation (8%/yr): $1,200,000/yr<br />
:Taxes, insurance (3%/yr): $450,000/yr<br />
:'''Total Manufacturing Cost = RMC + U + L + M + S + D + T = $15,690,000/yr'''<br />
:'''Gross Sales = Production * Product price = $24,000,000/yr'''<br />
:'''Gross Profit = Gross Sales - Manufacturing Cost = $8,310,000/yr'''<br />
<br />
==Market Effects on Process Design==<br />
Process design, like most other things, is inherently dependent on global markets, and the economy as a whole. Process design can depend on locational markets. For example, if an expected long run exchange rate between two specific currencies makes production of a specific chemical more profitable in a colder climate such as Russia as opposed to a warmer climate such as Mexico, this may have an effect on how the process to produce this specific chemical should be designed. Relative inflation rates of different countries can have similar effects on process design. Additionally, markets for key production inputs can have an effect on process design. If a company was designing a large scale production plant which required massive amounts of steel, the global market for steel (or to a lesser extent, iron mines) would impact how/where this company would choose to design and build their process. Also worth noting is the fact that the state of the economy as a whole may impact process design as well. If the economy is in a recession, funds are likely to be more tightly managed within a company. Thusly, it may make sense for a process engineer to design a process that operates on a smaller scale. Not only will this decrease the cost of running the process (lower utilities costs, lower costs for process inputs), but since less product will be produced, this will also reduce the risk of having the company running an inventory. This is a positive, as running an inventory can lead to losing even more money during a recession. Conversely, a healthy economy may encourage the design of facilities with larger production capabilities. This section aims to look at several recent examples of macroeconomic markets affecting process design, similar to the hypothetical situations described above.<br />
<br />
Back in 2009, when the USA, UK, and many other first world countries were facing a harsh recession, both the profitability and scale of many production processes diminished. The Royal Society of Chemistry found that many chemical companies with plants located in England, when faced with the recession, elected to shut down production entirely rather than simply slow down production (RSC, 2009). Many of these plant closings were accompanied with the construction of new plants in areas of the world less affected by the recession, including the Middle East. Many notable international companies, such as Dow, were involved in this migration from countries entrenched in recession to countries where production would be more economical. While this may seem obvious, it is still worth noting that a process engineer needs to take into account the state of the economy in the location where a production facility is looking to be built.<br />
<br />
Building off of what the Royal Society of Chemistry found, a 2010 KPMG case study aimed to further analyze the effects of the recession on chemical production plants. Though the KPMG case study was based on United States production, the economic climate in the UK was similar to that of the United States, so similar conclusions can be drawn from both studies. What KPMG found in the United States in 2010 backed up what the Royal Society of Chemistry found in England in 2009. The United States was seeing a similar outsourcing of both old production facilities that were being shut down and new production facilities that, though initially proposed for construction in the United States, were being outsourced. However, KPMG went further in depth to analyze how specific markets in the United States were affecting chemical production plants. They found that the domestic auto industry and the construction sector heavily influenced the number and scale of chemical plants being proposed for construction in the USA. This makes sense, as these are two of the largest markets for chemicals in the USA. By 2010, the domestic auto industry was rebounding, and this was found to be largely correlated with an increase in demand for US chemicals (KPMG, 2010). This highlights that a process engineer must consider what sectors use the output of his process (in this scenario, various chemicals) as an input to their process. Analyzing these specific markets in several competitive locations could provide the final input into determining where to build a production facility.<br />
<br />
Another case study, conducted by CNN, analyzed the world market for rare earth minerals. Rare earth minerals are vital to the preparation of catalysts, which impacts a large portion of the chemical engineering industry. Additionally, rare earth minerals are used extensively in a wide variety of consumer products, including but not limited to hybrid cars and smartphone chips. Any process engineer looking to design a plant where production utilizes rare earth minerals should analyze this global market in order to influence his decisions that go into designing the process. The case study conducted by CNN highlights the fact that China is the dominant player in the global market for rare earth minerals (Yan, 2015). Per National Center for Policy Analysis, with whom CNN consulted during their case study, China controls about 95% of global rare earth mineral production, and holds half of the world’s resources of these metals. A specific subset of rare earth minerals, rare earth oxides, are vital components of many catalysts. Figure 1 below shows just how strong of a monopoly China has on these rare earth oxides in the global market. The supply crunch brought on by China forced the lone United States producer of rare earth metals, Molycorp, into bankruptcy. China having a near monopoly on the global market for rare earth metals means that they can exert their dominance on the market in several fashions. One such example came in 2010, when Beijing abruptly reducing their export quota for rare earth minerals lead to skyrocketing prices. A process engineer who’s proposed design includes the use of any catalyst using these rare earth metals must take all of these possibilities into account. In this example, does the influence China has on the global rare earth metals market make it more sensible to build a plant in China? Or maybe the volatility of the prices is too concerning, which could lead to the process engineer being forced to redesign his process without the use of a catalyst. Though this would certainly reduce product yield, it could be the case that the markets for the metals that make the necessary catalyst render the process without the catalyst more profitable on a per unit basis. This rare earth metal example is just one of many; the overall lesson is that a process engineer must evaluate the global markets for all important inputs to his process, as the profitability of different designs will be heavily influenced by these markets.<br />
<br />
[[File:reo.jpg|frame|center|border|<div align=center> Figure 1. a) Flash conditions with respect to pressure in the last drum. b) Process metrics with respect to pressure in the last drum. <div>]]<br />
<div align=left><br />
<br />
<br />
A third case study conducted by the Economic Policy Institute analyzed broad world market effects in the form of trade agreements. A trade agreement is a tax, tariff and trade treaty between two or more nations that often includes investment guarantees. One prominent example of a trade agreement is the North American Free Trade Agreement (NAFTA), which was signed in 1993. NAFTA is a trilateral rules based trade bloc between the United States, Canada and Mexico. One of the biggest effects of NAFTA was the movement of chemical and manufacturing plants from the United States to Canada and Mexico. In the first decade of the act (1993-2002), approximately 880,000 U.S. jobs were lost to Canada and Mexico (Scott, 2003). As a result of the influx of jobs, real wages dropped in Mexico, causing the operation of plants and processing facilities in Mexico to be more profitable than in the United States. This in turn lead to more factories moving to Mexico, and the cycle perpetuated. NAFTA is just one example of many trade agreements which had this effect. Globally, trade agreements between developed and developing countries leads to a migration of jobs (and the accompanying lowering of wages) to the developing countries. A process engineer needs to analyze trade agreements between the country in which he works and any countries with which his country has trade agreements. Not only will he likely find cheaper labor in other countries, but depending on the trade agreement, there may be tax breaks, import deals, or a variety of other money saving clauses written in the trade agreement. Overall, regardless of whatever the process may be, a process engineer can save his company millions of dollars by simply analyzing trade agreements and optimizing plant location.<br />
<br />
====Example====<br />
<br />
As a final example to sum up some of these global market impacts on a process engineer’s decisions, imagine the following hypothetical process: the production of a new biofuel using ethanol as an input, a catalyzed packed bed reactor to produce the biofuel with carbon dioxide being one of the byproducts, and a distillation column required to purify the product. Let us say that the process engineer is deciding between China and the United States as a location for the plant.<br />
