KET050 Feasibility Studies on Industrial Plants Dept of Chemical Engineering, Lund Institute of Technology A feasibility study on conversion of an ethanol plant to a butanol plant Presented to StatoilHydro Oslo, Norway May 16, 2008 Principal investigators: Emil Larsson, Mark Max-Hansen, Anton Pålsson, Ragnar Studeny Tutors: Hans Eklund and Børre T. Børresen StatoilHydro Guido Zacchi and Hans T. Karlsson Department of Chemical Engineering
Transcript
KET050 Feasibility Studies on Industrial Plants Dept of Chemical Engineering, Lund Institute of Technology
A feasibility study on conversion of an
ethanol plant to a butanol plant
Presented to StatoilHydro
Oslo, Norway
May 16, 2008
Principal investigators: Emil Larsson, Mark Max-Hansen, Anton Pålsson, Ragnar Studeny
Tutors:
Hans Eklund and Børre T. Børresen
StatoilHydro
Guido Zacchi and Hans T. Karlsson
Department of Chemical Engineering
Page 1 of 81
Summary The interest for alternative fuels is increasing in the world, mainly because of the increasing oil demand
and the negative impact of fossil fuels on the environment. According to EU directives from 2003 5.75%
of all transportation fuel must be represented by biofuels by the year of 2010. Today ethanol is the most
common biofuel to replace gasoline and is produced from different types of biomass such as sugar,
starch and cellulosic material. Cellulosic raw material is of big interest because it is widespread and is
not included in the human food chain. Today the biggest producers of ethanol are Brazil and USA.
An alternative to ethanol is butanol which can be produced from the same raw materials as ethanol. The
main advantage with butanol is that it has higher energy density than ethanol. The energy content for
butanol is 28.5 MJ/l, compared to 23.5 MJ/l of ethanol. Another advantage is that butanol does not have
any corrosive properties which make it easier to distribute with existing piping infrastructure.
A preliminary study has been done to investigate the possibilities to convert an ethanol plant to produce
butanol. The ethanol plant uses 181 000 tons of dry spruce as raw material per year and the annual
production of ethanol is 50 000 m3. The butanol plant is based on the same amount of raw material, in
this way no changes are needed in the expensive pretreatment.
The butanol process is designed to use as much of the existing equipment available from the ethanol
plant as possible. Basically the only change is different microorganisms in the fermentation setup.
Research regarding other setups are not investigated but mentioned.
The ethanol process uses bakers yeast to ferment the sugar to ethanol. In the ABE process Clostridium
cultures are used instead. These bacteria yield not only butanol, but also acetone and ethanol and
therefore the name ABE. The main problem with these bacteria is their intolerance towards the
produced solvents, the total ABE concentration cannot exceed 20gL-1. This leads to heavily diluted broth
in the fermentation, which results in more energy demanding recovery. The technical calculations were
performed in Aspen Plus. Because of the great amount of water in the process both the existing
distillation and evaporation steps need additional process equipment, such as distillation columns and a
whole new parallel evaporation line. This is also true for the fermentation step where additional tanks
are needed.
A sensitivity analysis was performed in regard to the bacteria properties. This is achieved by stating
seven different cases where the bacteria properties are altered in different ways. Case I is based on
experimental data found in the literature, and the subsequent cases had improved selectivity towards
butanol, higher sugar conversion and higher solvent tolerance.
Case I: In this case the bacteria tolerate up to 20 gl-1 solvents and the conversion of the released sugars
is 60 %.
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The economical evaluation includes capital investment costs, payback time and operation costs for the
different cases. The investment cost for the ethanol plant is assumed to be paid off. The most important
parameter that affects the economical profit was identified as the tolerance towards ABE. The only
cases that turned out to be profitable were cases V, VI and VII. The investment cost for case I was
estimated to 41.51 million US $. The lowest investment cost was for case VI and was estimated to 29.73
million US $. The payback time for V, VI and VII, which are the only cases that are profitable, were 4.7,
3.3 and 3.5 years respectively. The payback time is based on an interest of 10 % and a deprecation time
of 15 years. The butanol production cost for these three cases were estimated to 0.92, 0.87 and 0.92 $/l
respectively.
The largest part of the investment cost is the new evaporation line and it is also the most power
demanding step in the process. A potential alternative to evaporating the stillage is using biogas-
fermentation, an alternative method using anaerobic microbes to ferment the compounds in the stillage
into biogas, however this has not been researched, so experimental data and research needs to be
made on the subject.
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Table of Contents Summary ....................................................................................................................................................... 1
1.2 Aim and Method ................................................................................................................................. 6
1.3 Ethanol and butanol ............................................................................................................................ 7
2. The Ethanol Process .................................................................................................................................. 9
2.1 Raw material ....................................................................................................................................... 9
2.9 Evaporation and Drying .................................................................................................................... 13
2.10 The existing ethanol plant ............................................................................................................... 14
3. The Butanol Process ................................................................................................................................ 17
4.4 Pay-back time .................................................................................................................................... 31
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5. Alternative methods and Research. ........................................................................................................ 32
Cellulose and hemicelluloses must be hydrolyzed to monomeric sugars before they can be
fermented. There are three ways of doing that: concentrated acid, dilute acid or with enzymes. The acid
attacks the bindings randomly and are therefore not selective to the bindings between the sugar
molecules. Because of this the formed sugars can be broken down to so called sugar degradation
products, these products are mainly furans and acids, and are not wanted in the process, as they may
inhibit the fermentation. Hydrolysis with concentrated acid and with diluted acid works the same way
with the difference that the diluted acid method needs higher pressure and temperature.[3,9]
2.4 Enzymatic hydrolysis To break the cellulose and hemicelluloses to monomeric sugars, three kinds of enzymes are used, endo-glucanases (EG),cellobiohydrolases (CBH) and β-glucosidase, these enzymes are called cellulases. Simplified EG attacks the amorphous parts of the cellulose and cuts of long fragment from the chain, which creates a lot of chain ends. The CBH enzymes are then attached to these chain ends and break the chain down to cellobiose which is two glucose units. Finally β-glucosidase breaks down the cellobiose into simple sugar molecules as shown in figure 2. The enzymes have maximum activity at a temperature between 50-60 0C and pH 4,5 – 6,0. Compared to acid hydrolysis, this method has a better selectivity. A disadvantage with this method is that the cellulases stop working when the concentration of sugars is high. This problem is solved if the hydrolysis and fermentation step is run at the same time, using so called simultaneous scarification and fermentation (SSF). The enzymatic method is also a very slow process compared to the acid method. The enzymatic method takes several days and the acid method needs only minutes.
Figure 2. EG attacks the cellulose, CBH then divide polysaccharides into shorter chains. Finally b-glucosidase cuts it into monosacharides.
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2.5 Fermentation
The fermentation can be run in separate stages or simultaneous with the hydrolysis. If the fermentation is separate from the hydrolysis, the solid materials mostly consisting of lignin is separated after the hydrolysis and the rest goes to the fermentation process. The fermentation is where sugars are transformed by the yeast Saccharomyces Cerevisiae into ethanol. C6H12O6 2 C2H5OH + 2CO2
2.6 Simultaneous hydrolysis and fermentation
In SSF the sugar is transformed to ethanol immediately after it is produced and therefore does not inhibit the enzymes. This method leads to a higher yield and productivity compared to the method with separate hydrolysis and fermentation steps. If the fermentation and hydrolysis are to be simultaneous, the hydrolysis must be enzymatic. A disadvantage with this method is that the enzyme cannot be reused, and that the process conditions must be chosen so that they fit both enzymes and the yeast.
2.7 Distillation
After the fermentation the ethanol is separated from the byproducts using multiple stage distillation,
consisting of two strippers and a rectifier column.
The broth is fed into the top of the first column. The distillate contains concentrated ethanol. In the
bottom of the column the byproducts, also known as stillage or syrup is obtained, and parts of this
stream can be used to dilute the stream going in to the fermentation. In this way the use of freshwater
can be minimized. The rest of the stillage goes to the evaporation stage.
The ethanol and water vapor from the top of the first column is not free from byproducts, this stream
also includes substances such as methanol, propanol and aldehydes. The separation from these
substances take place in the second column called rectification column. The bottom fraction (liquid)
goes with the stillage to the evaporation stage. The main goal in the second column is to obtain as pure
ethanol as possible but because of an azeotrope between water and ethanol the maximum
concentration of ethanol that can be obtained is 95,6 %w/w.[9]
2.8 Dewatering Because of the azeotrope in the distillation further concentration is needed. In this process a molecular
sieve is used to remove water. The molecular sieve consists of synthetic zeolites, and under pressure
these zeolites adsorb small molecules with dipoles. Water is such a kind of molecule and the ethanol
passes through the sieve. The lower limit of water in the outgoing ethanol is 20 ppm. The sieve is
saturated after 5-10 minutes and needs to be regenerated. To regenerate the sieve the pressure is
lowered, hot ethanol vapor is passed through in opposite direction and the vapor brings along the water
molecules. The ethanol and water is then recirculated into the distillation column. To make the process
continuous more than one molecular sieve is needed.[3,9]
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2.9 Evaporation and Drying
There are two kinds of byproducts from the distillation, solids and dissolved byproducts. The solid part consists of mostly lignin and can be used to make pellets or be burned to produce energy for the process. The dissolved byproducts are burned to produce energy for the process. To do this the stillage needs to be concentrated. The stillage goes to a filter press where the liquid is separated from the solids. The filter cake contains non fermentable lignin residues which have a high heat of combustion and therefore is suitable to use as fuel. The liquid from the filter press is concentrated in a multiple stage evaporator to form syrup with approximately 50 % dry substance. The filter cake is dried in a steam dryer. The remaining solid material could either be burned as it is to produce energy in the process or used to make pellets which can be sold as a solid fuel. [3,9]
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2.10 The existing ethanol plant This chapter gives a detailed description over the existing ethanol plant. A schematic figure of the
existing ethanol plant is shown in figure 3.
