Alternatives for Bio-Butanol Production

Alternatives for Bio-Butanol Production Feasibility study presented to Statoil 12/5-2011 Tutors Sixten Dahlbom Hans T. Karlsson Hanna Landgren Christian Hulteberg Peter Fransson Industry advisor/statoil representatives Børre Tore Børresen Per Nygård I

Abstract Title: Authors: Advisors: Course: Aim: Methodology: Theory: Data collection: Results: Key words: Alternatives for Bio-Butanol Production Sixten Dahlbom, Hanna Landgren and Peter Fransson Hans T. Karlsson, Christian Hulteberg, Per Nygård and Børre Tore Børresen KET050 (Projektering) The aim of this project is to improve the production of butanol through comparison with a butanol process investigated in 2008, and some other alternative butanol processes. The processes consist of: Identical butanol production plant as in the report from 2008, but with other purification processes. As well as using a similar ethanol plant as in the report from 2008 but with a chemical conversion of ethanol into butanol. The latter can be performed in a single- or multi-step procedure. For these alternatives the group will consider energy balances as well as investments and production costs. The group will first perform a literature survey and then simulate the best alternative for production of butanol. When the simulation is finished the group will calculate the costs of production and investment. ASPEN PLUS is used for the simulations and the Ulrich method is utilized for economical calculations. Mostly secondary data will be used; electronic sources and books. Also some primary data will be used; help from the advisors are used in this project. The results point at obvious flaws in the production of butanol through indirect synthesis using a condensation reactions, these drawbacks are mostly consisting of the poor productivity and large production of side products. None of the studied alternatives seems to be profitable, although out of the five compared alternatives the two perstraction methods are the most promising. The two perstraction methods have pay-off times of 15.2 years compared to the process using HAP as a catalyst which has a pay-off time of 40.1 years, with the current cost and revenue situation. Vacuum Distillation, Butanol Production, Separation Methods, Investment Costs, Operating Costs I

Table of contents 1 Introduction... 1 1.1 Company Presentation... 1 1.2 Aim... 1 1.3 Method... 2 1.3.1 The Simulation Program Aspen plus... 2 1.4 Economic Assessment... 2 1.5 The Feedstock... 2 1.6 The Bacteria, Clostridiaceae... 3 1.7 Background... 3 1.9 Risk Aspects in the Butanol Process... 4 2 Separation Processes to the ABE Fermentation... 5 2.1 Gas Stripping... 5 2.2 Membrane Separation Methods... 6 2.2.1 Perstraction... 6 2.2.2 Simulation Results for Perstraction... 8 2.3 Adsorption... 12 2.4 Liquid-Liquid Extraction... 14 2.5 Vacuum Distillation... 16 2.5.1 Simulation Results for Vacuum Distillation... 18 3 Direct and Indirect Butanol Synthesis from Ethanol... 20 3.1 Direct Synthesis... 20 3.1.1 Hydroxyapatite (HAP) as Catalyst... 20 3.1.2 Magnesium Oxide (MgO) as Catalyst... 23 3.1.3 Zeolites as Catalyst... 23 3.1.4 Process Flow Diagram, Direct synthesis... 23 3.1.5 Simulation Results for Direct Synthesis... 25 3.1.6 Further Investigations and Improvements... 28 3.2 Indirect Butanol Synthesis... 28 3.2.1 Ethanol Acetaldehyde... 28 3.2.2 Acetaldehyde Crotonaldehyde... 30 3.2.3 Crotonaldehyde Butanol... 30 3.2.4 Process Flow Diagram, Indirect Synthesis... 31 4 Financial Analysis... 32 4.1 Investment Expenditure for the five Alternatives... 33 4.2 Operating Costs for the five Alternatives... 35 4.3 Revenue for the five Alternatives... 37 4.4 Production cost for the five Alternatives... 37 4.5 Net Present Value for the five Alternatives... 37 4.6 Pay-Off Time for the five Alternatives... 38 4.7 Sensitivity Analysis for the Five Alternatives... 39 4.8 Financial Comparison of the five Alternatives... 41 5 Recommendations... 43 5.1 Synthesis of Butanol from Ethanol... 43 5.2 Separation Processes... 43 5.3 Economy... 43 7 References... 44 II

1 Introduction For 20 weeks, a group consisting of four students from Lund University in Sweden will work together with Statoil ASA, Norway. The group will continue the study made by Larsson, E. et al. [1] : A feasibility study on conversion of an ethanol plant to a butanol plant. More in detail, the group will investigate different separation methods for purification of the products from the butanol fermentation described by Larsson, E. et al. [1]. The group will also study whether it is economically sustainable to convert ethanol to butanol. The ethanol will be produced by the existing plant described by Larsson, E. et al. To assist the group, tutors from both Lund University (Hans T. Karlsson and Christian Hulteberg) and from Statoil ASA (Børre Tore Børresen and Per Nygård) are involved. 1.1 Company Presentation Statoil is an international company with activity in 34 countries and with the headquarters in Norway. The company is large, with around 20 000 employees worldwide and is publicly traded on the New York stock exchange. Statoil wants to grow even more and operate in more countries than they do today. [2] Figure 1 below shows in which countries Statoil has offices today. [3] Figure 1: Shows which countries Statoil operates in. [3] Statoil produces electricity, oil, methanol as well as ethanol and has a large focus on natural recourses. Because of the fact that Statoil believes that natural gas will be more and more important in the future, they put large effort on finding supplies and harvesting it. Even though Statoil is currently coupled to negative environmental impact they try to compensate for it by paying respect to issues that can be improved in order to decrease their impact on the environment. [2] 1.2 Aim In order to reduce the dependence of fossil fuels, production of butanol as a gasoline substitute, diesel substitute or additive is an interesting alternative. Butanol is historically mainly produced by fermentation, but the toxicity of butanol to the organism producing it leads to low production, the concentration in the fermenter will be at most 1-2wt% butanol. However, today almost all butanol is produced from propane using the oxo process. Another problem with the fermentation route is that the butanol has to be separated from the 1

