Table 1. Milling Equipment Summary
The total cost of the milling equipment is summarized in Table 2.
A detailed discussion of the milling and pretreatment equipment and material streams is
available in the attachment Milling Material Streams.
Key assumptions made for material balances and sizing were found in laboratory papers
. in sizing the equipment, the multiple effect evaporators were assumed to be of equal size to
make calculations simpler. In reality, each evaporator would be smaller than the previous one.
Although the milling of sugarcane closely followed the standard process used today, the
sucrose extraction step and clarification step were researched and investigated for better options.
As shown in Table 1, the sucrose extraction step is completed by shredding the sugarcane and
thus breaking the cells in which sucrose is located, making it easier to extract. The clarification
step is important because it removes the impurities from the cane juice that would otherwise only
interfere with the product purity.
Conventionally, roller mills are the equipment of choice used to extract the sucrose from
sugarcane. However, sucrose extraction by diffusion has recently emerged in the industry. The
sugarcane goes through twelve stages where it is heated co-currently by imbibition water,
immediately raising the temperature of the cane to about 75 °C . This heating even opens the
cells that were not opened by the previous shredding process. Another advantage of this heating
is that sugar destroying bacteria cannot survive at this temperature. However, in a mill tandem,
the average temperature is between 30-35 °C, where these bacteria are still active . Extraction
by diffusion boasts a higher extraction percentage, lower operating and manufacturing costs, and
better mechanical reliability. The ratio of capital costs of a plant using a diffuser compared to
one using a mill is about 1:1.5 . One downside of diffusion is that the shredding process is
significant because the cane needs to be the right size for efficient diffusion to occur. Therefore,
a heavy-duty shredder is needed, which is initially more expensive and requires more power.
However, looking in the long run, the appeal of lower capital cost, lower operating and
maintenance cost, and high extraction percentages made diffusion the better alternative.
The clarification process is important because it removes the impurities from the cane
juice and also neutralizes its acidity. There were many options to consider: defecation
(Ca(OH)2), magnesia (MgO2), sulphitation (lime and SO2), carbonation (lime and CO2), and
phosphatation (lime and H3PO4) . Each differ in cost, the time needed to precipitate the
impurities, and product purity. With these factors in mind, calcium hydroxide, also known as
milk of lime, was chosen because its reaction time was the shortest and cost the least. The
comparisons are shown in Table 3.
Milling Possible Errors
There were a few sources of possible error in the milling process. When sizing the
diffuser, the main equipment piece in this section of the process, dimensions were extracted from
existing diffuser models. Using the flow rate of sugarcane needed to produce the desired amount
of ethanol, the size of the diffuser was estimated. Also, because the evaporators were assumed to
be the same size, ICARUS produced a larger cost than it should be (viewable in ICARUS
equipment list attachment). These possible errors would affect the overall cost of the design.
Bagasse Hydrolysis Results
Bagasse consists of the fibers that come out of the diffuser and in industry, it is generally
burned in a cogeneration system that produces electricity to be used in different parts of the
ethanol production process. However, recent processes have begun to process the bagasse in a
way that it can be added into the fermentation step to increase ethanol production. About 90% of
the bagasse produced is used to make ethanol while the remaining 10% is used to generate
electricity. This process pretreats the cleaned bagasse with 2 wt.% sulphuric acid  and steam
and through reactors, filters, distillation, and evaporation, a 25 wt.% glucose solution is produced. Sterilized sugarcane juice is diluted and added to the glucose solution to before
entering the fermenters. The addition of the bagasse sugars increases ethanol production. Table 4
summarizes the equipment chosen for this process.
- All pretreatment and hydrolysis equipment will be made of stainless steel because sulphuric
acid is being used.
The total cost of the bagasse hydrolysis equipment is summarized in Table 5 below.
A detailed discussion of the bagasse hydrolysis equipment and material streams is available in
the attachment Milling Material Streams.
Bagasse Hydrolysis Possible Errors
Possible sources of error in the Organosolv process stem from the fact that the process
has yet to be produced on an industrial scale. Therefore, all data has been interpolated from
laboratory data. This would potentially affect the final glucose yield in the process as well as the
overall cost of the design.