<br />
First of all, seeing as how ethanol is the main feed, if is obviously necessary to look into the global market for ethanol. As can be seen below in Figure 2, the United States and Brazil are the two dominant players in terms of production. This would mean that, despite the relative volatility of ethanol prices, it is reasonable to expect that ethanol would be the cheapest in either of these two countries, since acquiring ethanol would not include shipping costs and/or import tariffs. This favors the United States as the location for this plantAs a quick tangent, since a large amount of ethanol is derived from corn, a process engineer should look into future expected trends for corn globally.<br />
<br />
Next comes the catalyzed reactor. As highlighted previously, China dwarfs the United States in terms of control of the rare earth metals that go into producing many catalysts. Due to the dominance of China in the rare earth oxide catalysts market, this would favor China as a location for production. Based on the fact that the two important inputs for this process (ethanol and the catalyst) aren’t both cheaper in either the U.S. or China, a process engineer could alter the process to favor one location. For example, if plant production in the U.S. was desirable, perhaps a lesser quality catalyst without as many rare earth metals could be utilized, which would require a larger amount of ethanol in the feed to meet a desired product quantity. This would cut costs on the catalyst (which is much more expensive in the U.S.) while the increase in costs for ethanol (due to increased quantity) would not outweigh the savings on catalyst.<br />
<br />
The two main products exiting this hypothetical reactor are product biofuel and byproduct carbon dioxide. China recently announced that they will implement a national cap-and-trade system for greenhouse gas emissions in 2017, whereas the United States currently has no policy for taxing greenhouse gas emissions. However, it is highly likely that the United States will soon implement a carbon tax policy similar to Canada’s successful carbon tax policy introduced in 2008. Let us suppose for the sake of this example that a carbon tax will be implemented in the United States in 2017, the same time as China’s policy will come into effect. The cap-and-trade system in China and carbon tax system in the U.S. are very different policies. A carbon tax imposes a flat tax on each unit of carbon dioxide emitted, whereas a cap-and-trade system sets a maximum level of pollution and sells emissions permits to companies allowing them to emit up to the maximum level. Depending on the scale of the process and the quantity of carbon dioxide being released, one policy might be much more favorable, and thusly a process engineer must look into this further.<br />
<br />
A final aspect to consider is simply the costs associated with any plant. Though there are many costs associated with building and running a plant, two large costs will be analyzed here. The cost of steel required to build the plant, and the cost of labor required to operate the plant. Even though the price of steel is relatively volatile, the per unit costs in China and the U.S. have remained quite close to one another. As for labor, costs are generally lower in China overall, but a process engineer must look into specific salaries for each position at the production facility. Even though China as a whole has less expensive labor, a plant operator for biofuels production may be a much higher paying position in China than it is in the U.S. Overall, taking all of these factors into account, it may make sense for a process engineer to either favor production in one country over another, or alter the process in order to favor one country if he is not in control of where the production facility is being built.<br />
<br />
The above analyses cover only a fraction of what a process engineer must consider. The analyses conducted highlight some of the major components, but there are many others that must be considered as well. Global market effects are endless, and play a large part in helping a good process engineer determine how to design a process and where to build a production facility.<br />
<br />
[[File:Ethanol.JPG|frame|center|border|<div align=center> Figure 2: Global ethanol production from 2007-2011 (evsroll.com) <div>]]<br />
<br />
<br />
[[File:Steel.JPG|frame|center|border|<div align=center> Figure 3: Global steel prices from 2006-2012 (Gue, 2012) <div>]]<br />
<div align=left><br />
<br />
==Conclusion==<br />
Thermodynamics and kinetics are essential to designing an operational plant, but at the end of the day profits and margins are what make plants go from the engineering paper pad to operating continuously. Before any ground is broken, estimation of production costs and revenues are absolutely necessary to assure CEO's and shareholders that this process is a profitable and worth while venture. There are many avenues to achieve these answers with some being more accurate than others. The best indicator of these answers will be in pilot plant design which will provide appropriate estimations for scaled up processes.<br />
<br />
==References==<br />
<br />
Biegler LT, Grossmann IE, Westerberg AW. Systematic Methods of Chemical Process Design. Upper Saddle River: Prentice Hall; 1997. <br />
<br />
Peters MS, Timmerhaus KD. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw Hill; 2003.<br />
<br />
Seider WD, Seader JD, Lewin DR. Process Design Principles: Synthesis, Analysis, and Evaluation. 3rd ed. New York: Wiley; 2004.<br />
<br />
Towler G, Sinnott R. Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013.<br />
<br />
Turton R, Bailie RC, Whiting WB, Shaewitz JA, Bhattacharyya D. Analysis, Synthesis, and Design of Chemical Processes. 4th ed. Upper Saddle River: Prentice-Hall; 2012.<br />
<br />
Chemicals Sector Struggles in Recession. Royal Society of Chemistry, rsc.org. July 29, 2009<br />
<br />
The Outlook for the US Chemical Industry. KPMG, kpmg.com. 2010.<br />
<br />
Yan S. China is About to Tighten its Grip on Rare Earth Minerals. CNN, money.cnn.com. June 5, 2015<br />
<br />
Tse P. China’s Rare-Earth Industry. USGS. pubs.usgs.gov. 2011. Accessed February 21, 2016<br />
<br />
Scott R.E. The High Price of “Free” Trade: NAFTA’s Failure has Cost the United States Jobs Across the Nation. Economic Policy Institute, www.epi.org. November 17, 2003<br />
<br />
Renewable Energy Ethanol. EVs Rock. http://evsroll.com/Renewable_Energy_Ethanol.html. Accessed February 21, 2016<br />
<br />
Gue E. Where Steel Prices are Headed. Investing.com, www.investing.com. April 3, 2012. Accessed February 21, 2016</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=File:Watson_aspen.JPG&diff=5077File:Watson aspen.JPG2016-03-11T17:19:30Z<p>WFu: </p>
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<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> WCT = \frac{AR}{WC}</math><br />
<br />
WCT = Working capital turnover<br />
<br />
AR = Annual revenues($)<br />
<br />
WC Working capital($)<br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Debt Financing===<br />
<br />
As stated, the debt financing involves the issuing of bonds. Buyers of the bonds can be either individual investors or banks and others institutional investors. After the bond is sold, the company who sold the bond is now in debt to the buyer. The buyer is also known as a creditor, and has priority over a stockholder in the event of a bankruptcy in the company. Bonds can have a variety of different capital amounts, also known as bond principle. In addition, bonds can have a variety of different payback times and interest rates. An interesting application of debt financing is for seasonal products. A company could release short term bonds in order to make the stream of revenue more consistent. During times of low sales, they could sell bonds, and during times of high sales, they could pay back bonds.<br />
<br />
===Equity Financing===<br />
<br />
Equity financing is accomplished through the sale of stock, also known as equity, in the company. In other words, the company is selling ownership interest in order to achieve a certain amount of funds. Equity financing occurs throughout the lifetime of a company. In the beginning and startup of the company, angel investors and venture capitalists are the major contributors to equity financing. Both give capital in order to obtain equity in the company. The chemical industry has angel investors and venture capitalists who operate as individuals, and there are groups of individuals with the same interests who pool money in order to have a larger ownership interest. As chemical companies grow, many decide to become publicly traded, or "go public". This involves an initial public offering (IPO) and the beginning of trading of the company's stock on stock exchanges. However, private companies can still sell stocks. Mergers and takeovers can occur when one company assumes majority ownership of another company.<br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<br />
<math> DR = \frac{TD}{TA}</math><br />
<br />
DR = Debt ratio<br />
<br />
TD = Total debt($)<br />
<br />
TA = Total assets($)<br />
<br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<br />
<math> ROE = \frac{NAP}{SE}</math><br />
<br />
ROE = Return on equity<br />
<br />
NAP = Net annual profit($)<br />
<br />
SE = Stockholders equity($)<br />