fermentation broth, which costs a lot of money and energy. To improve the process of butanol production, the group will compare a process studied in 2008 [1] with some other processes. These are: Identical butanol plant as in the report from 2008 [1], but with other purification processes. Using a similar ethanol plant as in the report from 2008 [1] and afterwards convert ethanol to butanol. This can be performed in a single- or multi-step procedure. For the three alternatives listed above the group will consider: Energy balances Investment costs Production costs 1.3 Method The project will start by making a literature survey where facts will be gathered from primarily peer-reviewed articles, but also from books. The literature study will be used as a basis for performing calculations on parameters such as investment and production costs, as well as energy efficiencies. Simulations of the process will be performed in the computer program Aspen. Data that might be of interest can be viewed in appendix A. 1.3.1 The Simulation Program Aspen plus Aspen plus is a flow sheeting software, that lets the user draw large and advanced industrial processes. It also simulates and calculates mass and energy balances. The software includes physical properties for a numerous chemicals, and features a vast variety of models for phase equilibrium calculations. Aspen will be used in the project for simulation of the overall process, along with the occasional integration of MATLAB code. 1.4 Economic Assessment The production economy will be calculated using the Ulrich method where the results contain both operation and investment costs. 1.5 The Feedstock The feedstock that will be used to produce butanol is lignocellulose, which is a combination of cellulose, hemicellulose and lignin. Lignocellulose is the largest of the world s carbon based renewable natural resources and exists in plants and trees. [4] The lignocellulose to be used in this project consists of 40-50% cellulose 15-25% hemicellulose and 15-30% lignin. Cellulose is a polysaccharide consisting of multiple glucoses on a chain, hemicelluloses is similar but consists of several different sugars, not only glucose. Lignin is a large macromolecule which acts as a glue binding together all the polysaccharides. These three substances are the main components of trees and the lignocellulose to be fermented. Since the sugars are bound in polysaccharides they will need to be broken down before the butanol producing cells can utilize them as feed. The composition of spruce and its bark is seen in table 1 below. [1] 2

Table 1: The raw material composition. [1] Component Spruce Bark Glucan 44.0 31.3 Xylan 6.0 4.0 Galaktan 2.3 3.4 Arabinan 2.0 5.0 Mannan 13.0 4.2 Lignin 27.0 31.9 Ash - - Other 3.5-1.6 The Bacteria, Clostridiaceae To produce butanol a bacteria family called Clostridiaceae is used. These bacteria produce solvents (butanol, ethanol and acetone, therefore the fermentation is called ABEfermentation), organic acids (acetate and butyrate) and gases (carbon dioxide and hydrogen). Depending on which bacteria that is used, different parameters are relevant. One possible bacterium is Clostridium saccharoperbutylaccetonicum which during anaerobic immobilized cell batch systems can reach productivities of 0.36gABE L -1 h -1 at PH 6.0 and a temperature of 30 C. C. Saccharoperbutylaccetonicum is a hyper butanol producing strain and it is somewhat resistant to the product inhibition, therefore it can produce a higher concentration of butanol than the commonly used bacteria. In large scale production it might be best to use immobilized microbial cells since they have many advantages, they are e.g. easier to separate from the product, they can reach a higher cell density and the productivity is greatly improved. At most these bacteria can grow up to a total solvent concentration of 20gABE L -1. The results are based on cultures growing in a medium containing a pure carbon source, in our case a complex medium is used consisting of lignocellulose as carbon source which will most likely lower the production rate. [5] 1.7 Background The purity of butanol shall be 99.5 wt%. The amount of ingoing dry spruce is 181 000 tons per year. In 2008 Larsson, E. et al. [1] made a feasibility study on converting an ethanol plant to a butanol plant. A flow diagram describing the existing ethanol plant can be viewed in figure 2. To convert the ethanol plant to a butanol plant the following changes were suggested: The yeast in the fermentation step shall be changed to the bacteria Clostridium Acetobutylicum. The total fermenter volume is to be increased. The distillation part of the plant has to be expanded, this because much more water is included in the butanol processes. As with the distillation, the evaporation part is to be expanded. For a detailed description of the existing ethanol plant or the suggested modifications please review the report by Larsson, E. et al. [1]. A simplified flow diagram describing the butanol process is shown in figure 3. 3

Figure 2: Simplified flow diagram over the existing ethanol plant [1] Figure 3: Simplified flow diagram over the, by Larsson et al., suggested butanol process. 1.9 Risk Aspects in the Butanol Process In the butanol process there are only a few risks concerning the extra steps with regards to producing butanol instead of ethanol. The bacteria cultures used (Clostridium) can cause stomach problems if ingested [6]. Of course there are some risks coupled to the equipment but the risks are not larger than for the existing production of ethanol. 4

2 Separation Processes to the ABE Fermentation The high cost of distillation in the current ABE process [1] is a consequence of product inhibition. Due to the inhibitory effect, butanol can only be produced as a diluted component. In order to lower the cost for product purification different methods have been studied, these are: Gas stripping Perstraction Pervaporation Adsorption Liquid-liquid extraction Vacuum distillation A comparison between different refining methods was published [24], the results are shown in table 2 and 3. Table 2: A comparison between different refining methods [24] Stripping Adsorption Extraction Pervaporation Capacity Moderate Low High Moderate Selectivity Low Low High Moderate Fouling Low High Moderate Low Operational High Low Low High Table 3: A comparison between different refining methods, A-acetone, B-butanol, and E-ethanol [24] Recovery method Product Estimated total heat of recovery (MJ/kg ABE) Stripping B and AE mixture 21 Adsorption B and AE mixture 33 Extraction and perstraction ABE mixture 14 Pervaporation B and AE mixture 9 2.1 Gas Stripping Gas stripping is a simple way to separate butanol because it does not require expensive equipment and it lowers the butanol concentration. Inert gas will be sparged through the fermentation broth during fermentation and volatile butanol will vaporize and go out with the gas stream in the top of the reactor. Inert gases are appropriate to use in gas stripping, examples of such can be N 2 and CO 2 (care will have to be taken with the ph-decrease associated with CO 2 dissolving into the liquid). Gas stripping is a better technique than membrane filtration and adsorption since it does not remove the reaction intermediates from the broth. [17] Gas stripping is also viewed as one of the most economic techniques. [18] One disadvantage is that the butanol has the highest boiling point of the solvents and therefore is disfavored when the stream takes up the volatile substances from the fermentation broth. The concentration of ABE will be high in the gas stream and therefore contain much butanol [17]. N. Qureshi and H.P. Blaschek found that by using gas stripping as shown in figure 4 the productivity can be increased by 41% and the bacteria will use 50% more of the lactose than normal. The concentration in the recovery stream can be as high as 75.9 g ABE l -1 which is much better than in a batch reactor. Of course these results are from lab scale experiments so it is not sure that the same numbers will be reached in a full scale. [17] The effect on the butanol 5

