Design S1

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Project Justification

1,4 Butanediol is an important organic chemical used mainly as an intermediate for the production of Tetrahydrofuran (THF), Gamma-Butyrolactone (GBL) and Polybutylene Terephthalate (PBT). In 2011, the global market for BDO reached 1,725.0 kilo tons. Current projections suggest that worldwide BDO consumption will grow to 2550 kilo tons by 2017. The Technology Division of Evanston Chemicals has requested a preliminary design and economic evaluation on the feasibility of producing 1,4 Butanediol from a biomass derived succinic acid.

Technology Review

BDO can be produced by hydrogenation of succinic acid in the presence of a catalyst, at high temperatures and pressures (approximately 150-200°C and 20,000 kPa). However, to the best of the authors’ knowledge, there does not exist a means to produce only BDO. Instead, a mixture of THF, GBL, and BDO is produced. The choice of catalyst is the major factor in controlling yields and selectivity of a certain product. Current research suggests that bimetallic catalysts have a stronger activity and lead to higher conversion to BDO. TiO supported 2%Pd/X%Re catalyst and Carbon supported 2%Ru/4%Re catalyst both have been shown to be particularly effective. It is interesting to note that both catalysts contain Re. Research suggests that the synergy between Re and Pd/Ru leads to higher selectivity for BDO. Additionally, the chosen catalyst for this study was a 0.4% Fe, 1.9% Na, 2.66% Ag, 2.66% Pd, 10.0% Re on 1.5 mm carbon support. Taken from an ISP Investments patent from 2011, this catalyst has exceptionally high yields, upwards of 99.7% conversion of succinic acid, as well as high yields of BDO. In total, the percentage conversions of succinic acid to various products are: BDO: 85.51 wt % GBL: 2.04 wt % THF: 9.28 wt % Butanol: 2.87 wt % The lifetime of the catalyst is approximately 5 years. Furthermore, multiple reactor types have been discussed in the literature, including CSTRs and PFRs. However, the literature also shows that gas liquid induction reactors are the best for hydrogenation reactions in industry. They are preferred for their safety with respect to explosive gas phase reactants and complete utilization of the gas phase reactant; being especially useful for expensive inputs. These reactors use energy efficient impellers for gas induction and dispersion along with special recovery equipment to prevent waste of the gas phase, thus GLIR was chosen for this study.

Design Basis

The plant has been proposed to be built in Lake Providence, Louisiana. This is due primarily to its proximity to suppliers of succinic acid in the south east region, as well as close to distributors by way of the Mississippi River and the Gulf of Mexico. Additionally, the plant has been designed to have a capacity of approximately 50MM kg/yr of BDO production. This will require feeds of compressed hydrogen gas, and an aqueous feed of water and succinic acid, in a 50/50 wt % composition.

Technical Approach

The primary tool used to model the plant process was Aspen HYSYS. Because of the non-idealities involved in the NRTL HYSYS fluid package, an additional fluid package was created in Aspen. Additionally, the GLIR reactor chosen was modeled as a conversion reactor. Selectivity and yield of the reactor was modeled based off the specific data from the ISP Investments patent (discussed in Introduction: Technology Review). Although this provided very accurate modeling, the operating pressure and temperature were required to be held constant. This prevented any optimization within the reactor with respect to these operating conditions. The distillation columns were modeled as fractional distillation columns, with bottoms or distillate flow rate, and component recovery as active constraints. Additionally, Aspen Energy Analyzer was used to create the heat exchanger network, which can be seen in the Appendix. The economic analysis was performed in Aspen ICARUS.

Figure 1. Screenshot of Aspen HYSYS process simulation. As mentioned above, this plant will produce technical grade BDO and GBL, but the THF stream will be treated as waste. Originally, attempts were made to purify the THF stream in order to sell it; this process produces approximately 5M kg/yr, and at a selling point of ~$3.00/kg, this would generate an additional $15M/yr in revenue. However, the stream containing THF also contains water and n-butanol. As shown in Appendix XX, THF and water have a pressure dependent azeotrope; Figures XX1 and XX2 show that the azeotrope can be broken by pressurizing the stream after initial separation and atmospheric pressure. However, the presence of n-butanol makes separation of THF and water impossible; butanol and water have a tight azeotrope that cannot be broken simply by pressurizing the stream. Instead, absorption or ion exchange must be used to purify the stream. After extensive efforts were made to purify the THF, it was decided that the undertaking would not be profitable. In future iterations of this project, redoubled efforts would be made to force separation of THF to generate large revenues.