<br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity. Because of this, an increasing return on equity shows that more profit is being made relative to the amount of equity invested.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return for any given cost in a company. The expected return is a combination of the dividends the company pays and the growth of the company's stock price. The cost used in this calculation is usually the stock price of the company.<br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<br />
<math> i_c = DR*i_d+(1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
===Example of Project Financing: Sadara Integrated Chemicals Project===<br />
<br />
In 2013, one of the largest financing projects in the chemicals industry occurred in the petrochemical sector. Two of the companies that corroborated in this financing project were Dow Chemical and Saudi Aramco. At the time, Dow Chemical Co. had annual sales of over 57 billion USD and produced over 5000 products in varying sectors. Saudi Aramco was on of the largest oil companies in Saudi Arabia, and a leader in many aspects of the petrochemical sector including production, refining, shipping, and even hydrocarbon exploration.<br />
<br />
The project financing for Sadara Chemical Company began in 2011 with the issuance of a sukuk, which often referred to as Islamic bonds. The bonds were sold at an interest rate of 2.95%. The term, or duration of the bond, was 15.75 years. In total, the sale of these bonds earned about 2 billion USD (Dewar). This was the debt financing aspect of the project financing.<br />
<br />
Then in 2013, Dow Chemical and Saudi Aramco contributed about 17 billion USD total. In this case, the equity financing came from two well-established chemical companies. Currently, Sadara Chemical Company is evaluated at about 20 billion USD, and Saudi Aramco has 65% ownership and Dow Chemical Company has 35% ownership (Fletcher).<br />
<br />
The next aspect of equity financing for Sadara Chemical Company will happen in 2016, when it is scheduled for Sadara to be traded publicly after an IPO. It is anticipated that about 30% ownership in the company will be floated, or traded actively (Fletcher).<br />
<br />
Because the company is not currently traded, it is difficult to apply many of the quantitative measures to study the effectiveness of the project financing techniques. In addition, for companies that have recently been started, many of these quantitative measures may be skewed because of the heavy investment on research and growth. More specifically, chemical companies often do not start full production until 2-3 years after plant construction is completed. Sadara Chemical Company is planning full-scale production in mid-2016 (Fletcher). Sadara Chemical Company provides an insight into real world project financing. In addition, the magnitude of the project gained worldwide recognition.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Dewar, John. Sadara Project Sukuk: Heralding a New Era? N.p.: Butterworths Journal of International Banking and Financial Law, Mar. 2014.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Fletcher, Phillip, et al. Sadara – Redefining the Possible. N.p.: MILBANK TWEED HADLEY & MCCLOY LLP, Sept. 2013.<br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=4559Estimation of capital2016-02-21T19:58:14Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> WCT = AR/WC</math><br />
<br />
WCT = Working capital turnover<br />
<br />
AR = Annual revenues($)<br />
<br />
WC Working capital($)<br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Debt Financing===<br />
<br />
As stated, the debt financing involves the issuing of bonds. Buyers of the bonds can be either individual investors or banks and others institutional investors. After the bond is sold, the company who sold the bond is now in debt to the buyer. The buyer is also known as a creditor, and has priority over a stockholder in the event of a bankruptcy in the company. Bonds can have a variety of different capital amounts, also known as bond principle. In addition, bonds can have a variety of different payback times and interest rates. An interesting application of debt financing is for seasonal products. A company could release short term bonds in order to make the stream of revenue more consistent. During times of low sales, they could sell bonds, and during times of high sales, they could pay back bonds.<br />
<br />
===Equity Financing===<br />
<br />
Equity financing is accomplished through the sale of stock, also known as equity, in the company. In other words, the company is selling ownership interest in order to achieve a certain amount of funds. Equity financing occurs throughout the lifetime of a company. In the beginning and startup of the company, angel investors and venture capitalists are the major contributors to equity financing. Both give capital in order to obtain equity in the company. The chemical industry has angel investors and venture capitalists who operate as individuals, and there are groups of individuals with the same interests who pool money in order to have a larger ownership interest. As chemical companies grow, many decide to become publicly traded, or "go public". This involves an initial public offering (IPO) and the beginning of trading of the company's stock on stock exchanges. However, private companies can still sell stocks. Mergers and takeovers can occur when one company assumes majority ownership of another company.<br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<br />
<math> DR = TD/TA</math><br />
<br />
DR = Debt ratio<br />
<br />
TD = Total debt($)<br />
<br />
TA = Total assets($)<br />
<br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<br />
<math> ROE = NAP/SE</math><br />
<br />
ROE = Return on equity<br />
<br />
NAP = Net annual profit($)<br />
<br />
SE = Stockholders equity($)<br />
<br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity. Because of this, an increasing return on equity shows that more profit is being made relative to the amount of equity invested.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return for any given cost in a company. The expected return is a combination of the dividends the company pays and the growth of the company's stock price. The cost used in this calculation is usually the stock price of the company.<br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<br />
<math> i_c = DR*i_d+(1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
===Example of Project Financing: Sadara Integrated Chemicals Project===<br />
<br />
In 2013, one of the largest financing projects in the chemicals industry occurred in the petrochemical sector. Two of the companies that corroborated in this financing project were Dow Chemical and Saudi Aramco. At the time, Dow Chemical Co. had annual sales of over 57 billion USD and produced over 5000 products in varying sectors. Saudi Aramco was on of the largest oil companies in Saudi Arabia, and a leader in many aspects of the petrochemical sector including production, refining, shipping, and even hydrocarbon exploration.<br />
<br />
The project financing for Sadara Chemical Company began in 2011 with the issuance of a sukuk, which often referred to as Islamic bonds. The bonds were sold at an interest rate of 2.95%. The term, or duration of the bond, was 15.75 years. In total, the sale of these bonds earned about 2 billion USD (Dewar). This was the debt financing aspect of the project financing.<br />
<br />
Then in 2013, Dow Chemical and Saudi Aramco contributed about 17 billion USD total. In this case, the equity financing came from two well-established chemical companies. Currently, Sadara Chemical Company is evaluated at about 20 billion USD, and Saudi Aramco has 65% ownership and Dow Chemical Company has 35% ownership (Fletcher).<br />
<br />
The next aspect of equity financing for Sadara Chemical Company will happen in 2016, when it is scheduled for Sadara to be traded publicly after an IPO. It is anticipated that about 30% ownership in the company will be floated, or traded actively (Fletcher).<br />
<br />
Because the company is not currently traded, it is difficult to apply many of the quantitative measures to study the effectiveness of the project financing techniques. In addition, for companies that have recently been started, many of these quantitative measures may be skewed because of the heavy investment on research and growth. More specifically, chemical companies often do not start full production until 2-3 years after plant construction is completed. Sadara Chemical Company is planning full-scale production in mid-2016 (Fletcher). Sadara Chemical Company provides an insight into real world project financing. In addition, the magnitude of the project gained worldwide recognition.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Dewar, John. Sadara Project Sukuk: Heralding a New Era? N.p.: Butterworths Journal of International Banking and Financial Law, Mar. 2014.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Fletcher, Phillip, et al. Sadara – Redefining the Possible. N.p.: MILBANK TWEED HADLEY & MCCLOY LLP, Sept. 2013.<br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=4558Estimation of capital2016-02-21T19:57:13Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> WCT = AR/WC</math><br />
<br />
WCT = Working capital turnover<br />
<br />
AR = Annual revenues($)<br />
<br />
WC Working capital($)<br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Debt Financing===<br />
<br />