production varies with the gas recycle rate, acetone and ethanol concentrations in the broth and gas bubble size. [19] Figure 4: Flow sheet for gas stripping. 2.2 Membrane Separation Methods Two methods involving a membrane for purification have been investigated by the group. These two methods are pervaporation and perstraction. [20,21,22] 2.2.1 Perstraction The substance to be separated diffuses through a membrane into a solvent on the other side. This way of separating liquids does not differentiate much from the method known as extraction. However there is at least one benefit; the bacteria is separated from the solvent. This enables the use of a solvent harmful to the bacteria. It is said that the diffusion of butanol through the membrane is the rate controlling step [20]. However as long as the production rate in the reactor is even lower, this should not be a problem. Qureshi, N. and Maddox, I.S. have studied the reduction in butanol inhibition using perstraction. They used a silica tubing as membrane and oleyl alcohol as solvent. As substrate they used lactose/whey permeate. Their result was satisfying; a lactose concentration of 227gL -1 could be used (compared with 28.6gL -1 normally), at this lactose concentration the productivity was 0.07gL -1. Using perstraction ca. 99 gl -1 h -1 (fermentation broth volume) ABE was produced. The experiment was run for 391 hours and the concentration of ABE in the oleyl alcohol was at maximum 9.75gL -1. Figure 5 shows butanol, ethanol and butyric acid concentrations in oleyl alcohol at various levels of these chemicals in aqueous phase during the experiment. According to Qureshi, N. and Maddox, I.S no acetone can be found in the organic solvent. 6

Figure 5: Butanol, ethanol and butyric acid concentrations in oleyl alcohol at various levels of these chemicals in aqueous phase during fermentation perstraction experiment. [20] In situ product separation in butanol fermentation by membrane-assisted extraction has been studied by Jeon, Y.J.; Lee, Y.Y. [21]. They used the same membrane and the same organic solvent as Qureshi, N. and Maddox, I.S. The results of Jeon and Lee are depicted in table 4. Table 4: Overall performance of membrane-extractive butanol fermentation. [21] In figure 6 (below), a flow diagram that shows where the perstraction shall be implemented in the ABE process. 7

Figure 6: Simplified flow diagram over the ABE process using perstraction. 2.2.2 Simulation Results for Perstraction First of all it is important to mention that it is not sure that any suitable membrane exists, therefore these results are speculative. However, as long as the solvent is not harmful to the bacteria, about the same result can be achieved using extraction. Since the bacteria are separated from the solvent by the membrane, a wide range of solvents can be used. The group has mainly focused on two different solvents, mainly due to limited time. The focus was on a 50/50 wt% decane/oleyl alcohol mixture and on mesitylene. The reasons: the oleyl/decane mixture seems to be the most commonly used in experimental reports and mesitylene was suggested by a research group to lower the energy demand in the case of separating the products using extraction [35]. According to this group s results as low as 5.7 MJ/kg butanol was needed using mesitylene (compared to 15 MJ/kg butanol using pure oleyl alcohol). Since the solubility of mesitylene in water is extremely low they suggested that the toxicity towards the bacteria might be negligible. If not, this can be solved using perstraction, as long as a membrane with sufficient properties exists. The fact that acetone, ethanol and butanol are removed continuously allows the process to be run as a fed-batch process. Therefore the project group estimated that 20 fermenters are needed (compared to originally 23), the retention time is 60 hours. The volume of each fermenter shall be 1 039 m 3 using mesitylene as solvent (the fermenter volume is 1000 m 3, the extra volume is reserved for the solvent) and 1 062 m 3 using oleyl alcohol/decane as solvent. Inside every fermenter a membrane shall be placed, figure 7. The effective fermentation volume is 800 m 3. 8

Since the process is run as a fed-batch, fewer fermenters are needed, however they will probably have to be washed and the lignin in the feed has to be removed somehow, emptying the tank seems to be the easiest way of doing it. The outlet from the fermenter consists mostly of water and lignin. When the fermenter is emptied, the outlet contains the same mass of lignin as the mass of lignin retained from separation (only) by distillation, based on the same time. The outlet is passed to an evaporator line. The design of the evaporation line depends on how often the fermenters are to be emptied (energy demand and sizing), this will be left to later studies. Due to time limitations, the group will consider the energy and the size to be the same as in separation by distillation. The same goes for the pretreatment (size and energy). Figure 7: A fermenter fitted with a membrane which on the other side has the solvent flowing Depending on the membrane s properties, different answers will be returned. The group has made two assumptions, these are: Only acetone, butanol, ethanol and water can diffuse through the membrane. The mass transfer limitations are small enough to be neglected (since the production rate is low). If not, a larger membrane area is needed to obtain the same result. Whether water can diffuse through the membrane or not, as well as if the limitations in mass transfer can be neglected, can be discussed. 2.2.1.1 Decane/Oleyl Alcohol A suggested flow sheet is depicted in figure 8. Five distillation columns are needed, these will operate under different pressures and energy is to be added to two of them (table 5). A total of 16.1 MW is needed for separation and 3 412 kg pure butanol will be produced every hour. This means that 17.0 MJ/kg butanol is needed. Important streams are shown in table 6. Table 5: Specifications for the distillation columns needed in the Decane/Oleyl alcohol alternative Col-1 Col-2 Col-3 Col-4 Col-5 Pressure [bar] 0.28 0.28 5.0 1.0 1.0 Reboiler temp. [ºC] 143 115 215 117 74 Energy demand [MW] 13.1-3.0 - - Table 6: Specifications of the streams for the Decane/Oleyl alcohol alternative Mass fraction [wt/wt] Flow Total flow [kg/h] Water Butanol Ethanol Acetone Oleyl alcohol Decane Temp. [ºC] Feed 271 232 0 0,013 7.3e-4 3.6e-3 0.49 0.49 30 Butanol 3 427 0 0,996 4.5e-4 0 0 3.7e-3 40 Et/Bu/Wa 229 0.042 0,051 0.82 0.081 0 6.4e-4 75 Acetone 975 4.0e-3 0 6.0e-3 0.99 0 0 55 Solvent 266 616 0 0 0 0 0,50 0.50 73 9