Figure 3 Schematic figure of the existing ethanol process
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The ethanol plant which is to be converted into producing butanol is described as follows. The process is
dimensioned to produce 50 000 m3 of pure ethanol per year. In the process lignin is taken out and can be
sold as a fuel in form of pellets, the amount of lignin from the process is 61 000 ton per year.[3]
First there is a mill which crushes the trees into smaller parts. The mill is a high speed hammer mill and
has a capacity of 22600 kg dry substance per hour.
Next sulfuric acid (2 weight %) is added to start the cleaving. After that the pressure is increased to 20 bar
using steam. The acid and the high pressure split the chips into smaller particles. After that the pressure is
dropped in a flash chamber to atmosphere pressure.
The fine particles are diluted with water before the fermentation. When diluted the temperature drops
from approximately 95 degrees to 37 degrees Celsius.
In the fermentation tanks yeast and enzymes are added to the mixture. The fermentation step consists of 9
stirred tanks. The tanks have a volume of 1000 m3 each and the effective volume is 800m
3.
The stream from the fermentation is heated from 37 degrees to approximate 118 degrees before entering
the distillation process. In table 3 the heat exchanger areas are shown.
Table 3 Heat exchanger areas.
Heat exchanger HE 6 HE7 HE 8 HE9
Area (m2) 55 55 140 109
The distillation process consists of two strippers and one rectifier. The trays are called donut trays are
used to avoid the strippers to clog. Table 4 shows the dimensions for the strippers and the rectifier.
Table 4 Dimensions for strippers and rectifier.
Stripper 1 Stripper 2 Rectifier
Height (m) 14.6 14.6 20.1
Diameter (m) 1.4 1.4 2.3
Number of trays 25 25 35
Rebolier area (m2) 598 687 839
Condenser area (m2) - - 278
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The stillage from the two strippers is sent to a filter press where the solid material is separated from the
liquid. The liquid is the sent to the evaporation step. The filter cake is then sent to a drier and formed to
pellets. The evaporation step consists of 5 effects. In the first effect 4 bar live steam is used as heating
medium and all the following effects are driven by secondary steam from the previous effect. The sizes
for the five effects are shown in table 5.
Table 5 Dimensions for the evaporation
Effect Area (m2)
1 438
2 240
3 271
4 341
5 1288
In table 6 a summary of the important streams in the process are shown.
Table 6 Flow rate of important streams.
The energy demand for the ethanol process is presented in table 7.
Table 7 Energy demand for the ethanol process
Process Pretreatment Distillation Evaporators Dryer Total
Energy demand (MW) 6.3 18 11 0.9 36.2
Mass frac (% of total flow)
Flow Total flow (kg/h) Ethanol Temperature (°C)
To distillation 131 700 3.7 37
To filter press 121 000 0.02 120
To evaporation 107 600 0.02 120
Stream 1 4938 100 35
From evaporation 8 400 0 61
From filter press 13 400 0 97
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3. The Butanol Process
3.1 Cases
Seven different cases have been studied and simulated. The inhibition and bad yield of the bacteria restrict the process and therefore seven cases where the
bacteria are assumed to have been improved in different ways were simulated in Aspen Plus.
In the first four cases the bacteria are assumed to tolerate a concentration of 20 gl-1 produced solvent
which corresponds to the maximum found in the literature for cellulose based processes. In case III and
IV the sugar conversion is improved by 20 %. In case V , VI and VII the bacteria is assumed to tolerate a
20% higher concentration of solvents. Table 8 presents the important parameters in the different cases.
Case I: This is the original case and based on information found[10] and referenced in appendix A.
Case II: The same as case I except that the selectivity towards butanol is assumed to have been
improved so that produced solvents are 98 % butanol, 1 % ethanol and 1 % acetone.
Case III: The same as case I, except that the conversion of sugar to solvents is assumed to be improved
by 20 %.
Case IV: The same as case III except that the selectivity towards butanol is assumed to be 98 %. Ethanol
and acetone are assumed to have selectivity of 1 % each.
Case V: The same as case I except that the bacteria is assumed to tolerate a 20 % higher concentration
before it gets inhibited and stops produce solvents.
Case VI: The same as case V but assuming the same selectivity improvement as case II and IV.
Case VII: The same as case IV except the tolerance is assumed to be the same as case V and VI.
Table 8. Comparison between cases. The ethanol production in the existing plant was 6.25 m3/h
3.4 Energy demand The process steps that need steam are divided into four posts: pretreatment, distillation, evaporation and drying. The theoretical energy demands for the process were calculated with Aspen Plus, a description of how the simulations were performed is shown in appendix B. There are two kinds of live steam in the pretreatment, 4 bar steam to heat the incoming wood and 20 bar steam to use in the high pressure dilute acid hydrolysis pretreatment. In the distillation, 4 bar live steam is used to preheat the streams and to heat the boilers in the high pressure strippers. The rest of the columns use steam from the strippers so the energy used in these are not included here. Steam with the pressure of 4 bar is used to heat the first two evaporators. In the pretreatment after the acid step there is a flash that gives 4 bar steam that can be used to heat the evaporators. The steam from the pretreatment is not enough to heat the evaporators and live steam is used to cover the energy shortage. The total power demand in MW for cases I-VII is presented in table 17.
Mass frac (% of total flow)
Flow Total flow (kg/h) Butanol Ethanol Acetone Temperature (°C)
To distillation 227 140 2.2 0.028 0.035 37.0
To filter press 222 070 0 0 0 113.3
To evaporation 207 690 0 0 0 113.3
Stream 1 155 2 39 49 45.5
Stream 2 4 916 100 0 0 90.4
From evaporation 6 260 0 0 0 62.7
From filter press 15 710 0 0 0 113.3
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Table 17 Total theoretical power demand in MW
Process Pretreatment Distillation Evaporators Dryer Total
Case I 18.6 23.6 19.4 - 61.6
Case II 18.6 23.1 19.5 - 61.2
Case III 18.6 28.0 25.0 - 71.5
Case IV 18.6 25.2 29.7 - 73.5
Case V 18.6 19.9 14.7 - 53.2
Case VI 18.6 18.4 13.1 0.35 50.5
Case VII 18.6 18.2 17.9 - 54,7
The steam is generated in a boiler where the syrup from the evaporation is used as fuel. If the energy from the syrup is not enough the lignin from the filter press can also be used in the burner. The burner has been assumed to have an efficiency of 90 %. Because of the high energy demand all of the lignin and the syrup have to be used as fuel in cases I-V and VII. In case VI some of the lignin has to be burned, and this has to be optimized so that the most possible amount of lignin can be used to produce pellets. Table 18 presents the actual energy need and the energy produced in the boiler. This process uses substantially more energy in comparison with the ethanol process, why the existing boiler might be under dimensioned, which leads to additional investments. This is however not considered in the present study.
Table 18 Realistic energy demand and production.
Process Power from boiler (MW)
Power demand for live steam generation (MW)
Surplus (MW)
Demand (MW)
Case I 59.1 68.4 - 9.30
Case II 58.7 68.0 - 9.27
Case III 52.8 79.5 - 26.7
Case IV 52.6 81.6 - 29.0
Case V 59.4 59.1 0.23 -
Case VI 56.1 56.1 0.05* -
Case VII 52.5 60.8 - 8.3
* Some of the filter cake is used to produce pellets, with a total effective combustion power of 2.82 MW.
The energy demand per m3 produced butanol is shown in table 19.
Table 19 The energy demand per m3 produced butanol
The depreciation of capital cost is estimated using the annuity method. The depreciation of capital cost
together with the operation costs will represent the total annual production costs. To calculate a
production cost for the butanol the total annual production costs are divided with the annual volume
produced butanol. The base investments were assumed to have a depreciation of 15 years with a real
interest rate of 10%. The annuity factor fA =0.13 was used. See Appendix E.
The production cost for butanol is presented in table 31.
Table 31 Calculated butanol production cost in US$/L
Case I Case II Case III Case IV Case V Case VI Case VII
Butanol production cost 1.13 1.14 1.24 1.23 0.92 0.89 0.92
The price of butanol is set to be able to compete with ethanol, in cases I-IV the production costs per liter
is higher than the competitive price. In case V and VI the production costs is lower than the market price
and a profit can be obtained.
4.4 Pay-back time
To calculate the number of years needed to pay back the investment the pay-back method is used. The
results are presented in table 32. Detailed calculations are shown in appendix E.
Table 32 Number of years needed to pay back the investment
Case I Case II Case III Case IV Case V Case VI Case VII
Pay-back year - - - - 4.7 3.3 3.5
Only three cases have an opportunity to pay back the investment.
The three cases with a payback time within 15 years are V, VI and VII. In these cases the bacteria is
assumed to be more resistant towards ABE. The most important parameter to make the process
economically feasible is the resistance towards the produced ABE. The selectivity towards butanol does
not affect the economy as much as expected.
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5. Alternative methods and Research. To rectify some of the problems associated with butanol fermentation, such as low solvent tolerance, and low selectivity towards butanol, and low conversion rates, research is being done in ways to extract the solvents from the fermentation broth, to maintain a low concentration of solvents, and thereby increase conversion rates. Furthermore a lot of research is being done on the microorganisms, and other bacteria, such as E. Coli are genetically modified to produce butanol.