Process Flow Diagram & Flow Sheets

Below is listed a complete flow sheet, with data taken from Aspen HYSYS. Table 1 Stream Name Vapour Fraction Temp. (C) Pressure (kPa) Molar Flow (kgmole/h) Liquid Volume Flow (m3/h) Heat Flow (kJ/h) Feed 0 25 101 576 15.45 -2.13E+08 Feed_2 0 29 22408 576 15.45 -2.13E+08 Feed_3 0 164 22408 576 15.45 -2.05E+08 Reactor_Vap 1 165 22408 75 3.97 -5.17E+05 Reactor_Liq 0 165 22408 869 22 -2.53E+08 H2 Feed 1 25 22408 340 18.21 9.10E+04 H2 1 165 22408 376 20.14 1.62E+06 Bot 0 225 2200 200 9.19 -6.85E+07 8 0.08 165 22408 944 25.97 -2.54E+08 Vap_Rec 1 75 22383 72 3.86 6.86E+04 Liq 0 75 22383 872 22.11 -2.61E+08 9 0.08 75 22383 944 25.97 -2.61E+08 H2Rec_2 1 75 22383 36 1.93 3.43E+04 H3 1 30 22383 376 20.14 1.25E+05 H2Rec_1 1 75 22383 36 1.93 3.43E+04 Purge 1 75 22383 36 1.93 3.43E+04 Distillate 0 187 2000 671 12.88 -1.82E+08 Liqa 0 80 2200 872 22.11 -2.61E+08 Liqb 0 211 2200 872 22.11 -2.50E+08 Vapor dis 1 187 2000 1 0.04 -2.17E+05 Dis_2 0 103 100 4 0.16 -1.36E+06 Product 0 233 120 63 5.57 -2.87E+07 Bota 0.31 150 400 200 9.19 -6.85E+07 Botb 0.67 202 400 200 9.19 -6.43E+07 Dis_Vap_1 1 202 400 133 3.46 -3.38E+07 Dis_Bot_1 0 202 400 67 5.73 -3.05E+07 H20Rec_2 0 25 1100 133 3.45 -4.10E+07 H20Rec_3 0 28 22408 133 3.45 -4.09E+07 Recycle_3 0 164 22408 133 3.45 -3.93E+07 H20Rec_1 1 202 400 133 3.45 -3.37E+07 ReactorFeed 0 164 22408 708 18.9 -2.44E+08 Dis2Feed 0 202 400 67 5.73 -3.05E+07 Distillate_1 0.68 125 90 4 0.16 -1.24E+06 THF 0.19 103 120 671 12.88 -1.82E+08 Waste_1 0 -189 100 14 0.97 -3.98E+06 WaterWaste 0 105 120 657 11.91 -1.84E+08 Waste 0.19 103 120 671 12.88 -1.82E+08 VapDis_Out 0 30 1000 1 0.04 -2.66E+05 Water_4 0 94 80 3 0.05 -8.04E+05 GBL_1 0 199 90 1 0.11 -5.45E+05

A detailed sizing list of the relevant components is listed below. Of particular interest are the distillation columns, which have heights between 8 and 12 meters. Furthermore, the reactor has a length of 6.2 meters and a diameter of 2.1 meters. A more complete table of sizing and sizing methodologies is listed in the Appendix. Table 2 Type Sizing Spec Purpose Compressor (E-11) Ṿ=18.2m^3/hr Compress inlet hydrogen to reactor pressure Heat Exchanger 1 (E-9) A = 634.21m^2 Heat feed hydrogen to reactor temperature Pump 1 (E-1) Ṿ = 3.5m^3/hr Duty = 27.4kW Pump SUC feed to reactor pressure Pump 2 (E-21) Ṿ=15.5m^3/hr Duty=137 kW Pump Recycle to reactor pressure Jacketed GLIR (E-3) L = 6.2m D=2.1m t=230mm Gas Liquid Induction Reactor used to convert SUC to BDO, GBL, THF and side products Heat Exchanger 2 (E-2) A = 4600 m^2 Heat reactor effluent in order to drive better separation Vapor-Liquid Separator 1 (E-12) H=3.0m D=1.5m t=15mm Separates unreacted hydrogen from reactor effluent Heater 1 (E-16) A = 575.15 m^2 Heats Distillation Column inlet to Column conditions Distillation Column 1 (13) H=10m D=1.82m #Trays=10 Drive separation Vapor-Liquid Separator 2 (E-20) H=2.0m D=.61m t=6.5mm Drive separation Distillation Column 2 (E-18) H=8m D=.60m #Trays=16 Purify product grade BDO Distillation Column 3 (E-26) H=8m D=.90m #Trays=8 Purify waste water to expedite water removal Distillation Column 4 (E-23) H=12m D=.65m #Trays=12 Purify product grade GBL