As stated, the debt financing involves the issuing of bonds. Buyers of the bonds can be either individual investors or banks and others institutional investors. After the bond is sold, the company who sold the bond is now in debt to the buyer. The buyer is also known as a creditor, and has priority over a stockholder in the event of a bankruptcy in the company. Bonds can have a variety of different capital amounts, also known as bond principle. In addition, bonds can have a variety of different payback times and interest rates. An interesting application of debt financing is for seasonal products. A company could release short term bonds in order to make the stream of revenue more consistent. During times of low sales, they could sell bonds, and during times of high sales, they could pay back bonds.<br />
<br />
===Equity Financing===<br />
<br />
Equity financing is accomplished through the sale of stock, also known as equity, in the company. In other words, the company is selling ownership interest in order to achieve a certain amount of funds. Equity financing occurs throughout the lifetime of a company. In the beginning and startup of the company, angel investors and venture capitalists are the major contributors to equity financing. Both give capital in order to obtain equity in the company. The chemical industry has angel investors and venture capitalists who operate as individuals, and there are groups of individuals with the same interests who pool money in order to have a larger ownership interest. As chemical companies grow, many decide to become publicly traded, or "go public". This involves an initial public offering (IPO) and the beginning of trading of the company's stock on stock exchanges. However, private companies can still sell stocks. Mergers and takeovers can occur when one company assumes majority ownership of another company.<br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<math> DR = TD/TA</math><br />
<br />
DR = Debt ratio<br />
<br />
TD = Total debt($)<br />
<br />
TA = Total assets($)<br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<math> ROE = NAP/SE</math><br />
<br />
ROE = Return on equity<br />
<br />
NAP = Net annual profit($)<br />
<br />
SE = Stockholders equity($)<br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity. Because of this, an increasing return on equity shows that more profit is being made relative to the amount of equity invested.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return for any given cost in a company. The expected return is a combination of the dividends the company pays and the growth of the company's stock price. The cost used in this calculation is usually the stock price of the company.<br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<math> i_c = DR*i_d+(1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
===Example of Project Financing: Sadara Integrated Chemicals Project===<br />
<br />
In 2013, one of the largest financing projects in the chemicals industry occurred in the petrochemical sector. Two of the companies that corroborated in this financing project were Dow Chemical and Saudi Aramco. At the time, Dow Chemical Co. had annual sales of over 57 billion USD and produced over 5000 products in varying sectors. Saudi Aramco was on of the largest oil companies in Saudi Arabia, and a leader in many aspects of the petrochemical sector including production, refining, shipping, and even hydrocarbon exploration.<br />
<br />
The project financing for Sadara Chemical Company began in 2011 with the issuance of a sukuk, which often referred to as Islamic bonds. The bonds were sold at an interest rate of 2.95%. The term, or duration of the bond, was 15.75 years. In total, the sale of these bonds earned about 2 billion USD (Dewar). This was the debt financing aspect of the project financing.<br />
<br />
Then in 2013, Dow Chemical and Saudi Aramco contributed about 17 billion USD total. In this case, the equity financing came from two well-established chemical companies. Currently, Sadara Chemical Company is evaluated at about 20 billion USD, and Saudi Aramco has 65% ownership and Dow Chemical Company has 35% ownership (Fletcher).<br />
<br />
The next aspect of equity financing for Sadara Chemical Company will happen in 2016, when it is scheduled for Sadara to be traded publicly after an IPO. It is anticipated that about 30% ownership in the company will be floated, or traded actively (Fletcher).<br />
<br />
Because the company is not currently traded, it is difficult to apply many of the quantitative measures to study the effectiveness of the project financing techniques. In addition, for companies that have recently been started, many of these quantitative measures may be skewed because of the heavy investment on research and growth. More specifically, chemical companies often do not start full production until 2-3 years after plant construction is completed. Sadara Chemical Company is planning full-scale production in mid-2016 (Fletcher). Sadara Chemical Company provides an insight into real world project financing. In addition, the magnitude of the project gained worldwide recognition.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Dewar, John. Sadara Project Sukuk: Heralding a New Era? N.p.: Butterworths Journal of International Banking and Financial Law, Mar. 2014.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Fletcher, Phillip, et al. Sadara – Redefining the Possible. N.p.: MILBANK TWEED HADLEY & MCCLOY LLP, Sept. 2013.<br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=4346Estimation of capital2016-02-20T23:44:10Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> WCT = AR/WC</math><br />
<br />
WCT = Working capital turnover<br />
<br />
AR = Annual revenues($)<br />
<br />
WC Working capital($)<br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Debt Financing===<br />
<br />
As stated, the debt financing involves the issuing of bonds. Buyers of the bonds can be either individual investors or banks and others institutional investors. After the bond is sold, the company who sold the bond is now in debt to the buyer. The buyer is also known as a creditor, and has priority over a stockholder in the event of a bankruptcy in the company. Bonds can have a variety of different capital amounts, also known as bond principle. In addition, bonds can have a variety of different payback times and interest rates. An interesting application of debt financing is for seasonal products. A company could release short term bonds in order to make the stream of revenue more consistent. During times of low sales, they could sell bonds, and during times of high sales, they could pay back bonds.<br />
<br />
===Equity Financing===<br />
<br />
Equity financing is accomplished through the sale of stock, also known as equity, in the company. In other words, the company is selling ownership interest in order to achieve a certain amount of funds. Equity financing occurs throughout the lifetime of a company. In the beginning and startup of the company, angel investors and venture capitalists are the major contributors to equity financing. Both give capital in order to obtain equity in the company. The chemical industry has angel investors and venture capitalists who operate as individuals, and there are groups of individuals with the same interests who pool money in order to have a larger ownership interest. As chemical companies grow, many decide to become publicly traded, or "go public". This involves an initial public offering (IPO) and the beginning of trading of the company's stock on stock exchanges. However, private companies can still sell stocks. Mergers and takeovers can occur when one company assumes majority ownership of another company.<br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<math> DR = TD/TA</math><br />
<br />
DR = Debt ratio<br />
<br />
TD = Total debt($)<br />
<br />
TA = Total assets($)<br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<math> ROE = NAP/SE</math><br />
<br />
ROE = Return on equity<br />
<br />
NAP = Net annual profit($)<br />
<br />
SE = Stockholders equity($)<br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity. Because of this, an increasing return on equity shows that more profit is being made relative to the amount of equity invested.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return for any given cost in a company. The expected return is a combination of the dividends the company pays and the growth of the company's stock price. The cost used in this calculation is usually the stock price of the company.<br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<math> i_c = DR*i_d+(1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
===Case study of Project Financing: Sadara Integrated Chemicals Project===<br />
<br />
In 2013, one of the largest financing projects in the chemicals industry occured <br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=4345Estimation of capital2016-02-20T23:33:07Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> WCT = AR/WC</math><br />
<br />
WCT = Working capital turnover<br />
<br />
AR = Annual revenues($)<br />
<br />
WC Working capital($)<br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Debt Financing===<br />
<br />