5.1 Bacteria Acetone, butanol and ethanol are produced by solventogenic Clostridia. The most commonly used microorganism is Clostridium acetobutylicum. However the relatively new developed Clostridium beijerinckii BA 101 is more attractive because it produces higher concentrations butanol. C.acetobutylicum produces butanol concentrations in the range of 10-16 gL-1 whereas C.beijerinckii produces butanol in the range of 18-25 gL-1.[13] The concentration of ABE in the fermentation is limited due to the toxicity for the culture. The highest possible concentration of ABE is 20 gL-1.[14] In acid hydrolysis of agricultural biomass compounds such as furfural , hydroxyl methyl furfural (HMF), ferulic-, acetic, glucuronic, r-coumaric acids are produced. These compounds are often associated with inhibitory effects on microorganisms. r-coumaric- and ferulic acids decreases the ABE production and bacteria growth significantly. Furfural and HMF are not inhibitory to C. Beijerinckii BA 101; rather they have stimulatory effect on the growth of the microorganisms and ABE production. Acetic acids are not considered to have any inhibitory effects [10]. The robustness of clostridium is considered stable enough for industrial fermentations. Compared to ethanol producing organisms, clostridiums are able to produce solvents from both hexoses and pentoses. In figure 7 the ABE production from different types of sugar is presented. [10]
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Figure 7 Production of ABE from individual sugars (55g/l) using C.beijerinckii Ba 101
[10].
Another type of bacteria that is reported to be able to produce butanol from biomass is E-coli. By leveraging the native metabolic networks in E-coli and alter its intracellular chemistry scientists at UCLA claims the bacteria produces butanol from glucose. The produced butanol includes: isobutanol, 1-butanol, 2-methyl-1-butanol and 3-methyl-1-butanol. In particular the research team achieved high yield, high-specificity production of isobutanol from glucose[15]. This may be an interesting application in the future. However it still remains on the research stage. No data regarding these bacteria is yet available and therefore no calculations performed in the subject.
5.2 Extraction by Adsorption
The extraction by adsorption is based on the use of a solid phase adsorbent [16], to adsorb the desired product. Silicalite has been found to be very selective towards the adsorption of butanol [17]. This results in a stream of concentrated butanol, with up to 10% wt/wt of water and acetone and ethanol, which is good because it is above the binary water-butanol azeotrope. [7] There are a few different methods to achieve this, such as using two adsorption towers, see figure 9. By loading on one of them and regenerating the other at the same time, continuous operation can be achieved. The adsorbed species are desorbed by heat treatment, with a temperature ramp up to 200°C, which is far below the degradation temperature of the silicalite, which is 1000 °C. This method produces a stream consisting of over 90% butanol, which needs to be distilled to reach 100%, but the energy amount that is saved, when compared to distilling it with a feedback to break the azeotrope is substantial.
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This process however requires a substantial amount of silicalite, due to the adsorption capacity of silicalite being 97g / kg, with only ethanol and butanol, or 48g / kg when used in an ABE-system. This would lead to approximately two columns of 50 ton each, under the assumption that the loading and desorption times would be 2.30 minutes.
Figure 9. Bioreactor and Adsorption columns for continuous adsorption/desorption.
One problem with this method is that it is uncertain how the microorganisms will react to the silicalite, and that SSF results in solids in the fermentation broth. A way to work around this is to use an ultra-filtration membrane, to make sure that the cells, and other solids, stay in the fermentation loop, see figure 10.
Figure 10. Bioreactor with UF membrane and two adsorption columns. This method keeps the cells and solids in the bioreactor, and uses simultaneous adsorption/desorption
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Due to the lack of data, such as adsorption and desorption isotherms and other useable data, no mass or
energy balances could be performed that would require both isotherms, and the rates at which
adsorption and desorption occurs. The data that is available is based on a series of experiments, in
which a solution of acetone and butanol is kept with the Silicalite for 24 hours. The regeneration was
done by heating the silicalite to 300 °C, which desorbs the butanol. This kind of data is irrelevant for
process simulations. Adsorption and desorption times are highly relevant, since they in turn decide the
size of the adsorption columns. According to Quereshi[16] , the energy required is 8150 kJ/kg butanol, in
a adsorption system, whereas the current setup with conventional distillation has been calculated to use
73250kJ/kg . However the calculations used in article [16] are not available, and have not been verified.
In our opinion there is still a lot of development required before such a system will be available at an
industrial scale.
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5.3 Extractive fermentation
A reasonable approach to increase productivity and effectively recover the products may be the
utilization of extractive fermentation system, which could remove the inhibitory components from the
broth by in situ liquid-liquid extraction.
This method is based on a two phase fermentation. One water phase where the fermentation take place
and one organic phase. Figure 11 shows a typical flow sheet for extractive fermentation using oleyl
alcohol as solvent. The produced solvents accumulate in the organic phase. The organic phase from the
fermenter is sent first to the preheated vacuum vessel with a cooled stripper overhead for acetone
recovery. Later the bottom stream enters the second flash vessel and the overhead product, butanol is
collected. The extractants (oleyl alcolhol) remains in the bottom, and is recirculated to the fermenter for
reuse. The disadvantage of this system is that a large amount of expensive extractants are required for
the fermentations, which leads to a high cost of product recovery and regeneration of the
extractants.[18]
Figure 11. Fermenter and Flash stages for extraction based solvents recovery.
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6. Conclusions
The aim of this study was to investigate the theoretical possibilities for converting an existing cellulosic
Ethanol plant to produce butanol from the same feedstock. The pretreatment consisting of mills and
hydrolysis steps were retained and unaltered.
The bacteria used instead of yeast are of the Clostridium family. The bacteria convert sugar into a
mixture of acetone, butanol and ethanol, why this method is called the ABE-process.
The Achilles heel of this process is the fermentation step. The Bacteria has a maximum tolerance of
solvent up to 20gL-1 resulting in heavily diluted broth and expensive recovery. Other problems is bad
sugar conversion (only 62% of the hydrolisys sugars are utilized) and bad selectivity towards butanol.
The existing plant requires a lot of changes, and most of the equipment in the butanol process is new:
After the pretreatment step, the factory requires heavy modifications, starting with an increase from 9
fermentation tanks to 23 in case I, or to 19 tanks in case VI.
In the distillation step, there will be 4 strippers instead of 2, and two additional distillation columns and
two decanters will be needed in the cases I, III and V , and for case II, IV, VI and VII one new distillation
column and one decanter was needed.
The filter press can still be used, since it's the same amount of dry materials being used, but after the
filter press, the evaporator capacity has to be higher, which is achieved by installing a new five stage
line, and running them parallel.
Most of the changes are a result of the increased amount of water present in the system.
These changes lead to investment costs from approximately 37 million US $ for the cheapest (case VI) to
50 million US $ for the most expensive (case III).
As a sensitivity analysis the bacteria were improved in different ways. Seven different cases were
studied, see chapter 3.1, the first case, which is the basic case, or case I, has a total solvent
concentration of 20gl-1, and a 6:10:4 ratio of ABE.
For case I the energy demand was higher than the obtained energy from the boiler when slurry and solid
material was burned. The energy shortage of 9.3 MW was solved by buying pellets which were burned in
the boiler. This results in case I not being profitable. Case II which has a higher selectivity towards
butanol was not profitable either, since case II has the same problem as case I and the profit from a
higher butanol production was not enough to finance the pellets acquisition.
Cases III and IV have higher yields which leads to more products for sale than in cases I and II. A higher
yield however leads to more water in the system and therefore a higher energy demand. This in turn
leads to more pellets has to be bought and burned in the boiler to satisfy the energy demand.
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Cases V and VI which have a higher tolerance for the bacteria leads to less water needed in the
fermentation step and therefore a lower energy demand compared with the other cases. For case V the
energy demand is 59.1 MW and for case VI 56.1 MW. The energy obtained from burned slurry and solid
materials were for the two cases 59.4 respectively 56.1 MW. To satisfy the energy demand in case VI not
all solid material was needed in the boiler. Surplus solid material was dried and sold as pellets
5.1kton/year.
The annual income from sold products in case V and VI are 45.6 million US$ respectively 47.2 million
US$, and the annual operating cost are 33.25 million US$ respectively 33.24 million US$ . The production
price for butanol in case V and VI was calculated to 0.92 respectively 0.89 US$ per liter. The lower
energy demand for these processes leads to a profitable result. The depreciation times for these two
cases are 4.7 respective 3.3 years, this is based on a lifetime for the plant of 15 years.
Case VII, the case with most improvements, 20% higher yield and concentration tolerance and a 98%
selectivity towards butanol. This case has an energy demand of 60.8 MW and only 52.5 MW is obtained
from the burned slurry and solid material. The annual income from butanol sales is 53.9 million US$ and
the annual operating cost is 39.15 million US$. This gives a butanol production price of 0.92 US$ per
liter. The depreciation time for case VII is 3.5 years based on a lifetime for the plant of 15 years.
Although case VII is assumed to have the best bacteria it is not the most profitable one. Compared to
case VI which is the most profitable one the improvement in sugar yield leads to more water in the
fermentation step and less slurry to burn. This leads to a higher energy demand and less produced
energy in the boiler. The increase in butanol production does not cover the costs for the increase in
energy demand.
The largest part of the investment cost is the new evaporation line and it is also the most power
demanding step in the process. A potential alternative to evaporating the stillage is using biogas-
fermentation, an alternative method using anaerobic microbes to ferment the compounds in the stillage
into biogas, however this has not been researched, so experimental data and research needs to be
made on the subject.
As long as no major improvements are made to the bacteria, or new strains are found, according to the
data found and calculated, this process does not seem economically feasible. An increase of solvent
tolerance is the most important characteristic that needs to be attained in order to gain process
economy.
If broth adsorption was readily available perhaps this could change things, since lower energy
requirements would make the solids a revenue generating byproduct, instead of being burned to make
energy for the process, and according to Quereshi[H] substantially decreased energy demand, however
as the data this research was based on wasn’t available, those results and conclusions have not been
verified.