Figure 8: Flow sheet for Decane/oleyl alcohol alternative 2.2.1.2 Mesitylene A suggested flow sheet can be viewed in figure 9. For the separation of the products only three distillation columns are needed, pressure, reboiler temperature and energy demand for each one of them are shown in table 7. Energy is only needed in the first column (12.6MW), 3 192 kg pure butanol will be produced every hour. This means that 14.2 MJ/kg butanol is needed. Important streams are shown in table 8. The experimental distribution coefficients are reported [35] to differ a lot from the coefficients generated by the UNIFAC model. The distribution coefficients are (in the same report) also reported to be quite a lot higher at 80ºC than at 25ºC, therefore the perstraction will be performed at 80ºC. In order to keep the temperature in the fermenter at 37ºC, it need be cooled instead of heated. Table 7: Specifications for the distillation columns needed in the Mesitylene alternative Col-1 Col-2 Col-3 Pressure [bar] 1 1 0.7 Reboiler temp. [ºC] 164 117 76 Energy demand [MW] 12.6 - - Table 8: Specifications of the streams for the Mesitylene alternative Mass fraction [wt/wt] Flow Total flow [kg/h] Water Butanol Ethanol Acetone Mesitylene Temp. [ºC] Feed 211 717 1.1e-3 0.016 9.3e-4 4.6e-3 0.98 80 Butanol 3 193 0 1.0 0 0 0 117 Et/Bu/Wa 658 0.35 0.35 0.29 4.6e-3 0 77 Acetone 987 3.6e-3 0 4.8e-3 0.99 0 46 Solvent 206 880 0 0 0 0 1.0 164 10

Figure 9: Flow sheet for the Mesitylene alternative 2.2.3 Pervaporation Pervaporation and different membranes for this application has been studied by numerous researchers. The group has focused on one report (due to limited time) and therefore it cannot be guaranteed that this setup is the best, but it is experimentally verified. The report describes ABE recovery by pervaporation using different silicate/silicone composite membranes from a fed-batch reactor [22]. Qureshi, N. et al. made their own membranes, the procedure is well described in their report. The different membranes were characterized for flux and selectivity (defined as in equation 1) at 78 C using model ABE solution and actual fermentation broth, the results are shown in table 9. Selectivity = (y/(1 y))/(x/(1 x)) [Eq. 1] y = weight fraction of component in permeate sample x = weight fraction of component in retentate sample In their further studies they choose the membrane with the highest selectivity towards butanol. Their experimental setup is shown in figure 10. A UF-membrane is used between the reactor and the pervaporation, this since the reactor temperature is 35 C, but the pervaporation temperature is 78 C. 11

Table 9: Flux and selectivity of a silicate-silicone and silicone membranes using model solution and fermentation broth at 78 C. [22] Figure 10: A schematic diagram of ABE production in fed-batch reactor and recovery by pervaporation (a) fermentation reactor; (b) ultra filtration membrane unit; (c) buffer tank; (d) pervaporation membrane unit (e) cold traps. (1.1) cell culture stream; (1.2) ultra filtration membrane permeate; (1.3) pervaporation membrane retentate recycle; (1.4) cell free fermentation broth recycle stream through pervaporation membrane; (1.5) pervaporation membrane permeate stream. [22] Qureshi, N. et al. showed that the butanol selectivity not was affected by the broth. The experiment lasted for 120 hours. They also showed that ethanol and acetic acid not diffused through the membrane at concentrations lower than 0.04gL -1. In total the fed batch reactor was operated for 870 hours and 155 gl -1 solvent was produced. The solvent yield (g solvents/ g glucose utilized) was 0.31 0.35. 2.3 Adsorption A stream of fermentation broth from a butanol producing fermenter can be run through a membrane filter or a centrifuge to separate the biomass from the water, butanol, ethanol and acetone. The solids can then be recirculated to the fermenter while the liquids are run through adsorption columns. The columns can be filled with hydrophobic adsorbents of almost any kind (often silica). In the case of butanol production, the substances which will bond to the column are butanol, ethanol, acetone and some water. The components not adsorbed will be taken back to the fermenter. Thus the fermentation process can be run continuously, see figure 11. 12

Figure 11: Shows the process for absorption of butanol [25]. After the packing in columns have adsorbed butanol it can be desorbed by increasing the temperature to around 200 C. This has been suggested to greatly decrease the energy costs as ordinary distillation would require 73.3 MJ/kg butanol, while adsorption only would need 8.2 MJ/kg. Two columns should be used to ensure continuous production, then the two columns can alternate between being loaded (adsorption) and unloaded (desorption). [1] Research has shown that some zeolites can favor adsorption of butanol over both ethanol and acetone resulting in highly concentrated desorbed solutions. Studies have been made on the adsorption of the ABE-system by zeolites where the feed consists of either filtered or nonfiltered fermentation broth. The results here also showed that butanol is favorably adsorbed on the solid phase adsorbent; results can be seen in table 10. The table shows that about 100g butanol/kg adsorbent can be reached which is a good result, but can be improved further if used on industrial scale as adsorption material and conditions could be optimized further. The desorbed solution will contain over 90% of butanol which is relatively high; the rest will mostly contain acetone and some water. [23] The butanol concentration can then be increased further by distillation [1]. Table 10: Shows the feed concentrations and the adsorbed concentration of two filtered and two non-filtered fermentation broths. [24] 13

2.4 Liquid-Liquid Extraction Liquid-liquid extraction is considered to be suitable for butanol recovery. The unit operation can be placed inside as well as outside the fermentation tank [24]. To choose an appropriate organic solvent in the extraction setup, some considerations have to be made [25] : The solvent has to be compatible with the bacteria used for fermentation. The solvent has to have a high capacity for the fermentation products. This to minimize the amount of solvent needed and the product recovery cost. If the products are recovered from the solution by distillation, the solvent should be less volatile than the products. However it is important that the solvent is not so nonvolatile that high pressure steam is needed in the reboiler. The solvent has to be barely soluble in water, this to minimize solvent losses. A solvent with more or less these properties has been reported by Roffler, S. et al., namely oleyl alcohol diluted to 50 wt% in decane [25]. One can probably think of some other solvents, but to the groups knowledge no other has been reported. It is imported to mention that oleyl alcohol is toxic to the culture when exposed over a long period of time [26]. Data for pure oleyl alcohol and for the oleyl alcohol/decane mixture can be seen in table 11. Table 11: Data for pure oleyl alcohol and for the oleyl alcohol/decane. [26] Extractant Oleyl alcohol Reference Oleyl alcohol /decane (50 wt%) Reference Distribution coefficient (g/l butanol in solvent)/ (g/l butanol in broth) 4 26 2.4 25 Selectivity 180 27 2 000 25 Density [kg/m 3 ] 840 25 780 25 Viscosity [cp] 17 25 3.1 25 Butanol diffusion coefficient [m 2 /s] 1.1*10-10 25 2.3*10-10 25 Specific heating capacity [kj/kg] 2.3 28 N.N. Solubility in water [mg/l] 100 25 N.N Roffler, S. et al. have studied the economic savings that can be made using a fed-batch extraction process for the production of butanol. They came up with a flow sheet describing the process, this can be seen in figure 12. They also came up with, for this feasibility study, some relevant remarks. Based on their calculations the production cost of butanol can be reduced by ca. 20% (year 1987). They also said that the productivity increased from 0.58 g/(l*hr) in batch culture to 1.5 g/(l*hr) in fed-batch culture. When the productivity is increased the numbers of fermenters are reduced. They also conclude that concentrated feedstocks can be fermented and stillage treatment costs are significantly reduced. 14