The fermentation operation was based on standard batch processes using the well
characterized yeast strain, Saccaromyces Cerevisiae. The glucose and sucrose solution from the
milling section is sent to two, 100,000 L storage tanks which will provide a buffer between the
continuous milling process and the batch fermentation process. As one of our 46 fermentation
vessels becomes available for filling, the sugar solution from the storage tanks is pumped out of
storage, mixed with the required yeast nutrients, and sent into the fermentation vessel. Yeast
culture is then pumped to the fermenter from propagation vessels which are the last stage in our
yeast recycle loop. The reaction will proceed within the fermenter for 27 hours which is based on
literature values. When the reaction is completed, the solution will pass to one of another two,
100,000 L storage tanks and then to a 41.7 cm diameter, continuous hydrocyclone. This
hydrocyclone will separate the yeast from the product mixture which will then be mixed in a 200
L mixing tank to a 50 wt % water slurry for pumping to one of 9 propagation fermenters. This
yeast separation step allows for recycle of our used yeast. Meanwhile the remaining product
solution is passed to the separations unit for purification. A summary of all required equipment is
given in Table 6.
Details pertaining to the design of the equipment listed in table 6 can be found in the attachment
entitled Fermentation Calculations.
Mode of Operation
Both continuous and batch processes were considered for our fermentation units. The
continuous process was based on immobilized cell bed reactors  though there were also
examples of fluidized bed reactors  which also performed the desired process. The batch
process, on the other hand, was very simple in that it only involved loading the specified amount
of material and reacting these components for a specified period of time.
Both modes of operation were compared in deciding our final process. A continuous process
such as an Immobilized cell bed offered a higher output rate for a given capital investment while
also allowing easy interfacing with the rest of the plant. The batch process, in comparison,
offered easy cleaning, flexible unit operation, and easier process control due to the simple reactor
When comparing the two modes, our team decided to pursue batch operation for our
fermentation units. This decision was largely motivated by the simplicity of the batch reactor and
the lack of industrial examples of continuous fermentation. Given the current economic climate,
it would be imprudent to pursue a risky venture such as continuous fermentation without
significant experience to draw upon so it was decided that batch fermentation was the best route
to take. In order to minimize costs in other areas of the plant, it was decided that though the fermentation units will be run as a batch operation, the units will be staggered so as to allow
continuous operation in the milling and separation processes. The fermentation vessels will be
preceded and followed by large storage tanks which will serve to buffer the small variations in
flow rate to or from the vessels. This will allow the overall plant to continuously produce
Choice of Bacteria
In deciding what type of organism to use, there were two main options. These two
options were Saccharomyces Cerevisiae and Z. Mobilis. Saccharomyces Cerevisiae is a strain of
yeast and is the most widely used organism in the beverage and fuel ethanol industry. This
organism is widely known and there is a wide range of literature detailing the optimization of its
alcohol production. Yeast produces alcohol in concentrations of approximately 10% and can use
a wide range of substrates as feed which simplifies the fermentation growth medium. This upper
limit for ethanol production is due to the toxicity of ethanol toward the yeast and is the major
disadvantage of this organism.
Z. Mobilis is a strain of bacteria which has become a major focus of research in recent
years. This recent focus has come about due to a higher ethanol tolerance when compared to the
traditional S. Cerevisiae strain of yeast. This higher ethanol tolerance allows more efficient use
of batch units as the process can produce more concentrated solutions of ethanol. The major
disadvantage of Z. Mobilis is a narrow range of substrates which may be digested for
fermentation. This disadvantage is likely a minor issue for sugarcane fermentation, however, due
to the simple sugars (glucose and fructose) which are produced from the milling process.
In the end, it was decided that S. Cerevisiae was the best organism for the process. Z. Mobilis is
largely a “laboratory-scale” strain as there are very few examples of its use in industrial
applications. Z. Mobilis is an option that should be under ongoing evaluation in the future,
however, due to the promise it shows as an ethanol producer.
The growth medium is a very important part of fermentation because without
supplementary nutrition, yeast metabolism and ethanol production is severely hampered. The
most important nutrition requirements for fermentation are a Nitrogen source to provide raw
material for protein synthesis and a source of Magnesium which allows for rapid yeast growth
. The original growth medium to be used in our process is shown in table 6 . This
medium, however, was adapted from laboratory practices and was found to be uneconomical on
an industrial scale. The laboratory medium was quoted at more than $20 million per year of
supplementary nutrients and this was determined to be unacceptable for our commercial process.