Page 39 of 81
7. References [1] Horn,Uwe Phd student LTH [2]www.SEKAB.com [3] Projekt höglandsetanol slutrapport-förstudie [4] Hans T.Karlsson Projekteringsmetodik 1992 [5]. Mörtstedt, Data & Diagram, 7th ed. 2005 [6] http://www.skogforsk.se/templates/sf_NewsPage____3011.aspx?sm=1&cpi=2169&ci=77 [7]. Kirk-Othmer _ Encyclopedia of Chemical Technology 5th ed. Vol.4 (355-360) [8]. Michelle Bryner, Chemical week, “DuPont and BP to build advanced biofuel facilities in UK” [9] Ingvarsson.M Jaeger.H Nilsson.H Swensson.E and Wendt.A Integrering av processer för production av drivmedelsetanol från stärkelse och cellolosa 2007 [10]. Ezeji.T, Qureshi.N Blaschek.H Butanol production from agricultural residues: Impact of degradation products on Clostridium beijerinckii growth and butanol fermentation, Journal of Biotechnology and Bioengineering, 2007, Vol: 97:6 pages: 1460-1469 [11] www.icispricing.com [12] Gunnarsson E.. Projektledare. Projekt Höglandsetanol slutrapport – förstudie. ÅF-Process AB. 2006-10-17 [13] Applied Microbiology Biotechnology (1998) 49: 639±648 P. Dürre New insights and novel developments in clostridial acetone/butanol/isopropanol fermentation [14] renewable energy 22 (2001) 557-564 N.Qureshi, H:P: Blaschek Recovery of butanol from fermentation broth by gas stripping [15]University of California, Los Angeles (2008, January 7). Efficient Biofuel Made From Genetically Modified E. Coli Bacteria. [16]. N. Qureshi et al, Energy-efficient recovery of butanol from model solutions and fermentation broth by adsorption. Bioprocess Biosyst Eng (2005) 27: 215_ 222 [17]. Milestone NB, Bibby DM (1981) Concentration of alcohols by adsorption on silicalite. Journal of Chemical Technology and Biotechnology 31:732_ 736 [18]. Shi.Z, Zhang.C Chen.J and Mao.Z Performance evaluation of acetone—butanol continous flash extractive fermentation process. Bioprocess and Biosystems Engineering ,2005 Vol 27:3 Pages: 175-183 [19] Hans T.Karlsson ProjekteringsHandboken 2007 [20] http://www.lr.dk/landbrugsinfo/informationsserier/dlbr_advice/nyhetsbrevlrfnov.pdf [21] http://www.skogssverige.se/skog/skogen/swe/lathund.cfm [22] www.kemirakemi.se [23] Wingren A.. Ethanol from Softwood – Techno-Economic Evaluation for Development pf the Enzymatic Process. Department of Chemical Engeneering. Lund University. Sweden. 2005 [24] http://www.okq8.se/servlet/ContentViewerServlet?contentUrl=cycube://internal/document/100663&nodeId=com.cycube.navigation.node.NodeId@100291 [25] www.vattenfall.se
The wood chips from the hammer mill are mixed with sulphuric acid and sent to a high pressure reactor
which is feed with 20 bar steam. The following reactions are taking place
Glucan + Water Glucose
Galactan + Water Galactose
Mannon + Water Mannose
Xylan + Water Xylose
Arabinan + Water Arabinose
The amount of sugars that are released in the high pressure reactor are shown in table 1
Table 1 The amount of sugars that are released in precent from the total amount of ingoing hemicelluloses and cellulose by the hydrolysis.
Glucose Mannose Galactose Xylose Arabinose
18 58.8 57.3 42 22.4
After the high pressure reactor the solution is fed to the fermentor tank and the cellulose enzymes
starts to attack the remaining cellulose molecules. The amount of the remaining sugar that is released
by the enzyme is shown in table 2
Table 2 The amount of sugar that are released in the enzymatic hydrolysis in precent of that available in the solid from the pretreatment.
Glucose Mannose Galactose Xylose Arabinose
91 80 80 80 80
When the numbers are summarized from the acid and enzymatic hydrolysis the total yield of sugar that
is released from the ingoing biomass is calculated and shown in table 3
Table 3 The total amount of sugars that are released in the high pressure reactor and the enzymatic hydrolysis in precent of the theoretical based on the composition of the raw material.
Glucose Mannose Galactos Xylose Arabinose
92.6 91.8 91.5 88.4 84.5
Page 41 of 81
Fermentation stoichiometric
The released sugars are fermented by the bacteria and transformed to butanol, ethanol, acetone and
some sugar is used by the bacteria to grow.
Very simplified the sugars are transformed by the following reactions.
Hexose sugars in the solution are glucose. galactose and mannose
Hexose sugars to butanol
C6H12O6 C4H9OH + H2O + 2CO2
Hexose sugars to ethanol
C6H12O6 2 C2H5OH + 2CO2
Hexose sugars to acetone
C6H12O6 2 CH3COCH3
The chemical equation for hexose sugar to acetone is not in balance so the theoretic amount of sugar
that is needed is based on the carbon balance.
Pentose sugars in the solution are xylose and arabinose
Pentose sugars to butanol
C5H10O5 C4H9OH+CO2+H2O
Pentose sugars to ethanol
3 C5H10O5 5 C2H5OH + 5 CO2
Pentose sugars to acetone
2 C5H10O5 3 CH3COCH3 +CO2+H2O
The chemical equation for pentose sugar to acetone and pentose sugar to butanol is not in balance so
the theoretic amount of sugar that is needed is based on the carbon balance.
The theoretic amount of one specific solvent that is possible to obtaion from 1 g hexose sugar is
calculated with equation 1.
g/g (1)
Where n = the number of solvent molecule obtaion from 1 molecule sugar.
Page 42 of 81
The same equation is used for butanol, ethanol and acetone from hexose sugar.
For pentose sugar almost the same equation is used with the difference that the molecule weight for
pentose sugar is used instead of hexose sugar. See equation 2.
g/g (2)
The result from equation 1 and 2 is shown in table 4.
The practical yields of ABE was obtained from a lab study[10] preformed in a 1 liter reactor with a sugar
concentration of 55g/l. After 60h the solution contained butanol, ethanol and acetone according to the
result shown in table 5.
Table 5 The amount of butanol, ethanol and acetone after 60h
55 g sugar Glucose Mannose Galactose Xylose Arabinose
Butanol (g) 13.5 12 8 13 13.5
Ethanol (g) 0.5 1 0.5 0.7 1
Acetone (g) 4 2 2 4 3
From the calculated theoretical yield it is possible to calculate how much sugar that has been consumed
to produce the amount of solvent measured from the experiment.
Equation 3 was used to calculate the theoretical amount of sugar that has been consume, see table 6.
(3)
Table 6 Theoretical amount of sugar consumed
Glucose Mannose Galactose Xylose Arabinose
Butanol (g) 32.93 29.27 19.51 20.31 21.09
Ethanol (g) 0.98 1.96 0.98 1.37 1.96
Acetone (g) 6.25 3.13 3.13 6.45 4.84
Sum 40.16 34.36 23.62 26.52 27.89
Page 43 of 81
From this it is possible to calculate how much of the 55 g sugar that has been transformed to butnol,
ethanol and acetone from each sugar type. Using equation 4 and the result is shown in table 7.
(4)
Table 7 The amount of sugar, in percent, which has been transformed to solvent from ingoing sugar
Glucose Mannose Galactose Xylose Arabinose
Butanol 60 53 35 37 38
Ethanol 2 4 2 2 4
Acetone 11 6 6 13 9
Total ABE 73 62 43 51 51
Page 44 of 81
Appendix B Simulation in Aspen Plus
Aspen Plus was used to calculate the energy demands and dimensions for the distillation- and
evaporation stages. The dimensions calculated with Aspen were used when calculating the additional
apparatus costs for the process.
Distillation in Aspen Plus
The distillation setup is different for the cases. For case I, III and V the setup consists of four strippers,
three rectification columns and two decanters. (see figure 5, chapter 3.3)
The new equipment are two strippers with 16 trays each, two rectifiers with 16 trays and two decanters.
The remaining cases uses the setup showed in figure 6 chapter 3.3. This setup includes four strippers,
two rectifiers and one decanter. The new equipment includes two strippers with 16 trays, one rectifier
with 16 trays and one decanter.
Because of solid substance in the feed open trays so called donut trays will be used. However in the
calculations valve trays with a 50% Murphree degree of efficiency where used instead. This results in a
more realistic description of the donut trays which are considered to be less effective than ordinary
valve trays. In the rectification columns a Murphree degree of efficiency of 75 % were used[9].
The model used describing the equilibrium between steam and liquid is an NRTL activity coefficient
model.
The task for the distillation plant is to separate the solvents from the other substances. The main goal is
to achieve pure butanol in the bottom of one of the rectifiers but in the cases where acetone is
produced it is also desirable to get the acetone as pure as possible to sell it as by product.
All the calculated dimensions are based on the condition to obtain 99.99 % pure butanol and in the
cases where acetone is produced, 99.99 % pure acetone. The dimensions are presented in tables 13-16
in appendix C.
Page 45 of 81
Evaporation in Aspen Plus
The bottom streams from the four strippers are sent to a filter press where 95 % of the solid material is
separated. The solid material however still contains water and the dry substance is 50 % . The stream
passing through the filter still contains a lot of dissolved dry material. The dry material can be used as
fuel to generate steam so it is desirable to remove some water and then burn it.
The evaporation line available from the ethanol process is not big enough to remove the desirable
amount of water and therefore an additional line is built. The flow going to the evaporation plant is split
into two different streams. The stream entering the existing line has about the same size as it had in the
ethanol plant which results in the same amount of live steam needed in that line.
The new line consists of five effects. The first effect has a pressure of 3 bar and the fifth a pressure of 0.2
bar. The pressures of the other effects are 1.7, 0.9 and 0.45 bar.
Live steam is used as heating medium in the first effect and the next effects are driven by secondary
steam from the previous effect. No consideration to the boiling point elevation was taken in the
simulation. In Aspen Plus each evaporator is described as a heat exchanger with a flash vessel.
The heat exchangers are assumed to have the following heat transfer coefficients (k): 1.97, 1.5, 1.5, 1.5
and 0.34 kWm-2C-1.[U]
The dry substance is 60 w/w % out of from evaporation plant.