As stated, the debt financing involves the issuing of bonds. Buyers of the bonds can be either individual investors or banks and others institutional investors. After the bond is sold, the company who sold the bond is now in debt to the buyer. The buyer is also known as a creditor, and has priority over a stockholder in the event of a bankruptcy in the company. Bonds can have a variety of different capital amounts, also known as bond principle. In addition, bonds can have a variety of different payback times and interest rates. An interesting application of debt financing is for seasonal products. A company could release short term bonds in order to make the stream of revenue more consistent. During times of low sales, they could sell bonds, and during times of high sales, they could pay back bonds.<br />
<br />
===Equity Financing===<br />
<br />
Equity financing is accomplished through the sale of stock, also known as equity, in the company. In other words, the company is selling ownership interest in order to achieve a certain amount of funds. In chemical companies, mergers and takeovers can occur when one company assumes majority ownership of another company. <br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<math> DR = TD/TA</math><br />
<br />
DR = Debt ratio<br />
<br />
TD = Total debt($)<br />
<br />
TA = Total assets($)<br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<math> ROE = NAP/SE</math><br />
<br />
ROE = Return on equity<br />
<br />
NAP = Net annual profit($)<br />
<br />
SE = Stockholders equity($)<br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity. Because of this, an increasing return on equity shows that more profit is being made relative to the amount of equity invested.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return for any given cost in a company. The expected return is a combination of the dividends the company pays and the growth of the company's stock price. The cost used in this calculation is usually the stock price of the company.<br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<math> i_c = (DR*i_d+((1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=4342Estimation of capital2016-02-20T23:30:02Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> WCT = AR/WC<br />
<br />
WCT: Working\ capital\ turnover<br />
<br />
AR: Annual\ revenues($)<br />
<br />
WC:Working\ capital($)<br />
</math><br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Debt Financing===<br />
<br />
As stated, the debt financing involves the issuing of bonds. Buyers of the bonds can be either individual investors or banks and others institutional investors. After the bond is sold, the company who sold the bond is now in debt to the buyer. The buyer is also known as a creditor, and has priority over a stockholder in the event of a bankruptcy in the company. Bonds can have a variety of different capital amounts, also known as bond principle. In addition, bonds can have a variety of different payback times and interest rates. An interesting application of debt financing is for seasonal products. A company could release short term bonds in order to make the stream of revenue more consistent. During times of low sales, they could sell bonds, and during times of high sales, they could pay back bonds.<br />
<br />
===Equity Financing===<br />
<br />
Equity financing is accomplished through the sale of stock, also known as equity, in the company. In other words, the company is selling ownership interest in order to achieve a certain amount of funds. In chemical companies, mergers and takeovers can occur when one company assumes majority ownership of another company. <br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<math> Debt\ ratio(DR)= Total\ debt($)/Total\ assets($)</math><br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<math> Return\ on\ equity(ROE) = Net\ annual\ profit($)/Stockholders\ equity($)</math><br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity. Because of this, an increasing return on equity shows that more profit is being made relative to the amount of equity invested.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return for any given cost in a company. The expected return is a combination of the dividends the company pays and the growth of the company's stock price. The cost used in this calculation is usually the stock price of the company.<br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<math> i_c = (DR*i_d+((1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=4009Estimation of capital2016-02-06T03:19:57Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> Working\ capital\ turnover = Annual\ revenues($)/Working\ capital($)</math><br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing usually involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<math> Debt\ ratio(DR)= Total\ debt($)/Total\ assets($)</math><br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<math> Return\ on\ equity(ROE) = Net\ annual\ profit($)/Stockholders\ equity($)</math><br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity. Because of this, an increasing return on equity shows that more profit is being made relative to the amount of equity invested.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return for any given cost in a company. The expected return is a combination of the dividends the company pays and the growth of the company's stock price. The cost used in this calculation is usually the stock price of the company.<br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<math> i_c = (DR*i_d+((1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=4008Estimation of capital2016-02-06T03:15:38Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. The working capital for a plant that produces a seasonal product may have a far larger working capital than a company that is not seasonal but has similar annual production. This is because the needs during the time when the product is in season has a more significant impact on the sales. Therefore, it is vital to ensure needs are met during the season.<br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> Working\ capital\ turnover = Annual\ revenues\ ($)/Working\ capital\ ($)</math><br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital turnover ratios for different companies over a 5-year period.<br />
<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
|+ '''Example working capital turnover ratios'''<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up and maintenance of a chemical plant, there are often different project financing methods required to cover the capital needs. The two main methods of project financing are debt and equity financing. Debt financing usually involves the issuing of bonds. Equity financing usually involves the issuing of common stock. However, most companies utilize a combination of these two methods to successfully finance a project.<br />
<br />
===Quantitative Measures===<br />
<br />
The discussion of project financing is centered on some quantitative measures that are often used to understand the growth and profitability of a company. More importantly, these measures also give insight into the health of the company beyond common metrics of revenue and costs. <br />
<br />
<math> Debt\ ratio\ (DR)= Total\ debt($)/Total\ assets\ ($)</math><br />
<br />
The debt ratio of the company can be used to understand how much risk there is on future earnings and cash flows of the company (Towler). As stated, the debt ratio is a comparison of the total debt and total assets of the company. For companies with large debt ratios, the interest that is deducted from earnings will be large because of the large amount of debt carried. <br />
<br />
<math> Return\ on\ equity\ (ROE) = Net\ annual\ profit\($)/Stockholders\ equity\ ($)</math><br />
<br />
The return on equity can be used to understand how effectively the company is managed from a fiscal point of view. As stated, the return on equity is a comparison of net annual profit and stockholder's equity.<br />
<br />
Another important measure is the cost of equity. This measure is the expected return <br />
<br />
===Cost of Capital===<br />
<br />
With the quantitative measures discussed in the previous section, an overall cost of capital can be calculated. This value is an interest rate that is the effective rate at which all of the capital is raised. This is the most useful measure for the economic evaluation of capital needed for projects.<br />
<br />
This rate can be written as:<br />
<br />
<math> i_c = (DR*i_d+((1-DR)*i_e</math><br />
<br />
where <math>i_c</math> is the cost of capital, <math>DR</math> is the debt ratio, <math>i_d</math> is the interest at which bonds are issued, and <math>i_e</math> is the cost of equity.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3996Estimation of capital2016-02-06T02:38:02Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. <br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> Working\ capital\ turnover = Annual\ revenues\ ($)/Working\ capital\ ($)</math><br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio. Below are different working capital ratios for different companies over a 5-year period.<br />
<br />
3.81 3.24 4.56 5.52 –<br />
LyondellBasell Industries N.V. 7.35 4.85 6.31 7.95 –<br />
Praxair Inc. 35.17 47.32 35.86 156.28 37.75<br />
<br />
{| class="wikitable" style="margin: 1em auto 1em auto;"<br />
! Date:<br />
! Dec 31, 2014<br />
! Dec 31, 2013<br />
! Dec 31, 2012<br />
! Dec 31, 2011<br />
! Dec 31, 2010<br />
|-<br />
| Dow Chemical Co.<br />
| 4.59 <br />
| 4.39<br />
| 4.66<br />
| 6.13<br />
| 5.24<br />
|-<br />
| E. I. du Pont de Nemours & Co.<br />
| 3.81<br />
| 3.24<br />
| 4.56<br />
| 5.52<br />
| -<br />
|-<br />
| LyondellBasell Industries N.V.<br />
| 7.35<br />
| 4.85<br />
| 6.31<br />
| 7.95<br />
| –<br />
|-<br />