The most significant contribution to the price was due to the presence of 5 g/L of peptone in the
growth medium which was quoted at around five dollars per kilogram or $19.3 million per year.
There has been some work in optimizing growth media for industrial applications,
however, and Pereira et. Al.  have performed an in depth investigation of this problem. They
found that the growth medium presented in Table 8 allowed for optimal yeast nutrition and
performance in industrial applications for a minimal price. The savings in this instance are
largely a result of substituting corn steep liquor, a complex nitrogen source, for yeast extract and
peptone. Corn steep liquor was quoted at five cents per kilogram  and offers a significant
economic advantage over typical laboratory supplements. This is the growth medium that was
chosen for our fermentation units.
Due to the low volume present in the propagation units, we decided to use a laboratory scale
growth medium for our propagation step. The cost of material for this step is insignificant given
the scale of our overall process. This growth medium is presented in table 9 .
Yeast Separation Method
After the fermentation step, it is desirable to remove the yeast solids from the reactor
product before sending the stream to the distillation unit. This prevents the yeast from interfering
with any of the distillation calculations but also allows for the possibility of yeast recycle back to
the reactor vessel. For the separation, we considered two options. First we considered a disc-stack centrifuge. This type of centrifuge is widely used in the industry and offered an attractive,
high efficiency separation. This unit requires the use of centripetal force to separate the yeast
solids from the outlet stream and this was the main disadvantage with this option. Spinning forty
kilograms of solution per second was an undesirable situation and the high shear stresses inside a
disc-stack centrifuge may have caused damage to the yeast which would limit the efficiency of
the recycle stream.
The other option for this separation was the use of a hydrocyclone. A hydrocyclone is a
cone with the inlet at the top and outlets at the top and bottom for the liquid and solids
respectively. The separation is driven by the pressure drop in the system with the mechanism
being the vortex created in the interior of the unit. Our specific outlet composition appears to be
on the lower end of the range for which hydrocyclones apply according to Towler . It should
also be noted that our solids concentration will likely be greater than that quoted in the flow
sheets due to excess yeast culture growth in both the propagation and fermentation steps. Due to
this, hydrocyclones should do an adequate job of separating the yeast from the ethanol product.
Hydrocyclones offer a number of benefits including minimal moving parts (safety and
maintenance), low energy requirements compared to centrifugation, and cheap capital costs. For
these reasons, a hydrocyclone was chosen for our yeast separation step.
The option of recycling yeast was explored in our design process. The yeast would be
separated in the hydrocyclone, then mixed with water into a 50 wt % slurry. This slurry would
then be pumped to a fermentation vessel where the yeast would be regenerated to form the
culture for the main fermentation tanks. This recycle process was compared to buying new yeast
to pitch into the fermentation tanks instead of regenerating the used yeast. We found that fresh
yeast may be acquired for approximately $2/kg  and, with only 35 kg required per batch, this
was an interesting option for the plant. It was found that pursuing yeast recycle and propagation
was likely more economical than the alternative, however, due to relatively equal capital cost
requirements but advantageous operating costs. Fresh dried yeast was found to require the use of
a tank to soak the yeast in order to activate the culture for fermentation. This increased the
capital costs beyond those initially estimated. Buying yeast would also cost at least an additional
$730/day extra when comparing the cost of the yeast versus the foregone ethanol production due
to sugar consumption in propagation. Due to the advantageous economics, we have decided to
install a propagation step in our plant to regenerate used yeast.
Based on the decisions discussed in the previous sections, a set of assumptions was developed
for the fermentation section. Information for input prices and key assumptions for the
fermentation reaction are given in tables 10 and 11.
Fermentation Possible Errors
There are a number of aspects of the design of our fermentation process unit which
deserve further research and verification. The most significant source of possible error is the
assumptions for the actual reaction. Most significantly, the reaction time and final ethanol
concentration may prove different upon scale up from laboratory scale. The industrial scale yeast
growth medium shown in table 8 may also affect the reaction time and final concentration. The
cooling requirements for the reactor are also a likely source of error. The reaction profile was
assumed to be linear for simplicity and this is known to be incorrect. An actual fermentation
reaction will have a slow build up stage, an exponential growth stage, and a slow nutrition
limited stage. Therefore, during the exponential growth phase, the cooling requirements will
likely be much greater than estimated. All of these sources of error are centered on the actual
chemical reaction characteristics. Therefore, it is recommended that a small reactor be used to
characterize these key variables prior to the final build.