Figure 12: Process flow diagram of a fed-batch fermentation employing oleyl alcohol/decane as Extractant. [25] 15

2.5 Vacuum Distillation A decreased pressure leads to a lower boiling point. The reduction of pressure means that less energy is needed to boil any components, however, more energy is then needed to sustain the low pressure. A diagram for the boiling points of butanol, ethanol and acetone at different pressures can be seen in figure 13. One has to consider that these boiling points are for the pure elements and not for the ternary mixture. The vacuum distillation set up does not need to be very different from the regular distillation set up. The pressure in the column has somehow to be controlled. In figure 14 the difference between the two processes is shown. Figure 13: A diagram for the boiling points for butanol (dots), ethanol (triangles) and acetone (crosses) at different pressures [30],[31],[32] Figure 14: Vacuum distillation (left), regular distillation (right) [29] In figure 15 (below), a flow diagram that shows where the vacuum distillation shall be implemented in the ABE process. 16

Figure 15: Simplified flow diagram over the ABE process using vacuum distillation. 17

2.5.1 Simulation Results for Vacuum Distillation The main idea behind the vacuum distillation process is to lower the pressure in the distillation columns, see distillation step in butanol process specification [1], and determine the changes in the process, e.g. the overall power demand. Figure 16: Flow sheet over vacuum distillation step The overall process diagram (figure 16) does not differ from the original butanol process [1] and in the calculation the project group assumed that process specifications (e.g. energy demand, size of apparatus and others) are the same as in the old process, except for the distillation step. Pressure reduction was mainly carried out in the stripper part of the distillation process, due to the fact that the main energy demand lies here. Two conditions were taken into accounts when determine the pressure in the stripper columns. 1) The temperature in the high-pressure columns should be high enough to run the reboiler in the low-pressure columns, a pressure differential of 0.9 bar was chosen. 2) The temperature in the top of the low-pressure column should be high enough to use cooling water, which is assumed to have a temperature of 17 C, for liquefaction of the top stream. With the above mentioned conditions, a pressure of 1 bar in the high-pressure columns and 0,1 bar in the low-pressure ones was chosen. Power requirement for the high pressure columns where calculated to 22.7 MW (11.3MW each).the pressure in the decanter is set to 3.5 bar and a temperature of 115 C, for a good separation. The pressure in the three remaining distillation towers kept at atmospheric pressure, except for rectifier number one, which is set to a pressure of 0.5 bar. The reason for this decision is because condense from rectifier number 2 should be able to heat the reboiler in distillation column 1 and 3. No changes were made in the second decanter. 18

Table 12: Distillation tower specifications HP 1 and 2 LP 1 and 2 R-1 R-2 R-3 Pressure [bar] 1.0 0.10 0.50 1.0 1.0 Reboiler temp. [ºC] 100 69 75 117 65 Energy demand [MW] 11.3 - - 4489 - Table 14: Flow specification for vacuum distillation. Mass fraction [wt/wt] Flow Total flow [kg/h] Water Butanol Ethanol Acetone Temp. [ºC] Feed 245 582 0.98 0.014 0.0080 0.0040 37 Acetone 856.4 0.0070 0 0.0060 0.99 92 Water/ethanol 384.8 0.095 0,0059 0.50 0.39 65 Butanol 3485 0 1 0 0 117 19

3 Direct and Indirect Butanol Synthesis from Ethanol There is an existing desire to find an alternative to the well-known ABE-fermentation for industrial production of butanol. One possible alternative is to use a catalyst for a single step synthesis of butanol from ethanol; another is to convert ethanol to acetaldehyde, acetaldehyde to crotonaldehyde and crotonaldehyde to butanol. 3.1 Direct Synthesis As for now, there are no existing industrial size plants in operation which uses this type of reaction. However, there exist numerous reports of experiments using different types of catalysts. The most common once are hydroxyapatit (HAP) [7], magnesium oxide (MgO) [8] and Zeolites. [9] 3.1.1 Hydroxyapatite (HAP) as Catalyst Hydroxyapatite has been reported as a useful catalyst for condensation of ethanol to butanol by Tsuchida et al. al [7]. Non-stoichiometric hydroxyapatite can be written as: Ca (10-Z) (HPO 4 ) Z (PO 4 ) (6-Z) *nh 2 O, where 0<Z<1 and n=0 2.5 The reaction mechanism as suggested by Tsuchida et al. is as follows: Ethanol adsorbs on the HAP-complex and a C-C bond is formed between a β-carbon in the ethanol molecule and an α-carbon in an n-c n H 2n+1 OH molecule. The products are n-c n H 2n+1 CH 2 CH 2 OH and water. Further it was suggested that a part of the formed n-c n H 2n+1 CH 2 CH 2 OH molecule will be adsorbed and form a C-C bond between a β-carbon, in the C n H 2n+1 CH 2 CH 2 OH molecule and an α-carbon in an n-alcohol, and a branched alcohol is produced. As shown above, besides n-butanol the reaction mechanism leads to numerous n-alcohols and branched alcohols as byproducts. These include aromatics, acetaldehyde, ethylene and ether, just to name a few. Takashi et al. determined the optimal condition for the butanol production by varying the ratio between calcium- and phosphorous atoms in the HAP-complex, the temperature and the contact time. The experiments were performed at atmospheric pressure. The highest butanol yield was obtained at Ca/P ratio of 1.64 [7]. At this ratio the ethanol conversion is about 15%. Further it was establish that a temperature of 300 C [7] and a contact time of 1.72 s [7] is best suited to gain as high butanol selectivity as possible; which in this case, according to the report, is 76% [1]. The effects of different contact times and the temperature can be seen in table 15. 20