Economic Analysis

Table 3 Manufacturing Cost Summary Capital Cost Operating Cost Compressor $ 513,300.00 SUC Raw Materials $ 111,139,488.18 Reactor $ 1,237,100.00 Catalyst $ 500,000.00 Distillation Columns $ 927,600.00 Water $ 158,018.83 Separators $ 507,200.00 Contingencies $ 2,600,000.00 Heat Exchangers $ 3,071,800.00 Utilities $ 9,500,000.00 Pumps $ 182,900.00 Hydrogen $ 2,706,102.05 Storage Tanks $ 658,300.00 S & OH $ 1,500,000.00 Flare $ 82,300.00 Engineering $ 2,960,000.00 Total $ 7,180,500.00 Total $ 131,063,609.05

Table 3 is a summary of the capital and operating costs required for the production of BDO and GBL. The expenditure on succinic acid clearly stands out; it makes up roughly 80% of the annual operating costs of this plant. The production plant consumes approximately 75.6M kg/yr of succinic acid, and at $1.47/kg, the costs add up quickly. Some other things to note from Table XX; the reactor is the most expensive singular piece of equipment; because it operates at exceedingly high pressure, it has extremely thick walls and is made of a high strength steel alloy. Additionally, the total cost of the heat exchangers is more than $3M; however, the production plant uses twelve heat exchangers, so the average individual cost comes out to a more reasonable $280,000. Lastly, this plant has a high annual utilities cost, mainly because of the 4 distillation columns and the large amount of cooling water necessary to keep the process running at the correct temperature. Below in Table XX+1 is a summary of the economic measures of return of this process. Table XX+1 Investment Summary Simple Pay-Back Period (yrs) 0.45 Avg Cash Flow ($MM/yr) 40.8 10yr ROI 264.13% 10yr NPV ($MM) 163.1 20yr NPV ($MM) 258.0 10yr IRR 128.37% 20yr IRR 128.50%

As shown clearly in Table 3, this plant has extremely encouraging economic potential, with an average yearly cash flow of $40.8M, a simple payback period of less than 6 months, and a breakeven point of approximately 9 months. These capital projections assume a 2 year construction time, 38% tax rate, 50% capacity in the first year of production, and a 7 year MACRS depreciation method. Revenue is split into two streams; main product (BDO) and by-product (GBL). The main product revenue is approximately $150M/yr, while the by-product revenue is $40M/yr, despite producing nearly 50 times less GBL than BDO. In future iterations of this design, it may be worthwhile to examine if it is more profitable to convert all of the succinic acid directly to GBL instead of producing the intermediary, BDO. If the economic returns seem astronomical, it is because they are; no economic investment actually produces these kinds of returns. The extremely favorable prediction could stem from several sources of error; supply and demand, bulk pricing, waste removal, and side reactions. The yearly supply of BDO is approximately 1M metric tons, meaning that this plant would produce 5% of the world’s supply. If the supply of BDO outpaced the demand, the selling price would fall, cutting into main product revenues. Additionally, the estimate of GBL selling price was based off a per kg price; if bulk pricing were used, the selling price of GBL would be between 40-50% lower than the price used. In addition to economic discontinuities, the problem of waste removal was never fully solved. Based on past estimates, waste removal could cost as must as $10M/year, further reducing the profitability of this plant. Lastly, the process of BDO synthesis was simplified for modelling purposes. In actual production, many side reactions would occur; driving down conversion and increasing separation costs.

A sensitivity analysis was performed on all important variables.