Research on industrial ethanol-water separations processes yielded an initial design for
the production of both hydrous and anhydrous ethanol. However, such industrial separation
processes also produce a number of byproducts, such as fusel oil, which ultimately reduce the
yield of hydrous ethanol . As the fundamental premise of this project is to produce as much
ethanol as possible from a given quantity of sugar cane feedstock, the proposed design differs
significantly from traditional ethanol separation processes. The separations for this method were
modeled using Aspen Hysys. The model is presented in Figure 2. A full output of the Hysys
simulation is available in the attachment Separations Workbook v2.
Figure 2. Simulation in Aspen HYSYS
It was deduced that the most effective means of achieving the desired ethanol product
purity was to design a separations process that consists of two parts: 1) the distillation of hydrous
ethanol and 2) the dehydration of hydrous ethanol by an azeotropic distillation process . The
azeotropic distillation design utilizes benzene as an entrainer . From this method, the desired
quality (99.9 mol% ethanol) and quantity (19,500 kg/h) of product is achieved. The
recommended design is presented in the Fermentation and Separations PFD and Figure 2. Table
12 summarizes the equipment utilized in the separation process.
From this design, not only are the ethanol production quality and quantity requirements
met, but byproduct carbon dioxide (CO2) produced from the fermentation process is successfully
separated and purified to 99.9 mol% and produced at a rate of 5200 kg/h. This process introduces
an additional $415 per hour cash flow based upon present carbon dioxide prices. This provides
not only an additional revenue stream for the plant but also a hedge against predicted new
regulations on carbon dioxide emissions.
The total cost of the separations and purification process is summarized in Table 13.
The sizing calculations for the separations vessels as well as one of the pumps, one of the
compressors, and one of the heat exchangers are available in the attachments BENZ STRIPPER,
DEHYDRATOR, DISTILLATION_SIZE, RECTIFIER, Phase_Splitters, P-800, C-2, and E-200.
One of the fundamental assumptions of the separations processes was the composition of
the feed from the fermentation units. After consulting a literature source  utilizing a similar
fermentation design to the one suggested in this report, and accounting for the ethanol production
of the yeast Saccharomyces cerevisiae, the inlet composition as given in Table 14 was
A second assumption was implementing the NRTL fluid package. From the Aspen HYSYS
manual, it was deduced that this fluid package would best represent the components in the
The final separations and purification design implements three rectifying columns, one
distillation column, four flash drums and a decanter, as summarized by the Fermentation and
Separations PFD and table 12. The flash drums serve a threefold process in the process: 1) to
eliminate volatile components from the product stream, 2) purify byproduct carbon dioxide and 3) purify recycle streams. Flash drums are preferred in separation processes as they retain
significantly lower capital, operation and maintenance costs, as evident in Table 13.
For the hydrous ethanol purification process, a rectifying and distillation column are
implemented in series. The rectifying column removes water and heavy components, while the
distillation column purifies the product stream to a composition of 91 wt% ethanol – a value
slightly under the ethanol-water azeotrope composition of 95.6 wt% ethanol . The choice to
utilize a rectifying column was made for both economic and practical reasons. As demonstrated
by the Fermentation and Separations PFD, the distillate vapor collected from the rectifying
column is sent directly to the distillation column. Maintaining the feed as a vapor ensures a more
effective separation of the ethanol from water and other undesired components. Furthermore,
such a design also reduces the energy demand of the distillation column reboiler.
The ethanol dehydration unit consists of two rectifying columns in series. The first
column of the dehydration unit is fed with two steams: 1) hydrous ethanol and 2) a recycle
stream containing benzene. The addition of benzene, which acts as an entrainer, leads to the
formation of a ternary azeotrope that possesses a boiling point significantly different from the
boiling point of ethanol. During the azeotropic distillation process, the ternary azeotrope would
distills to remove water. Then the binary azeotrope distills to remove any benzene, resulting in
99.9 mol% ethanol as the bottoms product. The distillate water, ethanol and benzene mixture is
fed to a decanter, where organic and aqueous phases are separated. The aqueous phase is
subsequently fed to a secondary rectifying column where wastewater is collected and benzene is
Despite its deleterious health effects and environmental effect, benzene was chosen as the
entraining agent. Although a number of alternatives act to form a ternary azeotrope with ethanol
and water, such as cyclohexane, it was determined that benzene proves to be the most cost
effective, as it requires the least amount of material to form the desired ternary azeotrope. It is
important to note that 0.33 kg/h of benzene exit the dehydration unit in the anhydrous ethanol
stream. Relative to the quantity of ethanol in the stream (19500kg/h), this quantity is minimal.