Table 15: Effect of temperature and contact time on selectivity, taken from Tsuchida et al. [1] As mentioned above, the experiments were carried out at atmospheric pressure, which is not well suited for an industrial process. A process at higher pressures is more suitable. To investigate how an increase in pressure will affect yield and conversion, a MATLAB simulation of the reaction rates were carried out. The reaction rate constants for the main reaction (S1) and the unwanted reactions (S2-S13) are given in table 16 together with activation energy. Table 16: Reaction rates, taken from Tsuchida et al. report [7] The result of the simulation was satisfying, plots showing yield, conversion and selectivity as functions of time can be seen in figures 17 to 19. With an increase in pressure the optimal contact time decreases, the ethanol conversion and selectivity, at optimal contact time, will decrease respectively increase. The butanol yield will however increase. 21

Figure 17: Butanol yield and optimal contact time at different pressures. Figure 18: Ethanol Conversion at different pressures. 22

Figure 19: Ethanol Selectivity at different pressures. As of now, HAP catalytic condensation of ethanol looks promising. However, the largest disadvantage is numerous byproducts which compete with butanol. 3.1.2 Magnesium Oxide (MgO) as Catalyst A.S. Ndou et al. studies of magnesium-oxide catalyzed synthesis of butanol from ethanol show a maximum butanol yield of 20%, byproducts/intermediates include acetaldehyde, crotonaldehyde, crotylalcohol and butanal The experiments were carried out at atmospheric pressure and around 400 o C [8]. A mechanism was proposed by Ndou et al. This is: butanol is derived from two different reaction mechanisms. In the first mechanism ethanol dehydrogenates to acetaldehyde and further acetaldehyde is converted to crotonaldehyde via an aldol condensation, and in the last step crotonaldehyde is hydrogenated to butanol. The second mechanism is the same as the one proposed by Tsuchida et al. (mentioned above). 3.1.3 Zeolites as Catalyst The mechanism behind the condensation reaction of ethanol, using zeolites as catalyst, has been studied [9]. Yang et al. [9] proposed two thinkable mechanisms for reaction, of which both has already been mention above. After analysis of the results from the experiments Yang et al. [9] came to the conclusion that the mechanism proposed by Tsuchida et al. [7] is the most likely one have taken place. The conclusion made by Yang et al. [7] seems to fit the other above mentioned catalysts. 3.1.4 Process Flow Diagram, Direct synthesis As a first approach, the project group suggests a process for single step synthesis of butanol from ethanol, based on the existing ethanol process [1]. 23

Figure 20: Simple process flow diagram over single step synthesis of butanol. The suggested process is the same as the existing ethanol plant, except for some changes after the distillation step (see figure 20). In the original process the ethanol/water stream from the distillation would be directed to a dewatering step to produce pure ethanol. Instead the ethanol/water stream will be sent to a tubular reactor packed with a catalyst. The outlet gas from the reactor will be distillated in a number of steps, to separate butanol and ethanol from byproducts. The separated ethanol will be recirculated to the reactor. 24

3.1.5 Simulation Results for Direct Synthesis The overall process is based on the ethanol plant [1] with an added HAP step, which includes a reactor for conversion of ethanol to butanol and five distillation columns for purification of butanol and ethanol recovery. The simulations and process description (figure 21, tables 17 to 19) below focus on the HAP step, for more information on the ethanol plant see the rapport from 2008 [1]. Figure 21: Flow sheet over HAP step The outlet ethanol stream from the existing ethanol plant, 4 721kg/h ethanol, 217 kg/h water will be vaporized and preheated up to 107 C. Next step is to compress the gas stream in a compressor from 1 bar to 10 bar. Further the gas stream needs to be preheated up to the desired temperature of 300 C and with this the pretreatment of inlet stream to the tube reactor concludes. The reactor is supposed to run isotherm at 300 C and because the overall net reaction is exothermic the reactor needs to be equipped with a cooling jacket The results from the simulation states that 750 kw need to be removed to keep the reactor isothermal, this gives an opportunity to preheat other streams in the process. The optimal size of the reactor were estimated with the help of a MATLAB simulation, the volumetric flow of the inlet stream were estimated to 897 m 3 /h (300 C, 10 bar) and the reactor volume was calculated to 0.24m 3 (the reactor is packed with hydroxyapatite and was assumed to have a void of 0.5, this corresponds to 385 kg catalyst). The outlet stream consist of 1 494 kg/h water, 3 068 kg/h ethanol, 1 787 kg/h butanol, 1 623 kg/h higher alcohols (170 kg/h hexanol, 687 kg/h 2-etyl- 1butanol, 766 kg/h 2-etyl-1butanol and others), 39 kg/h alkenes (13.3 kg/h 1,3-butadien, 9.7 kg/h hexen and others) and 0.75 kg/h hydrogen. The Outlet stream from the reactor will be distilled in the first distillation tower. Here ethanol will be separated from the butanol, where ethanol will go out from the top of the tower and the butanol in the bottom. The alkenes from 25

the inlet stream will pass through the distillation tower with the ethanol over the top and the other alcohols will follow the butanol. A small part of the top stream will be non-condensable, due to the hydrogen in the stream, and will pass over the top out from the process. The energy that needs to be removed in order to condensate the rest of the top stream is large enough to run the reboiler for distillation tower 5 and heat exchanger HX2, for vaporization and preheating of the ethanol feed. In order to achieve the above mentioned energy savings the temperature in the top of the column has to be high enough i.e. the pressure in the distillation column needs to be high enough. A pressure of 7 bar satisfies these conditions. The function of the second distillation tower is to separate butanol and water from higher order alcohols. The higher alcohols will leave in the bottom of the column and butanol/water leaves in the top. The energy received from condensing the top stream is enough to cover the power demand of distillation tower 4. In order to get pure butanol, the top stream from the second distillation column will be separated, in a decanter, into a water phase and an organic phase. The stream has a butanol weight percentage over the butanol/water azeotrope and now the organic phase can be distilled in distillation column 5. Pure butanol will pass through the column in the bottom, 1 699 kg/h, and the top water/butanol stream, 516 kg/h water and 706 kg/h butanol, will be recirculated to the decanter. Distillation tower number 4 separates ethanol and water from alkenes. Alkenes, 7 kg/h water, 133 kg/h ethanol, 4 kg/h butene, 9 kg/h hexene, 5 kg/h acetaldehyde, 12 kg/h 1,3-butadiene, leaves in over the top of the column and ethanol/water, 348 kg/h water, 2919 kg/h butanol, in the bottom. In the last distillation column water, 159 kg/h water, 4 kg/h butanol, pass through the bottom and ethanol/water stream, 188 kg/h water, 2 915 kg/h butanol, at the azeotrope, passes over the top of the column and mixes with the ethanol feed before the pretreatment step. Table 17: Distillation tower specifications for the HAP catalysis alternative Distillation Tower 1 Distillation Tower 2 Distillation Tower 3 Distillation Tower 4 Distillation Tower 5 Pressure [bar] 7 1 1 1 1 Reboiler temp. [ºC] 157 156 118 78 98 Energy demand [kw] 3 737 1 581 833 - - 26