Page 46 of 81
Appendix C Process calculations
Hydrolysis/Fermentation
No changes in the pre-treatment occur, as the same amount of raw materials will be used, and the same
pre-treatment equipment will be used. The first changes takes place in the hydrolysis/fermentation
step. As a result of the mixture being heavily diluted, more tanks are needed to compensate for the
volume increase. The numbers of extra tanks needed to maintain retention time is calculated. The
retention time is = 60 hours plus 12 hours for emptying, washing and filling. The flow rate, q, has been
calculated to 246 m3/h in case I. The inflows vary in the different cases and therefore the required
numbers of tanks vary. Each tank will have a working volume of 800 m3 each and a total volume of
1000m3 each. Table 1 shows the number of tanks needed in the different cases.
31772924672 mqV (1)
2316,22800
17729
tankV
Vn
(2)
Table 1 Total number of tanks needed in the different cases
Tanks needed additional tanks
Case I 23 14
Case II 23 14
Case III 27 18
Case IV 27 18
Case V 19 10
Case VI 19 10
Case VII 21 12
Steam dryer
The filter cake from the filter press is sent to a steam dryer, in order to remove water so that the filter cake can be pressed into pellets. The only case returning a surplus of energy is Case VI, why the calculations are based on this case. The filter cake is preheated to 140oC before entering the dryer, which give the dryer a pressure of 4 bar. The steam that is used in the dryer is therefore 4 bar which is overheated to 170oC with 20 bar steam. In table 3 the data for the steam and filter cake is shown. Table 4 presents the results from the calculations.
Page 47 of 81
Figure 1 Schematic picture of steam dryer
The amount of steam required in the pre heater is calculated with equation 3. Data for the dryer is shown in table 2, Where msteam is the mass flow rate of live steam used in the preheater, and mf is the mass flow rate of the filter cake.
TCm pF
vapsteam ΔHm (3)
Table 2 Data needed for steam dryer calculations
Data Case VI Units
Cp 4.21 kJ/kg°C
ΔHvap 2134 kJ/kg°C
Tin 113 °C
Tout 140 °C
Fm 1257 kg/h
The amount of steam that is required in the dryer is calculated with equations 4, 5, 6, 7 combined into
equation 8.
Mass balance:
ΔS = Sout - Sin (4)
Filter cake
12 w/w % water
Filter cake
50 w/w % water
Dryer Preheater
Steam Sin
(superheated)
Steam
Sin+ΔS
Page 48 of 81
Energy balance:
0ΔHΔSCpΔT SΔE steamsteamin (5)
saturatedoverheated TTΔT (6)
saturatedoverheated
saturatedoverheatedsteam
TT
hhCp (7)
steam
steam
inCpΔT
ΔHΔSS (8)
Table 3 Data for steam and filter cake
Data Case VI
Hoverheated (kJ/kg) 2798.96
Hsatured (kJ/kg) 2738.66
ΔHvap (kJ/kg) 2133.9
Tsaturated (°C) 140
Toverheated (°C) 170
Table 4 The steam requirement in case VI
Process ΔS [kg/h]
msteam [kg/h]
SDryer [kg/h]
Case VI 530.8 66.9 18784
Page 49 of 81
Steam recycling Overheated 4 bars steam is fed to the dryer and saturated when it comes in contact with the wet filter cake before leaving the dryer. The amount of steam that is generated from the material is condensed out and can be used for heating elsewhere in the process. The rest of the steam is overheated to 170oC with 20 bars steam in a counter current heat- exchanger. The power that is required in the heat exchanger is calculated from equation 9.
ΔTCSQ Poverheated (9)
And then the amount of 20 bars steam that is required is calculated using equation 10.
vap,20bar
20barΔH
Qm (10)
Data and results for equation 9 and 10 are shown in table 5.
Table 5 Calculated data and results.
Data Case VI
ΔT ( Co ) 25
Cp (kJ/kg) 2.01
ΔHvap20bar (kJ/kg) 1980.4
m20 bar (kg/h) 476.6
Calculation of heat exchanger areas for preheating of broth and distillation
The heat exchangers from the ethanol plant can be used as they are but the butanol plant needs more
heat exchangers due to the larger flow rates and more complex distillation setup. Five new heat
exchangers need to be bought and therefore the areas of those are calculated. The exchangers needed
to be bought are HE 1-5 in the flow sheet in chapter 3 figure 5 and 6.
All of the exchangers are counter current and the heat transfer coefficient is estimated to be 1 kW/ °C m2 for the heating and 0.9 kW/ °C m2 for the condensation.
Page 50 of 81
Thot in
Thot out
Tcoldt out Tcoldt in
All the calculations are made with equation 11.
LTAkQ ∆⋅⋅= (11)
where
( ) ( )( )( )
−−
−−−=∆
coldinhotout
cldouthotin
coldinhotoutcoldouthotinL
TT
TT
TTTTT
ln
Figure 1 Principle of ΔTL calculations
Exchangers 1 and 2 are the same in case I-VI and are calculated with Aspen Plus and are shown in table 6. Table 6 Exchanger area
Exchanger Area (m2) HE 1 75 HE 2 80
Heat exchanger 3-5 pre heating in distillation Heat exchangers 3 and 4 are used to pre heat the broth before the high pressure stripper and to pre heat the stream before the last distillation column where the pure butanol is obtained, respectively. Both these exchangers use 4 bar live steam as heating medium. Exchanger number 5 is used to condense the top product from the column where pure butanol is obtained. Water with a temperature of 10 °C is used as cooling medium. In the calculation it is assumed that the flow rate of this stream is high enough so that the temperature of the cooling medium does not change. The calculations for exchanger 3 and 4 are the same as above. For exchanger 5 the temperature difference is the same trough the whole exchanger, due to this the equation becomes as follows: (see equation1 2)
( )coldhot TTAkQ −⋅⋅= (12)
Table 7-13 shows all relevant data from Aspen Plus and for all the cases.
Page 51 of 81
Table 7-13 shows all relevant data from Aspen Plus and for all the cases. Table7 Case I heat exchanger 3-5
Case I
Heat Exchanger Heat duty (MW) T in (ºC) T out (ºC) T Cooling/Heating (ºC) k (kw/ ºC m2) Area (m2)
HE 3 5.3 90 118 144 1 139
HE 4 1.4 68 79 144 1 21
HE 5 0.8 67 67 10 0.9 21
Table 8 Case II heat exchanger 3-5
Case II
Heat Exchanger Heat duty (MW) T in (ºC) T out (ºC) T Cooling/Heating (ºC) k (kw/ ºC m2) Area (m2)
HE 3 5.1 94 121 144 1 148
HE 4 1.8 67 79 144 1 26
HE 5 1.0 67 67 10 0.9 20
Table 9 Case III heat exchanger 3-5
Case III
Heat Exchanger Heat duty (MW) T in (ºC) T out (ºC) T Cooling/Heating (ºC) k (kw/ ºC m2) Area (m2)
HE 3 6.3 90 118 144 1 165
HE 4 1.7 68 79 144 1 25
HE 5 1.0 67 67 10 0.9 19
Table 10 Case IV heat exchanger 3-5
Case IV
Heat Exchanger Heat duty (MW) T in (ºC) T out (ºC) T Cooling/Heating (ºC) k (kw/ ºC m2) Area (m2)
HE 3 5.5 93 121 144 1 169
HE 4 2.2 67 79 144 1 31
HE 5 1.2 85 85 10 0.9 18
Table 11 Case V heat exchanger 3-5
Case V
Heat Exchanger Heat duty (MW) T in (ºC) T out (ºC) T Cooling/Heating (ºC) k (kw/ ºC m2) Area (m2)
HE 3 4.4 90 118 144 1 115
HE 4 1.4 68 80 144 1 21
HE 5 0.8 67 67 10 0.9 16
Page 52 of 81
Table 12 Case VI heat exchanger 3-5
Case VI
Heat Exchanger Heat duty (MW) T in (ºC) T out (ºC) T Cooling/Heating (ºC) k(kw/ ºC m2) Area (m2)
HE 3 3.9 93 121 144 1 114
HE 4 1.8 67 80 144 1 26
HE 5 1.0 67 67 10 0.9 20
Table 33 Case VII heat exchanger 3-5
Case VII
Heat Exchanger Heat duty (MW) T in (ºC) T out (ºC) T Cooling/Heating (ºC) k(kw/ ºC m2) Area (m2)
HE 3 4.2 93 121 144 1 122
HE 4 2.2 67 80 144 1 31
HE 5 1.2 85 85 10 0.9 18
Strippers and rectifiers
The distillation setup requires that two strippers and two distillation columns have to be purchased. In order to calculate the cost for these columns the heights and diameters for the columns and the areas fore boilers and condensers have to be calculated. In the cases with higher concentration only one rectifier has to be bought. There are two new strippers. To calculate the required area in the reboilers and condensers the same equation as for the heat exchangers are used (see equation 2). All the trays are spaced equally with 0.61 m in all the columns and the diameter is calculated with Aspen Plus see Appendix B. The results for the different cases are presented in table 14-17.
The decanters design is based on a retention time of 10 minutes and a length diameter ratio of 3[5].
Equation 13, 14 and 15 are used to calculate the size of the decanters.
qV (13)
Ld
V
2
2 (14)
dL 3 (15)
Table 18 presents the result for the six cases.
Table 18 Decanter sizes
Decanter 1 Decanter 2
L (m) d(m) V(m3) L (m) d(m) V(m3)
Case I 4.3 1.4 7.0 2.4 0.8 1.3
Case II 4.1 1.4 5.9 - - -
Case III 4.6 1.5 8.6 2.6 0.9 1.5
Case IV 3.8 1.3 4.8 - - -
Case V 4.1 1.4 6.0 2.4 0.8 1.2
Case VI 3.8 1.3 4.8 - - -
Case VII 3.8 1.3 4.8 - - -
Page 55 of 81
Appendix D Investment costs
The first calculations refer to the first case where the process is designed according to the bacteria
efficiency found in the literature. Detailed description of the module costs are presented for case I while
the other cases are summarized in tables.