| Praxair Inc. <br />
| 35.17<br />
| 47.32<br />
| 35.86<br />
| 156.28<br />
| 37.75<br />
|}<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up of a chemical plant, there are often different project financing methods required to cover these costs. <br />
<br />
===Debt Financing===<br />
<br />
<math> Debt\ ratio = Annual\ Revenues\ ($)/Working\ Capital\ ($)</math><br />
<br />
===Equity Financing===<br />
<br />
===Cost of Capital===<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
"Dow Chemical Co. (DOW) Short-term (Operating) Activity Analysis." NYSE Stock Exchange Data. Web. 5 Feb. 2016. <br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3987Estimation of capital2016-02-06T02:21:13Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. <br />
<br />
====Working Capital Turnover====<br />
<br />
Optimization of working capital is a consideration that can greatly affect the success and growth of a company. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. Simply stated:<br />
<br />
<br />
<math> Working\ capital\ turnover = Annual\ revenues\ ($)/Working\ capital\ ($)</math><br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. There exists a relationship between revenue and working capital because of the aforementioned risk of being unable to address unexpected operations needs. In cases of low working capital, replacing faulty equipment may take long periods of time. This will correspond to down time for the process and decreased production and decreased revenues. Careful risk analysis will identify how to optimize the working capital turnover ratio.<br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up of a chemical plant, there are often different project financing methods required to cover these costs. <br />
<br />
===Debt Financing===<br />
<br />
<math> Debt\ ratio = Annual\ Revenues\ ($)/Working\ Capital\ ($)</math><br />
<br />
===Equity Financing===<br />
<br />
===Cost of Capital===<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3970Estimation of capital2016-02-06T02:04:38Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. <br />
<br />
====Working Capital Turnover====<br />
<br />
The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. <br />
<br />
<br />
<math> Working\ Capital\ Turnover = (Annual\ Revenues\ ($))/(Working\ Capital\ ($))</math><br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. Careful analysis will identify <br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up of a chemical plant, there are often different project financing methods required to cover these costs. <br />
<br />
===Debt Financing===<br />
<br />
===Equity Financing===<br />
<br />
===Cost of Capital===<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3947Estimation of capital2016-02-06T00:57:36Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler). Another suggested estimation of working capital is 10-20% of annual operating costs (Garrett). Both of these should be used as initial estimates, but further analysis of the aforementioned factors will yield a more useful value for working capital.<br />
<br />
But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance, and the aforementioned estimate will not work well for all situations. A simple example of unique characteristic that affects the size of working capital is the seasonality of a product. <br />
<br />
====Working Capital Turnover====<br />
<br />
The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. <br />
<br />
<br />
<math> Working\ Capital\ Turnover = (Annual\ Revenues)/(Working\ Capital)</math><br />
<br />
<br />
The objective for any plant is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. The second method is very simple to enact; rather than holding cash as working capital, it can be used in other aspects of the plant operations. Careful analysis will identify <br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up of a chemical plant, there are often different project financing methods required to cover these costs. <br />
<br />
===Debt Financing===<br />
<br />
===Equity Financing===<br />
<br />
===Cost of Capital===<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Garrett DE. Chemical Engineering Economics. 1st ed. New York: Van Nostrand Reinhold; 1989. p. 36-72.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3943Estimation of capital2016-02-06T00:28:03Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler 9.2.3). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler 9.2). But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance. <br />
<br />
<br />
<br />
====Working Capital Turnover====<br />
<br />
The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. <br />
<br />
<math> Working\ Capital\ Turnover = (Annual\ Revenues)/(Working\ Capital)</math><br />
<br />
The objective is to maximize the working capital turnover. There are two ways to accomplish this: increase annual revenues or decrease working capital. <br />
<br />
==Project Financing==<br />
<br />
Because of the magnitude of costs associated with the start-up of a chemical plant, there are often different project financing methods required to cover these costs. <br />
<br />
===Debt Financing===<br />
<br />
===Equity Financing===<br />
<br />
===Cost of Capital===<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3940Estimation of capital2016-02-06T00:20:47Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler 9.2.3). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler 9.2). But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance. <br />
<br />
<br />
<br />
====Working Capital Turnover====<br />
<br />
The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. <br />
<br />
<math> Working\ Capital\ Turnover = (Annual Revenues)/(Working Capital)</math><br />
<br />
==Project Financing==<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3939Estimation of capital2016-02-06T00:20:30Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler 9.2.3). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler 9.2). But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance. <br />
<br />
<br />
<br />
====Working Capital Turnover====<br />
<br />
The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. <br />
<br />
<math> Working\Capital\Turnover = (Annual Revenues)/(Working Capital)</math><br />
<br />
==Project Financing==<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3938Estimation of capital2016-02-06T00:19:26Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler 9.2.3). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler 9.2). But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance. <br />
<br />
<br />
<br />
====Working Capital Turnover====<br />
<br />
The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. <br />
<br />
<math> Working\thinmuskip Capital\thinmuskip Turnover = (Annual Revenues)/(Working Capital)</math><br />
<br />
==Project Financing==<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3937Estimation of capital2016-02-06T00:17:52Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant (Towler 9.2.3). It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of feedstocks and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. There are many aspects of plant operation that are considered when making an estimation for how much working capital is needed. The value of inventory, the value of products and by-products, magnitude of accounts payable, magnitude of accounts receivable, process equipment spare parts costs. When all of these factors are taken into account, a useful estimate of working capital needed is approximately seven weeks of productions costs minus two weeks of feedstocks costs (Towler 9.2). But it must be noted that for chemical plants of different processes, the individual factors that affect working capital can have large variance. <br />
<br />
<br />
<br />
====Working Capital Turnover====<br />
<br />
The working capital turnover is a metric that is used to determine how efficiently the working capital is managed. If the working capital is too low, it may not cover the costs of operations during a particular period of time. If the working capital is too high, it can be considered cash that is not gaining interest or value. <br />
<br />
<math> Working Capital Turnover = (Annual Revenues)/(Working Capital)(</math><br />
<br />
==Project Financing==<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3936Estimation of capital2016-02-06T00:02:10Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, and Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
===Working Capital===<br />
<br />
The capital costs associated with purchasing, building, and starting up any chemical plant can be estimated with methods discussed in previous and later sections. The working capital is a distinct cost associated with maintaining operations in a plant. It is important to differentiate it with costs of outflows during design and construction. It is also different from the costs of reactants and utilities that are paid during normal operations of the plant. Many of these costs have high value, but have a characteristic of being illiquid. For example, an expensive reactor in a process may be worth 3 million USD, but it can not be sold quickly for this price in the event 3 million USD is needed. <br />
<br />
The working capital of a plant provides liquidity and flexibility as it is cash kept in reserve. It can be thought of as money that is needed to address irregularities in process operation, that may or may not be spent. <br />
<br />
====Working Capital Turnover====<br />
<br />