Figure 1: Sensitivity analysis Figure 1 shows that the profitability of the plant (based on the 10 year NPV), is most sensitive to the selling price of BDO and GLB, the cost of raw materials, and the construction time. It is least sensitive to changes in the utilities cost and fixed capital costs. It is clear from the figure that this undertaking is profitable in the face of any singular change. However, if several of these variables react unfavorably to the coming economic climate, the plant could potentially make very little money. For example, if the construction time doubles (which happens frequently) and the selling prices of BDO and GBL are driven down by market saturation, the plant could have a 10 year NPV falls to approximately 0.


To conclude, this project is highly feasible, and our company stands to make a very high profit from the construction of this plant. Based on a two year construction time, a 38% tax rate, and 50% production over the first year, our simple payback period would be 0.45 years, with yearly revenue of over $190M. We strongly recommend Evanston Chemical to move forward with this facility. However, there are key next steps to be considered. First, a more detailed reaction scheme for the reactor is required. This scheme must be able to handle variable operating pressure and temperature, as well as contain a more detailed and complete incorporation of side reactions. Namely, the reaction of GBL to BDO is a necessary consideration. Moreover, a fully life cycle analysis and environmental health and safety analysis will be required. The removal of waste will also need to be addressed. Additionally, profits may be increased by performing a plant wide optimization project. Also, the assumption of a 5 year catalyst lifetime will need to be addressed in more detail, as the high cost of this component will greatly affect the bottom line. Another factor which may affect the overall profitability is the flocculating prices of rare metals, which make up the catalyst. Also, the market price of these commodity chemicals will almost certainly be affected by a 10% influx of global supply. Finally, while we have planned on also selling GBL, looking into purifying the THF will also be profitable.


Appendix 1

Sizing and construction information.

Component Diameter (m) Height/L (m) Material of construction Wall thickness (mm) Addt'l notes Reactor 2.06 6.19 Stainless steel 410 231 Hemispherical head, 104 mm Separator 1 1.52 2.96 Stainless steel 410 153 Hemispherical head, 69 mm Separator 2 0.61 2 Carbon steel 6.51 Torispherical head, 7.66 mm Distillation 1 1.82 10 - - 10 trays, 1 m tray spacing Distillation 2 0.6 8 - - 16 trays, .5 m tray spacing Distillation 3 0.6 8 - - 16 trays, .5 m tray spacing

Details: Reactor: Reactor Design Values 3575 psi 192 C Flow 16.56 m3/h LHSV 0.8 h-1

mu 0.001684 U 0.001376 ds 1.50E-03 rho 840.1 Pressure drop 5.52E+03 Pa 0.8 Psi (across react)

Design (as pressure vessel) Material SS 410 S (psi) 1.81E+04 E 1 t_hoop 0.231231 m t_long 0.098037 m Head type Hemisph High P t_head 0.103965 m

Volume 20.7 m3

L:D 3 to 1 L= 6.190242 m D= 2.063414 m

Separator 1: Settling velocity Towler p. 769 rho_l 974.6 kg/m3 rho_v 15.24 kg/m3 u_t 0.555387831 m/s

V_flow 146.6 kg/h V_flow 0.040722222 kg/s V_flow 0.002672062 m3/s

L_flow 6.31E+04 kg/h L_flow 1.75E+01 kg/s L_flow 1.80E-02 m3/s

Volume 5.39E+00 m3 Allowing for 5 minutes of hold up >Due to high volume flow of liquid D_min 0.07826726 m D_design 1.524 m For appropriate height

H 2.956815432 m From p.770 labels

Design specs (as Pressure V) Design Values 3575 psi ~50 C Material SS 410 S (psi) 2.00E+04 E 1 t_hoop 1.53E-01 m t_long 6.58E-02 m Head type Hemisph High P t_head 6.93E-02 m

Separator 2: Settling velocity Towler p. 769 rho_l 833.8 kg/m3 rho_v 2.702 kg/m3 u_t 1.227669556 m/s

V_flow 3481 kg/h V_flow 0.966944444 kg/s V_flow 0.357862489 m3/s

L_flow 5.87E+03 kg/h L_flow 1.63E+00 kg/s L_flow 1.96E-03 m3/s

Volume 5.87E-01 m3 Allowing for 5 minutes of hold up >Due to high volume flow of liquid D_min 0.609217549 m D_design 0.61 m For appropriate height