Quantitatively, benzene in the anhydrous product is 0.0015 vol%. This is significantly below the
Environmental Protection Agencies (EPA) maximum limit of 0.62 vol% stated in the Mobile
Source Air Toxics Rule .
Separations Possible Error
The likely sources of error in the separation process design pertain to the validity of the initial
assumptions. Specifically, it is possible that the composition of the fermentation feed stream 1)
retains different mole compositions or 2) addition (or fewer) components from those presented in
Table 14. Furthermore, it is presumed that the NRTL equation of state accurately describes the
vapor-liquid equilibrium (VLE) of all of the components in the separations process. From the
simulation output, it is evident that the NRTL fluid package does indeed capture the VLE
relationships accurately for simple binary mixtures of ethanol and water. However, for multicomponent
distillation systems (such as those in the hydrous purification system) it is likely that
some error arises naturally from the inability of NRTL to accurately predict VLE data for such
The plant utilities are displayed in the third page of the PDF, entitled Plant Utilities. A list
of the equipment shown is below in table 15.
Heat Exchanger Network
To economize utilities production, two heat exchanger networks were developed. One is
based on the milling and pretreatment area and the other on the separations area. Together, the
two networks save the plant the need for an additional 22,500 kW of heating and cooling (45,000
kW total), by using internal heat exchangers. These heat exchanger networks assumed a ΔTmin of
20℃ and can be viewed in the attachment Heat Exchanger Networks. The calculations are with
the related sections of the plant in the attachments Milling Material Streams, Separations
Workbook v2, and Utilities Streams.
Many sugarcane ethanol plants in the world utilize other materials besides the main sugar
cane feedstock for utility stream production . Our plant will separate up to 90% of the waste
bagasse material to be processed through hydrolysis and delignification process that will produce
glucose liquor. The glucose liquor will be fed to the fermentation process to increase ethanol
production. Produced lignin will be added to the remaining 10% of the bagasse material for
combustion in the steam boilers as can be seen in the Milling/Pre-treatment PFD and the Plant
Utilities PFD. These splits were based on research studies our team found on the utilization of
the bagassse not only as a fuel, but also as an additional source of ethanol. Research suggests that
these operational splits will allow for complete onsite electricity and steam generation for the
production facility in addition to increasing ethanol yields by up to 26% .
The size of the co-generation plant was based on several assumptions about the feed
streams and equipment. The plant will make use of the remaining 10% of the bagasse, mixed
with the lignin and some of the field trash, excess crop material that is often a waste stream for
the farmer, to generate steam. Table 16 contains values from literature and the material balances
regarding the boiler fuel .
Based on steam generation literature for bagasse combustion, the boilers are expect to run at an
efficiency of 85% . The cogeneration system will be backed up with #2 Fuel Oil for
emergencies, startup, and shutdown. The fuel oil has a lower heating value of 39.9 MJ/L. Current
designs include a single high pressure steam (HPS) stream utilized throughout the plant. The
HPS will be maintained at 500℃, 20 bar as super saturated steam. our boilers will generate 8.7
million kg of HPS per hour. Based on a 872,000 kg per hour throughput of water and steam, the
plant will require 10 field erected boilers. Calculations related to the co-generation steam
generation are available in the attachment Utility Streams.
According to EPA studies, the combustion of sugarcane bagasse will not be a significant
source of nitrogen oxides and sulfur oxides. It will, however, provide a significant source of
volatile organic compounds and particulate matter. The boiler is equipped with wet scrubbers to
maintain the boiler exhaust to within environmental regulations. For this matter, in absence of
finding environmental regulations from Costa Rica, we have followed the regulations of the US
EPA. The scrubbers will also provide relief when using the reserve #2 fuel oil .