The module costs are calculated with Ulrich’s method. Ulrich method consists of databases for
purchased equipment costs (Cp). Adjustment factors such as FBM for installation, FP for pressure and FM
for material choice are then introduced to calculate the bare module cost CBM for each process unit see
equation 1 and 2. This value contains the direct and indirect cost for every process unit. Values on Cp,
FBM, Fp, and FM are collected from Projekterings Handboken 2007.[19]
CBM=Cp FBM (1)
CBM=CP FP FM (2)
Pumps
The extra pumps needed in the process will be the same for all cases. No big differences in energy
demand between the different cases were observed why the cost for pump work will be the same for all
cases. According to the Aspen Plus simulations two more pumps must be installed, one before the
strippers and one before the filter press. The pumps are assumed to be centrifugal pumps with an
efficiency of 80 %. The material is stainless steel. The price of the pumps is based on their energy
demand, Sw .
Pump 1
$ 450 9C5.49.1
$ 100 23375.0BM
BMMP
PS
FFF
CkWw
Pump 2
$ 500 315.49.1
$ 700011BM
BMMP
PSC
FFF
CkWw
Total cost (includes two pumps with electric engines)
CBM = 40 950 $
Page 56 of 81
Case I
Fermentation and hydrolysis tanks
The tanks are assumed to be storage tanks with conical roof. The volume of the tanks is 1000 m3 each
but the working volume is 800 m3. The material used is stainless steel. No consideration to the reaction
enthalpy is taken and therefore no cooling system required is designed. Nine tanks already exist why 14
more need to be built. For one tank Ulrich method gives:
BMP
BM
CCsteelstainlesF
mV$ 000 30
) (5.4
1000 3
135 000 $
The agitation requirement is approximated to be about 20 W/m3 which results in 16 kW needed for each
tank [M].
$ 500 42$ 000 17) (5.2
16BMP
BM
CCsteelstainlesF
kWP
Total cost (includes 14 tanks with agitation)
CBM = 2.485 million $
Pre heating before distillation
Heat exchangers are described as fixed tube sheet/U-tube.
HE 1
000 4863
$ 000 880 2
BM
BMMP
CFFF
CpmA $
HE 2
000 4863
$ 000 875 2
BM
BMMP
CFFF
CpmA $
HE 3
000 7263
$ 000 12140 2
BM
BMMP
CFFF
CpmA $
HE 4
000 3063
$ 000 520 2
BM
BMMP
CFFF
CpmA $
Page 57 of 81
HE 5
000 2763
$ 500 415 2
BM
BMMP
CFFF
CpmA $
Total cost (includes all heat exchangers)
CBM = 225 000 $
Distillation
In the distillation process two new strippers need to be installed. These modifications are considered
necessary because of all the extra water present in the system. Two more rectifiers also have to be
installed, one to separate acetone and ethanol and one to obtain pure butanol.
The material used is stainless steel. For equipment costs the towers are assumed to be described as
vertically oriented vessels. Trays are assumed to be sieve trays and are spaced equally with 0.6096 m.
Reboilers are described as kettle reboilers and condensers as spiral plate heat exchangers.
Stripper 1
Vessel:
$ 400 235
7.108.4) (0.4
)4.2(2.1
$ 000 2239.1
4.10
BMBM
BMM
M
FCpC
FFFpsteelstainlesF
barFp
CpmDiameter
mHeight
Trays:
$ 060 13
16),16(2.1
)39.1$( 340
) (0.2
BM
q
BM
C
Ntraysf
mdiameterCp
steelstainlesF
Reboiler
Heating area: 393 m2
$ 000 39393.3100/$000 20) (5
)4.2(12
BMBM
BM
FCpCmCpsteelstainlesF
barFp
Page 58 of 81
Stripper 2
Vessel
$ 500 267
7.108.4) (0.4
)9.0(2.1
$ 000 2549.1
4.10
BMBM
BMM
M
FCpC
FFFpsteelstainlesF
barFp
CpmDiameter
mHeight
Trays
$ 824 13
16),16(2.1
)49.1$( 360
) (0.2
BM
q
BM
C
Ntraysf
mdiameterCp
steelstainlesF
Reboiler
Heating area: 572 m2
$ 000 57272.5100/$ 000 20)(5
)9.0(12
BMBM
BM
FCpCmCpteelstainlesssF
barFp
Rectifier 1
Vessel
$ 000 214
7.108.4) (0.4
)35.0(2.1
$ 000 2023.1
4.10
BMBM
BMM
M
FCpC
FFFpsteelstainlesF
barFp
CpmDiameter
mHeight
Trays
$ 520 11
16),16(2.1
)23.1$( 300
) (0.2
BM
q
BM
C
Ntraysf
mdiameterCp
steelstainlesF
Reboiler
Heating area: 146 m2
$ 000 14646.1$100/000 20) (5
)35.0(12
BMBM
BM
FCpCmCpsteelstainlesF
barFp
Page 59 of 81
Condenser
$ 000 88$ 000 22) (4
1.62 2
BMBM
BM
FCpCCpsteelstainlesF
mArea
Rectifier 3
Vessel
$ 700 224
7.108.4) (0.4
)35.0(2.1
$ 000 2131.1
4.10
BMBM
BMM
M
FCpC
FFFpsteelstainlesF
barFp
CpmDiameter
mHeight
Trays
$ 904 11
16),16(2.1
)31.1$( 310
) (0.2
BM
q
BM
C
Ntraysf
mdiameterCp
steelstainlesF
Reboiler
Heating area: 26.4 m2
$ 500 45$ 100 9) (0.5
)35.0(1BMBM
BM
FCpCCpsteelstainlesF
barFp
Condenser
$ 000 60$ 000 15)(4
2.21BMBM
BM
FCpCCpeelstainlesstF
mArea
Total cost ( includes towers, trays, reboilers and condensers)
CBM = 2.3 million $
Decanters
Page 60 of 81
The decanters are designed on the basis of a retention time of 10 minutes [Z]. A horizontal oriented vessel is assumed to describe the decanters. The ratio between length and diameter is assumed to be 3. Decanter 1
$ 800 49
3.88.4) (0.4
)35.0(2.1
$ 000 644.1
33.4
BMBM
BMM
M
FCpC
FFFpsteelstainlesF
barFp
CpmDiameter
mLength
Decanter 2
$ 240 23
3.88.4)(0.4
)35.0(2.1
$ 800 281.0
43.2
BMBM
BMM
M
FCpC
FFFpeelstainlesstF
barFp
CpmDiameter
mLength
Total cost of decanters: CBM = 73 040 $
Evaporators
Because of all the water present in the system an additional evaporation line is needed to remove all the
water. The material used is copper coated carbon steel which results in a material factor FM = 3. The
pressure factor FP is 1. In the ethanol plant falling film evaporators were used so it will be used here as
Rectifier 3 L=10.5 m D=1.55 m 0.35 bar 10.7 26 000 278 200
Total - - - - 738 300
Table 4 Trays in distillation
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Stripper 1 D=1.25 m 16 2.0 300 11 520
Stripper 2 D=1.35 m 16 2.0 310 11 904
Rectifier 3 D=1.55 m 16 2.0 360 13 824
Total - - - - 37 248
Table 5 Reboilers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Page 62 of 81
Stripper 1 A= 280 m2 2.4 bar 5 56 000 280 000
Stripper 2 A= 220 m2 0.9 bar 5 44 000 222 000
Rectifier 3 A= 17 m2 0.35 bar 5 7 000 35 000
Total - - - - 537 000
Table 6 Condensers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Rectifier 3 A=23.5 m2 0.35 bar 4 16 000 64 000
Total - - - 64 000
Table 7 Decanter
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Decanter 1 D=1.36 m L= 4.1 m 0.35 bar 8.3 6000 49 800
Total - - - - 49 800
Table 8 Evaporators
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Effect 1 A= 515 m2 <10 bar 3 910 000 2 730 000
Effect 2 A=525 m2 <10 bar 3 930 000 2 790 000
Effect 3 A= 640 m2 <10 bar 3 1 100 000 3 300 000
Effect 4 A= 735 m2 <10 bar 3 1 200 000 3 600 000
Effect 5 A=3750 m2 <10 bar 3 2 900 000 8 700 000
Total 21 120 000
Total cost case II CBM = 26.9 million $
CBM2004 = 42.37 million $
Case III 20% higher yield
Page 63 of 81
Table 9 Fermentation/hydrolysis tanks
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Vessel V=1000m3 18 4.5 30 000 135 000
Agitator P=16kW 18 2.5 17 000 42 500
Total - 36 - - 3 195 000
Table 10 Pre heating before distillation
Process unit Design parameter FM*FP FBM Cp ($) CBM ($)
HE 1 A= 80 m2 3 6 8 000 48 000
HE 2 A= 75 m2 3 6 8 000 48 000
HE 3 A=164 m2 3 6 13 000 78 000
HE 4 A= 25 m2 3 6 4 200 25 200
HE 5 A= 20 m2 3 6 4 000 24 000
Total 223 200
Table 11 Distillation towers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 L=10.5 m D=1.64 m 1.4 bar 10.7 28 000 299 600
Stripper 2 L= 10.5 m D= 1.76 m 0.9 bar 10.7 30 000 321 000
Rectifier 1 L=10.5 m D=1.26 m 0.35 bar 10.7 20 000 214 000
Rectifier 3 L=10.5 m D=1.42 m 0.35 bar 10.7 24 000 256 800
Total 1 091 400
Table 12 Trays in distillation
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Stripper 1 D=1.64 m 16 2.0 390 14 976
Stripper 2 D=1.76 m 16 2.0 400 15 360
Rectifier 1 D=1.26 m 16 2.0 310 11 904
Rectifier 2 D=1.42 m 16 2.0 340 13 056
Total 55 296
Table 13 Reboilers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 A= 525 m2 2.4 bar 5 105 000 524 000
Stripper 2 A= 770 m2 0.9 bar 5 150 000 768 000
Rectifier 1 A= 150 m2 0.35 bar 5 29 500 147 000
Rectifier 3 A=27 m2 0.35 bar 5 192 000 9 600
Total 1 448 600
Page 64 of 81
Table 14 Condensers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Rectifier 1 A=65.5 m2 0.35 bar 4 22 000 88 000
Rectifier 3 A =23.5 m2 0.35 bar 4 16 000 64 000
Total 152 000
Table 15 Decanter
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Decanter 1 D=1.54 m L= 4.6 m 0.35 bar 8.3 7 000 58 100
Decanter 2 D= 0.86 m L=2.6 m 0.35 8.3 2 600 21 580
Total 79 680
Table 16 Evaporation
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Effect 1 A= 667 m2 <10 bar 3 1 100 000 3 300 000