H 2.008137095 m From p.770 labels

Design specs (as Pressure V) Design Values 64 psi ~50 C Material Carbon steel S (psi) 1.30E+04 E 1 t_hoop 1.51E-03 m plus 5 mm corrosion 6.51E-03 m t_long 7.50E-04 m Head type Tori Low P t_head 2.66E-03 m

Distillation 1: Column diameter p.853 towler l_t 1 m Tray spacing rho_l 785.8 kg/m3 rho_v 14.99 kg/m3 u_v hat 0.380057 V_max 14.9 kg/s D_c 1.824833 m (six inch increments) Trays 10 H 10 m

Distillation 2: Column diameter p.853 towler l_t 0.5 m Tray spacing rho_l 794.7 kg/m3 rho_v 14.99 kg/m3 u_v hat 0.328154 V_max 1.38 kg/s D_c 0.597661 m (six inch increments) Trays 16 H 8 m

Distillation 3: Column diameter p.853 towler l_t 1 m Tray spacing rho_l 913 kg/m3 rho_v 14.99 kg/m3 u_v hat 0.410219 V_max 3.536111 kg/s D_c 0.855676 m (six inch increments) Trays 8 H 8 m

Appendix 2

Catalyst Pricing Information: Liter of catalyst 9072.580645 L per 42.0417 days Time on stream 1009 hrs-experiment Assume lifetime of 5 years

Void fraction 0.4 Avg. rho_catalyst 4.52 g/cm3 Avg. rho_catalyst 4520 g/L Amount of cat 41008064.52 g Amount of cat 41008.06452 kg Price of cat 32.61962751 $/kg Price (lifetime) (w. margin) $ 2,006,501.68 $ per 5 years

Catalyst components Frac (%) Price Units rho (g/cm3) Price ($/kg) Fe 0.4 160 dol/ton 7.874 3.1496 0.176367 0.004 0.000705 Na 1.9 2.5 dol/kg 0.97 1.843 1.9 0.019 0.0361 Ag 2.66 25 dol/oz 10.49 27.9034 25 0.0266 0.665 Pd 2.66 1000 dol/kg 12.023 31.98118 1000 0.0266 26.6 Re 10 50 dol/kg 21.02 210.2 50 0.1 5 Support 82.38 0.3858 dol/kg 2.15 177.117 0.3858 0.8238 0.317822 4.521942 32.61963 $/kg g/cm3

Appendix 3

Equipment Costs Area Name Component Name Component Type Total Direct Cost Equipment Cost Main Area H2 Feed Compressor DGC RECIP MOTR 513,300.00 427,600.00 Main Area Reactor Vessel DHT JACKETED 1,237,100.00 961,000.00 Main Area Separator 1 DHT HORIZ DRUM 420,400.00 273,200.00 Main Area Separator 2 DHT HORIZ DRUM 86,800.00 8,600.00 Main Area Distillation 1 DTW TRAYED 313,000.00 97,500.00 Main Area Distillation 2 DTW TRAYED 193,700.00 41,300.00 Main Area H.E. 1 DHE TEMA EXCH 198,500.00 48,000.00 Main Area H.E.2 DHE TEMA EXCH 286,000.00 94,500.00 Main Area H.E.3 DHE TEMA EXCH 277,000.00 86,400.00 Main Area H.E.4 DHE TEMA EXCH 83,200.00 12,000.00 Main Area H.E.5 DHE TEMA EXCH 1,203,300.00 607,000.00 Main Area H.E.6 DHE TEMA EXCH 111,900.00 22,800.00 Main Area H.E.7 DHE TEMA EXCH 107,500.00 18,800.00 Main Area H.E.8 DHE TEMA EXCH 138,600.00 28,100.00 Main Area H.E.9 DHE TEMA EXCH 298,500.00 115,100.00 Main Area H.E.10 DHE TEMA EXCH 214,100.00 62,100.00 Main Area H.E.11 DHE TEMA EXCH 153,200.00 32,900.00 Main Area Pump 1 DCP ANSI 35,900.00 8,800.00 Main Area Pump 2 DCP ANSI 147,000.00 79,400.00 Main Area Purge Flare DFLRSELF SUPP 82,300.00 11,300.00 Main Area Distillation Column 3 DTW TRAYED 206,900.00 46,500.00 Main Area Product Storage DVT STORAGE 229,300.00 131,800.00 Main Area Waste Storage DVT STORAGE 221,100.00 125,200.00 Main Area GBL Storage DVT STORAGE 207,900.00 79,300.00 Main Area Distillation Column 4 DTW TRAYED 214,000.00 81,400.00 Total 7,180,500.00 3,500,600.00

Appendix 3

Economic Sumamry

Appendix 4

Energy Stream Summary

Appendix 5

HYSYS Simulation

Appendix 5b

Stream Summary Table

Appendix 7

Summary of HYSYS stream compositions.