Several assumptions were also considered in sizing the electrical generation portion of
the plant. After use in operation, the remaining steam will drive electrical turbines (at 98%
efficiency) to generate electricity. After the electrical turbines, the steam should contain just
enough energy to remain saturated steam vapor. Ultimately, the vapor will be condensed to a
liquid in a shell and tube heat exchanger to preheat water entering the boiler. To calculate the
energy that can be convert by the steam turbines, it was the energy per hour of the steam leaving
the boiler was calculated. Then the base energy (the energy of the water the steam will be after
the turbines), and the process steam energy requirements were subtracted from the energy per
hour value. Finally, it was assumed that the steam loop would undergo 50% losses of energy to
the environment throughout the system loop. The remaining energy was converted at 98%
efficiency to electricity . Calculations related to the co-generation electricity generation are
available in the attachment Utility Streams.
Additionally, it was estimated that the plant would be capable of selling electricity at a
price of $0.08 per kilowatt hour. This is based on an industrial and commercial average
electricity price in Costa Rica of $0.12 per kilowatt hour. Since the electricity is highly regulated
a conservative estimate was chosen .
The electrical generation portion of the plant is estimated at a rating of 163 MW. A
tighter ultimate build on the plant will restrict losses, increasing electrical production. This
estimated generation will surpass the expected plant requirement of 19 MW, allowing the plant
to sell up to 144 megawatts (MW) of electricity per hour. Based on the electricity price in Costa
Rica of 0.08 cents per kilowatt hour, the plant can generate an extra revenue stream of $84
million per year selling electricity . The capital cost of the cogeneration plant is estimated at
$143 million installed.
Water Treatment Plant
In order to provide clean water for plant processes, the plant will utilize a two stage ion
exchange water treatment process. The plant will be equipped with 40 units capable of cleaning
52 cubic meters of water per hour. The sizing of these facilities is based on the water recycle
ratios and process water needs described below. These treatment units will be used to clean
Tempisque River water for use as steam, cooling water, and washing water. The units will also
be used to clean wash water after use in the milling process prior to disposal in the Tempisque River. The total cost of the water treatment equipment is estimated at $4.5 million installed.
Calculations related to the water treatment plant are available in the attachment Utility Streams.
In-Process Water Recycling
While the design would optimally recycle all of the in-process heating and cooling water,
it is realistically infeasible due to fluctuations in production, fowling, and degradation.
Therefore, initial set points have set the cooling water at a recycle rate of 90% and the steam
boiled water at a recycle rate of 98% (this is due to already accounted for pipe losses). Washing
water will be recycled at a rate of 50%. By these ratios the plant will withdraw approximately
675 m3 per hour of water from the Tempisque River. This water will be approximately 18℃
when drawn from the river. After plant startup, studies will evaluate these recycle ratios,
optimizing them to save utility costs. Calculations related to the water recycling are available in
the attachment Utility Streams.
After any necessary cleaning, water planned for ejection into the Tempisque River will be
held in a cooling pond for a period of a week. In accordance with the US Clean Water Act and
Costa Rican environmental regulations, these streams will be too warm to be dumped into a river
containing any wildlife. To remedy this, the plant will use a cooling pond, modeled off the one
used by the Central Utility Plant at Northwestern University. Reject cooling water and steam at
the university is emitted into a man-made lagoon allowing it circulate and cool prior to rejoining
waters of Lake Michigan. Our plant will use a 150,000 m3 cooling pond, allowing the reject
water to cool to ambient air temperature before returning it to the river. This will offset damage
that water at the incorrect temperature could do to the river ecosystem . The total cost of the
pond is estimated at just under a million dollars installed. Calculations related to the pond are
available in the attachment Utility Streams.
The design team has developed the entire plant design thus far with safety as the primary
concern. Therefore an effort has been made to reduce the use of carcinogenic and other health
hazardous chemicals within the plant. The two exceptions are benzene used for the entrainment
distillation and the refrigerant used for the CO2 purification process. The refrigerant was required
for cooling to extreme temperatures to ensure CO2 separation from the remainder of the stream.
We chose to use the refrigerant R-500, for its properties and common uses . For safety
precautions, the refrigerant is in a self contained loop where it cools the process stream in a heat
exchanger, then is re-cooled in a condenser. The closed loop system helps to prevent venting the
refrigerant, which acts as a greenhouse gas.