Effect 2 A=680 m2 <10 bar 3 1 100 000 3 300 000
Effect 3 A= 830 m2 <10 bar 3 1 300 000 3 900 000
Effect 4 A= 944 m2 <10 bar 3 1 400 000 4 200 000
Effect 5 A=4812 m2 <10 bar 3 3 500 000 10 500 000
Total 25 200 000
Total cost case III CBM = million $
CBM2004 = 52.044 million $
Page 65 of 81
Case IV 20 % higher yield 98 % selectivity
Table 17 Fermentation/Hydrolysis tanks
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Vessel V=1000m3 18 4.5 30 000 135 000
Agitator P=16kW 18 2.5 17 000 42 500
Total - 36 - - 3 195 000
Table 18 Pre heating before distillation
Process unit Design parameter FM*FP FBM Cp ($) CBM ($)
HE 1 A= 80 m2 3 6 8 000 48 000
HE 2 A= 75 m2 3 6 8 000 48 000
HE 3 A=170 m2 3 6 14 000 84 000
HE 4 A= 30 m2 3 6 4 900 29 400
HE 5 A= 20 m2 3 6 4 000 24 000
Total 233 400
Table 19 Distillation towers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 L=10.5 m D=1.40 m 2.4 bar 10.7 20 000 214 000
Stripper 2 L= 10.5 m D= 1.55 m 0.9 bar 10.7 25 000 267 500
Rectifier 3 L=10.5 m D=1.65 m 0.35 bar 10.7 27 000 288 900
Total 770 400
Table 20 Trays in distillation
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Stripper 1 D=1.35 m 16 2.0 320 12 288
Stripper 2 D=1.55 m 16 2.0 360 13 824
Rectifier 3 D=1.65 m 16 2.0 400 15 360
Total 41 472
Table 21 Reboilers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 A= 330 m2 2.4 bar 5 66 400 332 000
Stripper 2 A= 280 m2 0.9 bar 5 56 000 280 000
Rectifier 3 A= 12 m2 0.35 bar 5 6 700 33 500
Total 645 500
Page 66 of 81
Table 22 Condensers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Rectifier 3 A = 24 m2 0.35 bar 4 16 000 64 000
Total 64 000
Table 23 Decanters
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Decanter 1 D=1.30 m L= 3.8 m 0.35 bar 8.3 5 000 41 500
Total 41 500
Table 24 Evaporation
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Effect 1 A= 600 m2 <10 bar 3 1 000 000 3 000 000
Effect 2 A= 615 m2 <10 bar 3 1 050 000 3 150 000
Effect 3 A= 750 m2 <10 bar 3 1 200 000 3 600 000
Effect 4 A= 855 m2 <10 bar 3 1 300 000 3 900 000
Effect 5 A= 4350 m2 <10 bar 3 3 200 000 9 600 000
Total 23 250 000
Total cost case IV CBM = 23.98 million $
CBM2004 = 37.78 million $
Page 67 of 81
Case V 20 % higher concentration
Table 34 Fermentation/Hydrolysis tanks
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Vessel V=1000m3 10 4.5 30 000 135 000
Agitator P=16kW 10 2.5 17 000 42 500
Total - 20 - - 1 775 000
Table 26 Pre heating before distillation
Process unit Design parameter FM*FP FBM Cp ($) CBM ($)
HE 1 A= 80 m2 3 6 8 000 48 000
HE 2 A= 75 m2 3 6 8 000 48 000
HE 3 A=115 m2 3 6 10 000 60 000
HE 4 A= 20 m2 3 6 4 000 24 000
HE 5 A= 15 m2 3 6 3 400 20 400
Total 200 400
Table 35 Distillations towers cost
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 L=10.5 m D=1.30 m 2.4 bar 10.7 21 000 224 700
Stripper 2 L= 10.5 m D= 1.2 m 0.9 bar 10.7 20 000 214 000
Rectifier 1 L=10.5 m D=1.25 m 0.35 bar 10.7 20 000 214 000
Rectifier 3 L=10.5 m D=1.25 m 0.35 bar 10.7 21 000 224 700
Total 877 400
Table 28 Trays in distillation
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Stripper 1 D=1.30 m 16 2.0 320 12 288
Stripper 2 D=1.2 m 16 2.0 310 11 904
Rectifier 1 D=1.25 m 16 2.0 310 11 904
Rectifier 2 D=1.25 m 16 2.0 320 12 288
Total 48 384
Table 29 Reboilers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 A= 267 m2 2.4 bar 5 53 400 267 000
Stripper 2 A= 388 m2 0.9 bar 5 77 600 388 000
Rectifier 1 A= 148 m2 0.35 bar 5 29 600 148 000
Rectifier 3 A=13 m2 0.35 bar 5 6 500 32 500
Total 835 500
Page 68 of 81
Table 30 Condensers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Rectifier 1 A= 62 m2 0.35 bar 4 23 000 92 000
Rectifier 3 A = 16.5m2 0.35 bar 4 13 000 52 000
Total 144 000
Table 31 Decanters
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Decanter 1 D=1.40 m L= 4.0 m 0.35 bar 8.3 6000 49 800
Decanter 2 D= 0.8 m L=2.5 m 0.35 bar 8.3 2600 21 580
Total 71 380
Table 32 evaporation cost
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Effect 1 A= 385m2 <10 bar 3 800 000 2 400 000
Effect 2 A= 390 m2 <10 bar 3 800 000 2 400 000
Effect 3 A= 480 m2 <10 bar 3 900 000 2 700 000
Effect 4 A= 550 m2 <10 bar 3 950 000 2 850 000
Effect 5 A=2800 m2 <10 bar 3 2 500 000 7 500 000
Total 17 850 000
Total cost case V CBM = 23.4 million $
CBM2004 = 36.86 million $
Page 69 of 81
Case VI 20 % higher concentration 98 % selectivity
Table 33 Fermentation/Hydrolysis tanks
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Vessel V=1000m3 10 4.5 30 000 135 000
Agitator P=16kW 10 2.5 17 000 42 500
Total - 20 - - 1 775 000
Table 34 Pre heating
Process unit Design parameter FM*FP FBM Cp ($) CBM ($)
HE 1 A= 80 m2 3 6 8000 48 000
HE 2 A= 75 m2 3 6 8000 48 000
HE 3 A=115 m2 3 6 10 000 60 000
HE 4 A= 25 m2 3 6 4200 25 200
HE 5 A= 20 m2 3 6 4000 24 000
Total 205 200
Table 35 Distillation towers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 L=10.5 m D=1 m 2.4 bar 10.7 18 000 192 600
Stripper 2 L= 10.5 m D= 1 m 0.9 bar 10.7 20 000 214 000
Rectifier 3 L=10.5 m D=1.65 m 0.35 bar 10.7 28 000 299 600
Total 706 200
Table 36 Trays in distillation
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Stripper 1 D=1 m 16 2.0 290 11 136
Stripper 2 D=1 m 16 2.0 300 11 520
Rectifier 3 D=1.65 m 16 2.0 400 15 360
Total 38 016
Table 37 Reboilers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 A= 160 m2 2.4 bar 5 32 000 162 000
Stripper 2 A= 130 m2 0.9 bar 5 26 000 132 000
Rectifier 3 A= 10 m2 0.35 bar 5 6 000 30 000
Total 324 000
Page 70 of 81
Table 38 Condensers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Rectifier 3 A = 19.5 m2 0.35 bar 4 15 000 60 000
Total 60 000
Table 39 Decanters
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Decanter 1 D=1.30 m L= 4 m 0.35 bar 8.3 5000 41 500
Total 41 500
Table 40 Evaporation
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Effect 1 A= 335 m2 <10 bar 3 740 000 2 220 000
Effect 2 A= 340 m2 <10 bar 3 740 000 2 220 000
Effect 3 A= 420 m2 <10 bar 3 800 000 2 400 000
Effect 4 A= 480 m2 <10 bar 3 890 000 2 670 000
Effect 5 A= 2450 m2 <10 bar 3 2 200 000 6 600 000
Total 16 110 000
Total cost case VI CBM = 20.86 million $
CBM2004 = 32.9 million $
Page 71 of 81
Case VII 20 % higher concentration 98 % selectivity and 20 % higher yield
Table 41 Fermentation/Hydrolysis tanks
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Vessel V=1000m3 12 4.5 30 000 135 000
Agitator P=16kW 12 2.5 17 000 42 500
Total - 24 - - 2 130 000
Table 42 Pre heating
Process unit Design parameter FM*FP FBM Cp ($) CBM ($)
HE 1 A= 80 m2 3 6 8 000 48 000
HE 2 A= 75 m2 3 6 8 000 48 000
HE 3 A=125 m2 3 6 10 000 60 000
HE 4 A= 30 m2 3 6 4 900 29 400
HE 5 A= 20 m2 3 6 4 000 24 000
Total 214 800
Table 43 Distillation towers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 L=10.5 m D=1 m 2.4 bar 10.7 18 000 192 600
Stripper 2 L= 10.5 m D= 1.4 m 0.9 bar 10.7 22 000 235 400
Rectifier 3 L=10.5 m D=1.65 m 0.35 bar 10.7 28 000 299 600
Total 727 600
Table 44 Trays in distillation
Process unit Design parameter Number of units FBM Cp ($) CBM ($)
Stripper 1 D=1 m 16 2.0 290 11 136
Stripper 2 D=1.4 m 16 2.0 310 11 900
Rectifier 3 D=1.65 m 16 2.0 400 15 360
Total 38 400
Table 45 Reboilers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Stripper 1 A= 192 m2 2.4 bar 5 35 000 175 000
Stripper 2 A= 252 m2 0.9 bar 5 40 000 200 000
Rectifier 3 A= 12 m2 0.35 bar 5 6 700 33 500
Total 408 500
Page 72 of 81
Table 46 Condensers
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Rectifier 3 A=24 m2 0.35 bar 4 16 000 64 000
Total 64 000
Table 47 Decanters
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Decanter 1 D=1.30 m L= 4 m 0.35 bar 8.3 5000 41 500
Total 41 500
Table 48 Evaporation
Process unit Design parameter Pressure FBM Cp ($) CBM ($)
Effect 1 A= 444 m2 <10 bar 3 850 000 2 550 000
Effect 2 A= 459 m2 <10 bar 3 850 000 2 550 000
Effect 3 A= 562 m2 <10 bar 3 950 000 2 850 000
Effect 4 A= 640 m2 <10 bar 3 1 100 000 3 300 000
Effect 5 A= 3263 m2 <10 bar 3 2 800 000 8 400 000
Total 19 650 000
Total cost case V CBM = 23.27 million $
CBM2004 = 36.7 million $
Page 73 of 81
Appendix E Operating cost and Income
Raw material cost The raw material in the process is wood. The consumption of wood is 21574 kgh-1 D.S and have a price of 50.16 $m-3pub[20]. The density is assumed to be 400 kgm-3pub D.S [21]. The wood cost is calculated with equation 1.