Table 18: Flow specifications for HAP catalysis Stream, mass fraction (wt/wt) Component Feed Inlet Outlet Gas1 Dest1 Bot1 Dest2 Bot2 Water 0.044 0.050 0.19 0.085 0.10 0.25 0.39 0 Ethanol 0.96 0.95 0.38 0.81 0.89 1.3e-4 0 0 Butanol 0 0 0.22 0 1.7e-4 0.39 0.6 0.025 Hexanol 0 0 0.021 0 0 0.037 0 0.1 2-Etyl-1Butanol 0 0 0.086 0 0 0.15 0 0.41 Octanol 0 0 3.7e-3 0 0 6.5e-3 0 0.018 2-Etyl-1Hexanol 0 0 0.095 0 0 0.17 0 0.45 Ethene 0 0 1.0e-4 9.7e-3 1.9e-4 0 0 0 Butene 0 0 5.3e-4 5.6e-3 1.2e-3 0 0 0 Hexene 0 0 1.2e-3 0.028 2.7e-3 0 0 0 Acetaldehyde 0 0 6.4e-4 8.6e-3 1.4e-3 0 0 0 Hydrogen gas 0 0 9.0e-5 0.036 1.3e-5 0 0 0 Octene 0 0 6.8e-4 0 0 1.2e-3 1.9e-3 0 1,3-Butadiene 0 0 1.6e-3 0.016 3.7e-3 0 0 0 Total Flow (kg/h) 4 938 8 042 8 042 19 3 439 4 584 2 888 1 696 Table 19; Flow specifications for HAP catalysis Stream, mass fraction (wt/wt) Component Organ Wat Dest3 Bot3 Gas2 Dest4 Bot4 Dest5 Bot5 Water 0.18 0.96 0.42 0 6.7e-3 0.041 0.11 0.06 0.97 Ethanol 1.0e-3 5.0e-4 2.4e-3 0 0.14 0.78 0.89 0.94 2.2e-2 Butanol 0.82 0.039 0.57 1.0 0 0 1.8e-4 0 3.7e-3 Hexanol 0 0 0 0 0 0 0 0 0 2-Ethyl- 1Butanol 0 0 0 0 0 0 0 0 0 Octanol 0 0 0 0 0 0 0 0 0 2-Ethyl- 1Hexanol 0 0 0 0 0 0 0 0 0 Ethene 0 0 0 0 0.19 1.9e-3 0 0 0 Butene 0 0 0 0 0.11 0.023 0 0 0 Hexene 0 0 0 0 0.17 0.052 0 0 0 Acetaldehyde 0 0 0 0 0.025 0.029 0 0 0 Hydrogen gas 0 0 0 0 0.026 0 0 0 0 Octene 1.0e-3 4.0e-3 2.7e-3 0 0 0 0 0 0 1.3-Butadiene 0 0 0 0 0.33 0.072 0 0 0 Total Flow (kg/h) 2 925 1 191 1 228 1 69 9 1.7 170 3 267 3 103 163 27

3.1.6 Further Investigations and Improvements Further improvements need to be done in minimizing the losses of ethanol and butanol, as for now there is a loss of 88kg/h butanol (46kg/h in the water-phase stream and 42kg/h in the bottom stream of the second distillation column) and 133kg/h ethanol, e.g. recycle the waterphase stream to distillation column number 2. Some improvements can be made in the energy demand and recovery, e.g. with the help of pinch analysis. Further studies on the reaction kinetics need to be carried out in order to determine if there are any restrictions in mass transport or diffusion in relation to the reaction kinetics. 3.2 Indirect Butanol Synthesis Butanol can be produced if ethanol is reacted to acetaldehyde, acetaldehyde to crotonaldehyde and finally crotonaldehyde to butanol [10], reaction (1-3). These three reactions are all well studied and described in the literature. A drawback using this method of producing butanol is the additional need of process equipment. In this chapter the chemical processes used for production of acetaldehyde, crotonaldehyde and butanol will be described. [reac.1a] [reac.1b] [reac.2] [reac.3] 3.2.1 Ethanol Acetaldehyde Acetaldehyde can be produced from ethanol in two different processes; ethanol can either be dehydrogenated (reac.1a) or oxidized (reac.1b). The benefit with the dehydrogenation reaction is the simultaneous production of hydrogen. Hydrogen is needed in reaction three, so the hydrogen produced in reaction one can be used in the third reaction. On the other hand the catalyst life in the oxidation process is longer and the possibility of recovering energy is better [10] (due to the fact that dehydrogenation of ethanol is an endothermic reaction). 3.2.1.1 Dehydrogenation of ethanol In the temperature range 260 310 C ethanol vapor is passed over a catalyst. The catalyst is typically made of copper. However, some other catalysts have also been reported, for instance nickel supported by SnO 2, Al 2 O3 or SiO 2 [11], and palladium or platinum modified alumina [12]. The reason of using copper as a catalyst is that the lowest amount of decomposition products are reported using this catalyst [10]. A drawback using copper as a catalyst is that the catalyst is subjected to rapid deactivation, which results predominantly from sintering [12]. The reaction takes place in a tubular reactor and a conversion of 25 50% is obtained [10] per pass. Acetaldehyde selectivity up to 100% has been reported at 200 C using Ni/SnO 2, Ni/Al 2 O 3 or Ni/SiO 2 as catalyst, at higher temperatures the selectivity is decreasing [11]. In industrial processes the selectivity to acetaldehyde is 90 95% [13]. In order to separate ethanol and acetaldehyde from the exhaust gas (mainly hydrogen), the gas is washed with an ethanol/water mixture. The ethanol which has not reacted is recovered in a 28