<math></math><br />
<br />
==Project Financing==<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3935Estimation of capital2016-02-05T23:38:08Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2016] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
<br />
===Working Capital===<br />
<br />
In addition to installation and construction costs, all equipment and buildings need maintenance. To handle this, a certain amount of capital is kept in reserve to handle maintenance costs. This is termed the "working capital" of the plant, in addition to the fixed investment. Working capital is not money that has been spent yet, but is tied up for use in maintaining the plant. Due to the time-value of money, calculating the costs associated with keeping this money but not having spent it on depreciating equipment is non straightforward.<br />
<br />
==Project Financing==<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFuhttps://processdesign.mccormick.northwestern.edu/index.php?title=Estimation_of_capital&diff=3934Estimation of capital2016-02-05T23:36:47Z<p>WFu: </p>
<hr />
<div><br><br />
<br />
Authors: Alex Chandel<sup> [2014] </sup>, Eric Jiang<sup> [2014] </sup>, Minwook Kim<sup> [2014] </sup>, Todor Kukushliev<sup> [2014] </sup>, William Lassman<sup> [2014] </sup>, and Watson Fu<sup> [2015] </sup><br />
<br />
Steward: Daniel Garcia, David Chen, Fengqi You<br />
<br />
Date Presented: 2/9/2014<br />
<br />
<br><br />
<br />
<br />
==Introduction==<br />
<br />
One of the most important aspects of determining the overall economic viability of a chemical process is determining the capital cost. In addition to the purchase price of the equipment, capital costs include delivery and installation of equipment, preparation of land for construction, salaries of contractors and construction workers, and any other costs associated with building a chemical plant. For this reason, the cost associated with process equipment is not as straightforward as the sticker price.<br />
<br />
==Components of Capital Cost==<br />
<br />
===Fixed Capital Investment===<br />
<br />
The fixed capital investment is the total cost associated with constructing the plant. This cost includes design, site remediation, purchasing process equipment, developing infrastructure, and contingency charges, and includes the raw material costs as well as labor. It is divided into four categories.<br />
<br />
====ISBL (Inside Battery Limits) Plant Costs====<br />
<br />
ISBL (Inside Battery Limits) plant costs are the cost of procuring and installing all process equipment. ISBL costs include purchasing and shipping costs of equipment, land costs, infrastructure, piping, catalysts, and any other material needed for final plant operation, or construction of the plant. ISBL costs also include any associated fees with construction such as permits, insurance, or equipment rental, even if these items are not needed once the plant is operational.<br />
<br />
ISBL is often defined as the "inner" cost of the plant, in that it is the cost associated with building the plant itself, from unloading the raw materials to shipping final products. Any costs associated with developing the plant itself is considered ISBL. It is important and relatively straightforward to obtain an estimate for the ISBL of the plant, and as other costs are often estimated based on the result of the ISBL, it is critical that this value is as accurate as possible.<br />
<br />
====OSBL (Outside Battery Limits) Plant Costs====<br />
<br />
OSBL (Outside Battery Limits), or off-site costs, are still an important component of the plant cost, but deals with calculating costs associated with off-site developments that require the plant to run. For example, if water or electricity are being utilized from the main grid, and infrastructure needs to be expanded to accommodate the chemical plant's addition to these systems, these costs are considered OSBL because they are not directly associated with elements between the input and output of the chemical plant.<br />
<br />
Other examples of OSBL costs include fencing and security, utilities such as steam or electricity generators, sewers and waste treatment, firefighting and emergency equipment, offices and laboratories, and employee amenities. These facilities and pieces of equipment are not directly affiliated with the process but are critical costs associated with constructing any work site, and are filed under OSBL cost.<br />
<br />
OSBL costs are initially estimated as a percentage of the ISBL costs. If not a lot of information ins available, a rule of thumb is to use 40% of the ISBL costs as an estimate for OSBL. However, once detailed information such as the exact site and plant layout are known, OSBL costs can be calculated in a manner similar to the ISBL costs.<br />
<br />
====Engineering Costs====<br />
<br />
Many of the steps involved in designing detailed equipment or structures onsite fall outside the scope of chemical process design. Rather than having the plant engineer do these designs anyway, a contractor is usually hired to do this design. The costs associated with generating a design, and in some cases all the way through finished fabrication and installation of equipment is filed under engineering costs. Depending on the size of the project and the amount contracted to the outside, engineering costs may include 30% of the ISBL and up to all of the OSBL, or only 10% of the ISBL. This cost depends largely on the size of the parent company, and whether or not it has in-house capability to do detailed design of the many different processes and equipment within a chemical plant.<br />
<br />
====Contingency Charges====<br />
<br />
Once costs are determined, if one could instantaneously construct the plant, then there would be no need for contingency charges. Contingency charges exist though because prices change, unanticipated costs arise, and other unexpected events can cause changes in costs. Contingency charges ensure that there is enough capital on hand to deal with these unexpected changes. Usually, contingency charges are billed to the parent organization, or of the design is done by a contractor to the contracting organization directly at the start of the project, rather than asking for increased funding mid-project. An absolute minimum for contingency charges is 10% of the ISBL and OSBL, with a more realistic value being closer to 40%.<br />
<br />
<br />
===Working Capital===<br />
<br />
In addition to installation and construction costs, all equipment and buildings need maintenance. To handle this, a certain amount of capital is kept in reserve to handle maintenance costs. This is termed the "working capital" of the plant, in addition to the fixed investment. Working capital is not money that has been spent yet, but is tied up for use in maintaining the plant. Due to the time-value of money, calculating the costs associated with keeping this money but not having spent it on depreciating equipment is non straightforward.<br />
<br />
==Accuracy and purpose of Capital Cost Estimates==<br />
The accuracy of the total cost of a project will become more accurate as the project continues. The Association for the Advancement of Cost Estimating International (AACE International) classifies five types of estimates of capital cost.<br />
# Order of Magnitude. (±30–50%) First estimation conducted for screening purposes based on cost of similar processes.<br />
# Preliminary Estimates. (±30%) Based on only a few design detail.<br />
# Definitive Estimates. (±10–30%) Improved estimation with incorporation of more equipment detail.<br />
# Detailed Estimates. (±5-10%) Incorporation of individual equipment cost.<br />
# Check Estimates. (±5–10%) Final estimation based on completed design.<br />
<br />
==Order of Magnitude Estimates==<br />
For the early stages of the design process, it is often necessary to make quick capital cost estimates of total plant cost. The accuracy of these order of magnitude estimates are usually within ±50% accuracy. The quickest and most often employed order of magnitude process scales the cost of the new design based on the cost of similar processes. <br />
<br />
Towler gives the following equation to estimate the new design cost based on values which can be found in Towler and Sinnott (2013) Table 7.1: <br />
<br />
<math>C=aS^n</math><br />
<br />
C = cost of new plant<br />
<br />
a = constants <br />
<br />
S = size parameters, based on existing plants<br />
<br />
n = exponent constant<br />
<br />
==Estimating Purchased Equipment Costs==<br />
===Sources of Equipment Cost Data===<br />
Obtaining accurate and updated equipment costs is an important matter and there are a variety of sources to obtain this information.<br />
* Engineering, Procurement, and Construction (Contractors) companies<br />
* Cost engineering department (common in large companies)<br />
* Catalog or list prices<br />
* Cost estimation software<br />
* Cost correlations<br />
* Estimate total cost based on cost of components<br />
<br />
===Cost Correlation===<br />
Cost curves can be used as preliminary estimation of equipment costs if updated cost data is not available. <br />
<br />
<math>C_e=a+bS^n</math><br />
<br />
C_e = purchased equipment on a U.S. Gulf Coast basis<br />
<br />
a,b = constants <br />
<br />
S = size parameters<br />
<br />
n = exponent constant<br />
<br />
Correlations for constants can be found in Towler's Chemical Engineering Design (Towler and Sinnott, 2013).<br />
<br />
Example: Estimate the cost of a 30 m^2 double pipe heat exchanger.<br />