Appendix 8

Economic analysis summary

Appendix 9

Assorted economic data


    Basic Engineering	767,300.00	6,959.00
    Detail Engineering	1,552,700.00	14,893.00
    Material Procurement	537,300.00	 
    Home Office	98,700.00	1,030.00
    Total Design, Eng, Procurement Cost	2,956,000.00	 

Figure A1: Summary of Engineering Costs OPERATING LABOR AND MAINTENANCE COSTS

    Operating Labor 	 	 
         Operators per Shift	 	3
         Unit Cost	Cost/Operator/H	20
         Total Operating Labor Cost	Cost/period	480,000.00
         Cost/8000 Hours	 	83,300.00
         Total Maintenance Cost	Cost/period	83,300.00
         Supervisors per Shift	 	1
         Unit Cost	Cost/Supervisor/H	35
         Total Supervision Cost	Cost/period	280,000.00

Figure A2: Summary of Labor and Maintenance Costs PROJECT RESULTS SUMMARY

    Total Project Capital Cost	Cost	17,094,689.70
    Total Raw Materials Cost	Cost/period	0
    Total Products Sales	Cost/period	0
    Total Operating Labor and Maintenance Cost	Cost/period	843,300.00
    Total Utilities Cost	Cost/period	346,035.64
    Total Operating Cost	Cost/period	1,945,064.49
    Operating Labor Cost	Cost/period	760,000.00
    Maintenance Cost	Cost/period	83,300.00
    Operating Charges	Cost/period	190,000.00
    Plant Overhead	Cost/period	421,650.00
    Subtotal Operating Cost	Cost/period	1,800,985.64
    G and A Cost	 	144,078.85

Figure A3: Overall Summary of Project Results Appendix 10. VLE Data

Figure A4: Water vs THF VLE at 100 kPa Figure A5: Water vs THF VLE at 2000 kPa

Figure A6: Water vs Butanol VLE at 100 kPa Figure A7: Water vs Butanol VLE at 2000 kPa

Appendix 11

Project capital summary. PROJECT CAPITAL SUMMARY Total Cost Design, Eng, Procurement Construction Material Construction Manhours Construction Manpower Construction Indirects

    Purchased Equipment	Cost	3,562,500.20	 	3,562,500.20	 		 
    Equipment Setting	Cost	48,033.10	 		1,591.00	48,033.10	 
    Piping	Cost	1,839,521.60	 	1,300,326.60	18,056.00	539,194.80	 
    Civil	Cost	352,120.60	 	183,510.60	7,104.00	168,610.10	 
    Steel	Cost	71,926.20	 	58,874.10	493	13,052.10	 
    Instrumentation	Cost	1,109,755.50	 	924,111.60	6,171.00	185,644.00	 
    Electrical	Cost	684,348.20	 	596,287.40	3,084.00	88,060.80	 
    Insulation	Cost	594,591.20	 	362,250.20	10,562.00	232,341.10	 
    Paint	Cost	123,535.10	 	34,906.60	4,037.00	88,628.50	 
    Other	Cost	5,357,000.50	2,956,000.20	719,800.10	 		1,681,200.10
    Subcontracts	Cost	0	 		 		 
    G and A Overheads	Cost	323,619.90	0	232,277.00	 	40,906.90	50,436.00
    Contract Fee	Cost	597,333.90	174,404.00	159,496.90	 	117,975.60	145,457.40
    Escalation	Cost	0	0	0	 	0	0
    Contingencies	Cost	2,639,571.20	563,472.70	1,464,181.20	 	274,040.50	337,876.80
    Total Project Cost	Cost	17,303,857.20	 		 		 
    Adjusted Total Project Cost	Cost	17,094,689.70	 	 	 	 	 

Appendix 12

Heat exchanger network.

Appendix 13

Reaction mechanism summary