Other safety features include adequate eye wash and safety shower locations throughout
the operations area for emergency use. Emergency lighting will be supplied on a separate battery
and circuitry from plant electricity, and all elevated areas will have appropriate railings and
multiple escapes in the event of an emergency. All employees will have full safety training
starting their first day of work onsite, and prior to any work they will complete an area specific
safety tutorial highlighting the area’s specific hazards and safety operations.
On the design side, all pressure vessels were design to the ASME BPV code and are
equipped with pressure relief devices. All valves and rotating equipment were designed with fail
positions to prevent dangerous situations such as pressure build up in power or control loss
situations. Since our facility will be operating with extensive rolling and crushing machinery in
the milling section, those machines will be designed with guards and two-handed required
operating controls to keep appendages away from the pinch points. Additionally, throughout the
indoor processes of the plant, mostly the milling and bagasse hydrolysis steps, the air handling
system will be designed to adequately remove dust and other particulate matter from the air to
prevent dust explosions. All electrical equipment will be rated for anti-spark areas and electrical
rooms will be pressure positive to prevent dust from travelling into them. Later in the process
when there is an extensive use of hot and cold streams, the design includes pipe insulation, even
if it is not needed to maintain process temperature, but to protect operators from accidental scalds
Finally, Personal Protective Equipment (PPE) will be a must. All operators will be
outfitted with hard hats, chemical resistant work gloves, chemical goggles, fire and chemical
retardant jump suits and steel toed boots. Those in the milling area will be outfitted with ear
protection as well, as the machines will be loud. Lastly, a smoke free site and a cell phone ban in
vehicles will help prevent industrial accidents.
Several measures have already been incorporated to increase the reliability of the plant.
First is a through storage system necessary to take the batch-wise fermentation and have it run
continuously with the remainder of the plant. The storage system offers two days of production
storage before and after the fermentation, and prior to the entrainment final dehydration step.
This will allow for the separations processes to continue running for two days in the event of a
fermentation issue. Additionally, in the fermenter sizing, the calculations allotted for each
fermenter to be taken out of service and cleaned for 4 hours at the end of each fermentation run.
This will increase product quality. Lastly, a large holding area was sized for incoming sugarcane
with the capacity to store up to a week supply at a time. The design specifications of the storage
vessels can be viewed in the attachment Storage Sizes.
Another reliability design feature is redundant pumps. Not only was each pump
purchased in double or triple, most of the pumps throughout the plant share a make, type, size,
and rating with several other pumps. This way, if one is broken and its reserve is broken, it may
be able to use a “brother” pump’s reserve.
Lastly, as the project cost was being completed, seeing that we were on target for
payback, we provided liberal corrosion allowances and materials selection on material to
improve the lifetime of the equipment.
Figure 3 is an in-depth process controls diagram for the refrigeration loop, used in the
CO2 processing equipment. Stream FD2-V is a vapor stream, containing CO2, water, ethanol, and
other byproducts. For the final CO2 bottling stream, pure CO2 is required, with impurities only in
the few parts-per-million. The stream is at 25℃ and 1.4 bar, with a flowrate of about 5300 kg per
hour. To achieve the desired purity, contents of FD2-V will be chilled to -78℃, then sent through
a flash separator, isolating the CO2. We are looking at the chilling process, FD2-VC (leaving the
diagram), takes the stream to the flash separator.
Figure 3. Process Control PFD
To chill the stream, we are using Refrigerant-500, which is in the Refrigerant and Refrig
Recycle streams at 100 wt% concentration. The R-500 is fed to the heat exchanger at -90℃,
cools the FD2-V, and leaves the heat exchanger at 10℃. A compressor unit re-cools the refrigerant to -90℃. Due to fluctuations in previous steps, it is possible that FD2-V will not
always be at the same flowrate and temperature. It is key to know the values of the incoming
stream to ensure enough cooling without too much cooling. If FD2-VC is too warm, the
separator will not work as well, and the CO2 will be filled with impurities. If FD2-VC is too
cold, then CO2 will sublime, no separation will occur, and no CO2 will be captured.