C_e = 1900 + 2500*S^1.0 for S = [1 m^2, 80 m^2]<br />
C_e = $76900<br />
<br />
===Estimation based on component cost===<br />
If the process of design and construction of a piece of equipment is known, then it is preferred by professional cost estimators to estimate total cost based on the cost of materials, labor, and manufacturer profit. Estimation of cost based on component cost will allow an unbiased estimation of real cost, allowing accurate estimation as well as possible price negotiation.<br />
<br />
==Estimating Installed Costs: The Factorial Method==<br />
Before the chemical plants can be built, capital cost estimates must be made. This is done by using the factorial method. Accuracy and the reliability of the estimate will heavily depend on the availability of the data and the level of the design at the time. Lang proposed capital cost equipment by given equation:<br />
C = F * Sum(C_e)<br />
C is the total capital cost, F is the installation factor also known as Lang factor, and C_e is the cost of major equipment. Lang factor is 3.1 for solid processing plant and 4.74 for fluids processing plant. Better estimate can be made when the different factors are used for corresponding equipment. Lang factor for different equipment can be found in calibrated data chart. <br />
Usually, the above method is used as a preliminary estimate. When more detail has been acquired, installation factor are more rigorously estimated. In detailed factorial estimates, other direct costs are compounded into the Lang factor. Installation factors are usually based on a specific material for its equipment, usually carbon steel. Failure to properly correct installation factors for materials of construction is one of the most common sources of error with the factorial method. Material factor, however, does not linearly scale with the installation factor since the transportation cost, labor cost, and fabricator’s cost does not scale with the material of the equipment. Many variations of the factorial method exist as different assumptions can be made which will determine the rigorousness and the accuracy of the estimate.<br />
<br />
==Cost Escalation==<br />
<br />
Cost estimation is a method base that basis its calculation from historical data. The prices of the construction and the labor are subject to inflation; therefore, a method has to be used to update old cost data. The method relates present costs to past costs that are based on statistical digests. To get the best estimate, each job should be broken down into its components and separate indices should be used for labor and materials. A composite index for the United States process plant industry is published in the journal Chemical Engineering. For oil refinery and petrochemicals projects, the Oil and Gas Journal publishes the Nelson-Farrer Refinery Construction Index. Both indices are updated monthly and indices for forty types of equipment are updated quarterly. There are also other indices for building the plants offsite. All cost indices should be used with caution and judgment. They do not fully represent the true costs for any particular piece of equipment or plant, nor the effect of supply and demand on prices. The closer the date of the estimate made from the date of indices published, estimate is more reliable.<br />
<br />
==Location Factors==<br />
Because of the abundance of chemical engineering plants in the U.S Gulf Coast, it is often the standard for plant and equipment cost. Cost of plant construction will differ based on:<br />
* Construction Infrastructure<br />
* Labor costs<br />
* Transportation costs<br />
* Tax Rates<br />
* Exchange Rates<br />
<br />
It is common to convert cost of construction to locations other than the U.S. Gulf Coast by applying a location factor around the U.S. Gulf Coast in which: <math>\mbox{Cost of Plant Construction} = (\mbox{Cost of Plant in Gulf Coast}) \mbox{X} (\mbox{Location Factor})</math><br />
<br />
Location Factors fluctuate with currency exchange rates and time. A rule of thumb is to that every 1000 miles away from the nearest major industrial center adds 10% to the location factor. Specific location factors can be found in the most recent edition of Aspen Richardson's International Construction Cost Factor Location Manual (Costdataonline.com).<br />
<br />
==Estimating Offsite Capital Costs==<br />
<br />
As mentioned above, OSBL costs are usually estimated as a percentage of ISBL costs until detailed site information and site layout are available for design.<br />
<br />
For new sites, the OSBL costs are often estimated as a higher percentage of the ISBL due to a greater need for remediation. Especially in cases involving handling solids, OSBL costs can be as high as 100% of the ISBL cost.<br />
<br />
The other extreme is utilizing an existing, underused site with no solids handling requirement, when fabricating a low-volume specialty chemical. In these cases, OSBL will be as low as 20% of the ISLB. For most cases, however, a typical value is 40%, and will be slightly higher for new plants, lower for existing sites with high capacities.<br />
<br />
Once requirements for onsite steam and electricity are determined, more detailed design can be done. Usually, specialized suppliers install the entire utilities system, or the entire fencing system, or provide the entire firefighting service, so many of the components of OSBL capital costs are simply negotiated with contractors.<br />
<br />
If the scope of the project changes, or if the project undergoes "scope creep," it is often easier to add capacity buy purchasing additional utilities from the outside once existing utilities have been constructed. However, this can lead to rapid changes in utility costs and the engineer should be aware of scope creep, as it can quickly change a viable process into an economically undesirable one.<br />
<br />
==Computer Tools for Cost Estimating==<br />
It is difficult for smaller companies that do not specialize in process design to maintain accurate data on process costs and perform the necessary analysis for this data to be useful. Instead, most companies use costing software and other computer tools to perform economic analysis.<br />
<br />
Several computer tools by Aspen Tech are available for estimating capital costs. Aspen's Economic Evaluation Product Family builds off of its original ICARUS technology. In the aspenONE product suite, the primary capital estimation tool is Aspen Capital Cost Estimator. It couples with Aspen Economic Evaluation to provide capital evaluations during process design and operation.<br />
<br />
Some issues that have arisen in the past utilizing ICARUS, or Aspen Capital Cost Estimator are as follows:<br />
<br />
*Mapping equipment from process simulations to ICARUS can simplify design or map dummy equipment that is not real process equipment.<br />
<br />
*It is good practice to include design factors for safety throughout the process. However, Aspen will map the equipment exactly as specified in HYSYS and therefore will not include an design factors in calculating the capital costs<br />
<br />
*Pressure vessels are costed exactly according to ASME Boiler and Pressure Vessel Code Section VIII Division 1. However, in some cases, this may an inadequate pressure vessel design. In these cases, the design should be manually entered.<br />
<br />
*Some processes require nonstandard components that HYSYS has no way of modeling correctly and for which ICARUS has no appropriate equipment category. Aspen has the capability to include non-standard equipment libraries which often can be obtained by equipment manufacturers. Adding these libraries allows use of the costing software for cost estimates.<br />
<br />
==Validity of Cost Estimates==<br />
<br />
One thing to keep in mind is that cost estimates are inherently associated with relatively high uncertainty. By leaving many aspects of the plant unspecified, the error grows dramatically. This should be kept in mind when utilizing cost estimates to perform economic analysis of the chemical process. A process that appears viable but has 50% error associated with capital costs, may quickly become undesirable as the project evolves. For this reason, it is essential that cost estimates include the most detailed design data possible.<br />
<br />
==Conclusions==<br />
While determining the capital cost of a chemical plant is difficult, it is an extremely vital aspect of determining of construction of a given plant is feasible given realistic financial constraints. For this reason, a number of tools have been developed to produce capital cost estimates at relatively early phases of plant construction including order of magnitude estimates, cost curve calculations, and more detailed costing of designed process equipment and other ancillary buildings and equipment.<br />
<br />
==References==<br />
<br />
Costdataonline.com. Richardson International Construction Factors Manual [Internet]. Pahrump: Cost Data On Line, Inc.; c2008- [cited 2015 Feb 26]. Available from: http://www.icoste.org/Book_Reviews/CFM-Info.pdf.<br />
<br />
Mecklenburgh JC. Plant Design and Economics for Chemical Engineers. New York: Halsted Press; 1985.<br />
<br />
Peters MS, Timmerhaus KD, West RE. Plant Design and Economics for Chemical Engineers. 5th ed. New York: McGraw-Hill; 2002.<br />
<br />
Towler G, Sinnott R. Capital Cost Estimating. In: Chemical Engineering Design: Principles, Practice and Economics of Plant and Process Design. 2nd ed. Boston: Elsevier; 2013. p. 307–354.</div>WFu