Wood cost = 21574.8000 . = 21.64 million US$year-1 (1)
SO2 cost The consumption of SO2 is 300 kgh-1 and have a price of 0.167 $kg-1[22]. The plant is calculated to run 8000 h per year. The annual SO2 cost is calculated with equation 1. SO2 cost = 300
. 8000 . 0.167= 400 800 US$ (2) Enzyme cost The cost for the enzymes is 3.2US$/106 FPU[23] where FPU stands for Filter Paper Units. and the amount of enzymes that is required is 20 FPU/g cellulose[4]. The cost for the enzymes is calculated with equation 3. Enzyme cost = 13937
. 8000 . 20 . 10-6 . 3.2 = 7.136 million US$year-1 (3) Where 13937 = The amount of Cellulose kgh-1 feed to the process
Fuel purchase and sales Butanol sale The selling price of the butanol is 1.11 $l-1 and calculated with equation 4 The price is set to be the same as that for ethanol based on energy content. Ethanol has a price of 0.92 $l-1[9]. Energy content in butanol 28.48MJl-1 Energy content in ethanol 23.58MJl-1
1
butanol
ethanol
ethanol $L US1.11contentEnergy contentEnergy
priceprice Butanol (4)
IncomeButanol
butanol
ρ
pricem (5)
Page 74 of 81
Acetone sale The price of the acetone is 0.864 $l-1 and the profit is calculated with equation 6[11].
Incomeacetone
acetone
ρ
pricem (6)
Pellets sale The price of pellets is assumed to 0.41$kg-1[24] and the income is calculated with equation 7. Income = mpellets per year . pellets price (7) The income from products sales are shown in table 1. Table 1 Annual income from sales of products in millionUS$
Case I Case II Case III Case IV Case V Case VI Case VII
Butanol Sales 37.4 44.8 44.8 53.9 37.4 44.8 53.9
Acetone Sales 8.2 - 9.8 - 8.2 - -
Pellets Sales - - - - - 2.4 -
Sum (Ii) 45.6 44.8 54.6 53.9 45.6 47.2 53.9
Total annual income Equation 8 is used to calculate the total annual income which consists of income from butanol, acetone and pellets sales.
Annual income Ii =million $ (8)
Energy demand If the process produces less steam than required pellets has to be purchased and burn in the burner. The price of pellets is 0.41$kg-1[24] and have a energy content of 4.76kWh/kg[20].
timeoperating pellets pricepelletsin content energy
demandenergy cost Pellets (9)
The process not only needs steam from the burner but also electric energy to run pump. The price of electric energy is 0.075US$/kWh[25].
timeoperatingenergy electric pricedemandenergy electricyElectricit (10)
The energy demand cost is shown in table 3.
Page 75 of 81
Investment calculation
The base investments G were decided with the total module cost and Ulrichs adjustments factors for
contract/unpredicted co/unf and help facilities fhelp .This adjustment factors include the cost for the work
to design the plant. It also includes costs for unpredicted costs for example delays. These costs are assumed to be 15 % of the total module cost.
1.15f co/un =
The total module cost is calculated in appendix F and the base investment G is calculated in equation 11.
G=Total module cost co/unf⋅ (11)
The base investments were assumed to have a depreciation of 15 years with a real interest rate of 10%. The annual plant cost is calculated see equation 12.
( ) 0.13X11
Xf where
fGcostplant Annual
NA
A
=
+−=
⋅=
−
(12)
The total module cost, base investment and annual plant cost is shown in table 2. Table2 Annual cost and base investment million US$
Case I Case II Case III Case IV Case V Case VI Case VII
Total module cost 44.04 42.37 52.044 37.78 36.86 32.9 36.7
Base investment (G) 50.65 48.73 59.85 43.45 44.67 37.84 42.2
The fixed capital is represented by the storage of raw material. product and spare parts to the process. The raw material and enzymes are assumed to be stored for 30 days and the products are stored for10 days before sold. To calculate the fixed capital from stored products and raw material the annuity factor fA =0.13 was used.
Page 76 of 81
Raw material and enzymes The annual costs for raw material storage and enzymes are calculated with equation 13. Where raw material consist of wood, SO2 and pellets.
Storage cost per year = raw material cost per year . 365
days storage . fA (13)
Products The annual costs for the storage of the products are calculated in the same way as for raw material. The results for the six cases are presented in table 3. Spare parts The cost for spare part storage is calculated as 10% of the maintenance and reparation. The cost of maintenance and reparation is calculated in the chapter direct variable costs in this appendix. Direct variable costs Direct variable costs are the costs for material, process operators, shift management, laboratory staff and maintenance and reparation. Raw material and enzymes The annual costs for the raw material and enzymes have been calculated and are presented in table 3. Maintenance and reparation The cost for maintenance and reparation is 2 % of the annual plant cost calculated with the annuity method. This is presented in table 3. Staff The annual cost for personnel is determined by the number of process operators needed to run the plant. A percentage is applied on the cost for the process operators to determine the costs for the rest of the staff. Process operators The amount of workers needed to operate the plant is assumed to be 5. The process is a continuous process and therefore 5 shift is applied, therefore 6 shifts teams are needed. The plant has an operating time of 8000 hours per year. The operators are assumed to have a monthly salary of 3344 US$. Equation 14 is used to calculate the annual cost for the operators. For both shift management and laboratory staff a percentage of 15 % of process operators is applied this is shown in equation 15 and 16. Process operators 5. 6.12.3344 = 1.2 million US$ (14) Shift management
15% of shift staffs $18.02.10.15 millionUS (15)
Page 77 of 81
Laboratory staff
15% of shift staffs US$18.02.10.15 million (16) Indirect variable costs The indirect variable costs consist of overhead for staff, administration and research and development. Overhead for staff Overhead for staff is estimated to be 70 % of the cost for operating personnel and to 50 % of the management and laboratory staff. These costs are shown in table 3. Administration The administration cost work is set to 25 % of the overhead for staff. Se table 3. Research and development (R&D) The cost for R&D for this plant is assumed to correspond to 2 % of the total operating costs. Total annual payments Equation 17 is used to calculate the total annual payments which consist of fix capital direct and indirect variable costs.
Annual outgoing payments. Ui =million $ (17)
Page 78 of 81
Table 3 Annual payments in million US$
Fixed capital Case I Case II Case III Case IV Case V Case VI Case VII
Butanol production cost The butanol production cost per liter is calculated with equation 20 and 21.
GfUI AiapbN (20)
Butanol production cost per liter =-b
b
P
N (21)
Where Iap= Annual income from acetone and pellets sales Pb = Annual production of butanol in liter The butanol production cost is calculated in table 5 for the different cases. Table 5 Calculated butanol production cost in US$
Case I Case II Case III Case IV Case V Case VI Case VII
Butanol production cost 1.13 1.14 1.24 1.23 0.92 0.89 0.92
Pay-back method The payback time for the investment, n, is calculated with a real interest rate of X= 10 %. The payback time is calculated with equation 22 and the result is shown in table 6. If the process has an annual deficit, it will never be profitable.
X)ln(1
aXGln(1n i
(22)
Table 6 Number of years needed to pay back the investment
Case I Case II Case III Case IV Case V Case VI Case VII
Pay-back year - - - - 4.7 3.3 3.5
Page 80 of 81
Appendix F Investment costs Ethanol plant
The ethanol plant the butanol plant was based upon had the following investment costs calculated:
Step Investment Cost (million $)
Raw materials handling 4.15
Hydrolysis/Pretreatment 42.2
Fermentation/SSF 48.9
Distillation 13.6
Evaporation 15.4
Drying, pelleting 22.3
Steam generation 22.9
Water purification 13.7
Stock 15.2
Total Investment Cost 198.3
This data was taken from a report on cellulosic ethanol using SSF[3].