The first set of control systems controls the flowrate of the refrigerant based on the inlet
flowrate and temperature of FD2-V. A flow meter (FE1) and a temperature element (TE1) on
FD2-V are read by a flow transmitter (FT1) and a temperature transmitter (TT1), respectively,
sending a signal to a flow indicating controller (FIC1) and a computer, which using the
information, calculates the correct flow rate needed of refrigerant for the heat exchanger. The
computer and the FIC send a signal to a flow controlling automated screw valve (FV1) which
controls the flow of the refrigerant. The screw valve is fail closed, so that in a situation where
some power is lost, but the pump still has power, the valve will stop refrigerant flow, causing the
pressure relief (explained below) on the pump to circulate fluid rather than send it through the
The refrigerant cycles in a loop, which is driven by a centrifugal pump. If the valve is
sufficiently closed, or some other back up occurs, increasing the pressure in the pipe, a relief
valve (V-5) will be triggered. This relief valve will signal an alarm, and circulate the refrigerant
around the pump, not causing excess pressure on the valve or the compressor. The alarm will
notify personnel to rectify the situation.
An additional instrument is on the refrigerant leaving the condenser. Since the computer
calculating the flow of the refrigerant is assuming the refrigerant is within certain temperature
constraints, there is a temperature element (TE3), a temperature transmitter (TT3) and a
temperature indicating controller (TIC3) in the refrigerant line. This sensor is equipped with both
high and low alarms to indicate the operators of a failed compressor.
If the refrigerant is returning colder than expected, the work of the compressor can be
adjusted. This is controlled by temperature element (TE2), a temperature transmitter (TT2) and a
temperature indicating controller (TIC2) which adjusts a variable speed drive on the drive shaft
of the compressor. A cold refrigerant recycle will slow the drive as less compression is needed.
A warm recycle will increase drive speed for more compression. Finally there is a check valve
(V-2) on FD2-V ensuring it does not back feed the previous flash separator. There are also two
check valves in the refrigeration loop; one to prevent back feed into the compressor (V-3) and
one to prevent the compressor from back feeding the pump (V-4).
Economic analysis of the proposed plant design, on a twenty year basis, yields a net
present value (NPV) of $240 million, a rate of return of 19.7% and a simple pay-back period of
4.10 years. For traditional chemical plants, this rate of return would be questionably large.
However, the proposed plant produces large quantities of excess electricity, as a result of the
cogeneration design, sold to local power plants for a substantial profit. Thus, the rate of return in
this specific instance is not unreasonable. Moreover, the simple pay-back period is less than the
period desired by GICC executive management. The gross margin percentage after plant startup
(excluding electricity production) is estimated as 70% – indicating strong financial plant
performance relative to other chemical plants, which typically operate in the range of 40-50%.
Government requirements to blend traditional gasoline with ethanol have resulted in a dramatic
increase in the global ethanol market in the past two decades. Market analysis indicates that this trend is likely to continue into the coming decades. In addition to the obvious financial potential
of this project, the design implements technologies to reduce its environmental impact.
Specifically, carbon dioxide capture from the fermentation tanks as well as a cogeneration plant.
The proposed plant design is therefore strongly recommended for future development to
diversify GICC product portfolio.
The following recommendations offer methods to further enhance profitability of the
proposed plant. From our present calculations, whereby 10% of the produced bagasse is diverted
to the boiler system to produce steam (in addition to lignin and field trash), 163 MW of
electricity is generated. This is sufficient to not only power the entire plant operation, but yields
144 MW of excess electricity that is sold to local power production companies for substantial
profits. It is recommended that the fraction of bagasse that is sent to the boilers be optimized
based upon economic conditions. For instance, if the demand for ethanol (either domestically or
internationally) is large, more bagasse should be sent to the fermentation tanks to yield a higher
ethanol production, thereby relying solely on lignin and field trash to generate steam.
Given the seasonal nature of the sugar cane crop, it is suggested to explore the option of
preserving sugar cane so as to enable the ethanol production plant to operate throughout the year.
Furthermore, in the rare event that sugar cane crops do not achieve their predicted yields,
pursuing alternative sources of glucose, such as corn, to supplement sugar cane is recommended.
For the current design, waste removal costs are in excess of $6 million – a significant fraction of
the total plant operating cost. To reduce this cost, it is recommended that any opportunities to sell
byproduct pentose liquor are thoroughly explored.
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Appendix A: Attachments
SCEP Design Basis vFP
Cost of Production
Heat Exchanger Networks