distillation step and recirculated. The final acetaldehyde yield in industrial processes is ca. 90 % [10]. A simplified flow sheet describing this process can be seen in figure 22. Byproducts in the dehydrogenation process are butyracetate, crotonaldehyde, higher alcohols and ethylene [13]. Figure 22: Schematic flow sheet over the process for dehydrogenation of ethanol. A reactor, B/C heat exchangers, D/E purification steps [10] 3.2.1.2 Oxidation of Ethanol One of the most known processes is the Veba-Chemie Process [10]. In this process ethanol vapor and air are passed over a silver catalyst at a pressure of 3 bars and a temperature in the interval 500 650 C. The conversion of ethanol varies between 30 50 % per pass with a selectivity of 85 95% [13]. The ethanol which has not reacted and the produced acetaldehyde are removed from the exhaust gas by washing with ethanol. Ethanol and acetaldehyde are separated by distillation. A simplified flow sheet describing the Veba-Chemie Process can be seen in figure 23. The byproducts in this process are acetic acid, formic acid, ethyl acetate, CO and CO 2. Figure 23: Schematic flow sheeting describing the oxidation of ethanol. A reactor, B heat exchanger, C purification [10] step/reactor, D purification step 29

3.2.2 Acetaldehyde Crotonaldehyde The common method to produce crotonaldehyde is the aldol reaction of acetaldehyde followed by dehydration of the acetaldol (reaction 4). [reac.4] In order to produce acetaldol, acetaldehyde is reacted in a tubular reactor at a temperature ranging from 20 up to 25 C. The residence time is several hours [13]. To catalyze the reaction an aqueous sodium hydroxide solution is used [14] (sodium hydroxide will at any time be found as an ion in water). The conversion of acetaldehyde is restricted to 50 60% to limit resin formation and secondary side reactions. The reaction is stopped and the reaction-mixture neutralized by adding acetic acid [13]. The selectivity to acetaldol is 85% and the main byproduct is crotonaldehyde. The dehydration of acetaldol occurs readily in the presence of acetic acid. In industrial processes the dehydration step is implemented in the purification steps. The reactor vessel (where the aldol condensation takes place) is followed by a stripping column. The acetaldehyde which has not reacted is recirculated back to the reactor. The acetaldol is fed to a distillation column, dehydrated and a crotonaldehyde/water mixture is distilled to the azeotrope and separated into water and an aqueous crotonaldehyde phase containing 10% water. Finally, the crotonaldehyde/water mixture is fed into a rectification column. A simplified flow sheet describing the process can be seen in figure 24. Figure 24: The acetaldehyde to crotonaldehyde process [14] 3.2.3 Crotonaldehyde Butanol This reaction can be carried out in both liquid- and vapor-phase processes [15],[16]. If the reaction is to be carried out as liquid reaction, a solvent must be used. The solvent has to be separated from the product later on, this is not needed if the reaction instead is a gas phase reaction. Using a Cu/Al 2 O 2 catalyst, crotonaldehyde conversion up to 100% and butanol selectivity up to 100% can be achieved [16]. The temperature should be at least 150 C. If this catalyst is exposed to sulfur it will be poisoned and has to be replaced. 30

3.2.4 Process Flow Diagram, Indirect Synthesis A flow diagram of the indirect synthesis of butanol is shown in figure 25. The production of ethanol is followed by dehydrogenation (of ethanol). This reaction is followed by formation of crotonaldehyde from the produced acetaldehyde. Finally, crotonaldehyde is reacted to butanol. The dewatering step might not be needed before the reactor, but at the end the water has to be removed somewhere in the process. Figure 25: A flow diagram over the indirect synthesis of butanol from ethanol. 31

4 Financial Analysis To compare the alternatives with each other, some calculations have been done on the investment expenditure and the operating costs. As a basis for the profitability analysis a lifespan of 10 years has been assumed together with an interest rate of 15 %. In the calculations five alternatives have been compared, these are: The butanol production from the 2008 report The same butanol production as in the report from 2008 but with altered separation method by using vacuum distillation Oleyl alcohol/decane perstraction Mesitylene perstraction. Producing butanol from ethanol through a HAP-catalyzed condensation reaction. To compare the five different methods for producing butanol all the boxes from figure 26 below except for the box called Distillation has been estimated to one cost by looking at the report from 2008. This have been done due to lack of time and the fact that more focus have been put on the simulation and on the financial analysis of the new parts in the butanol production. The calculations from the report 2008 do not include the originally suggested ethanol plant and therefore these costs have been added to the investment and operating costs for the alternatives. The HAP-method includes this reports calculated costs and all costs for the ethanol plant suggested by Per Sassner (2007). The costs for the additional investment expenditure for the butanol plant excluding distillation have been estimated to 41.9 MUSD. To these cost, the costs for the different separation processes have been added. After doing this all five methods can be compared with each other and a total cost for producing butanol can be presented. Figure 26: A diagram over the by Larsson et al. suggested butanol process. Additional costs for things such as overhead and R&D have been added to all operating costs including the estimated parts. The investment expenditure, operating cost and the additional costs have then been summarized to one cost for each method of producing butanol by using 32

the net present value method. Also the pay-off time has been calculated for the different methods. To investigate the accuracy and reliability of the calculations a sensitivity analysis has been performed regarding the most important parameters in the calculations. All results for the five alternatives are compared in the chapter called Financial Comparison of the Alternatives. 4.1 Investment Expenditure for the five Alternatives Investment costs for the butanol plant from the 2008 report excluding the separation method and Per Sassner's (2007) ethanol plant have been estimated to 41.9 USD using the equation below. The costs for the total butanol production (41.5) minus the cost for the separation method distillation (3.7) have been calculated. Then the cost has been updated to 2011 costs by using the consumer price index (I) as seen below in equations 2 and 3. I 2004 =279.4 I 2011 =310.11 C 2004 = the cost for the butanol plant excluding the distillation step in 2004 costs. C 2011 = the cost for the butanol plant excluding the distillation step in 2011 costs. [Eq. 2] [Eq. 3] The calculations for the investment expenditure are based on Ulrich method. From the Ulrich method Cp, F M, F P, and are given by tables and graphs. F M is based on the material that the equipment is made from. Cp depends on the equipment size and F P is based on what pressure the equipment needs to withstand. [36] Then C BM (which is the investment cost including additional costs such as transportation and installation) can be calculated by the equations below (equation 4-6) in which F BM is taken from a graph including F M *F P. [37] [Eq. 4] C BM for all unit operations are then summarized and a total plant C BM is given which is multiplied by 1.15 to compensate for the cost of support facilities. Since the tables are from 1982 the total C BM is converted to 2011 years prices using the consumer price index (I) from 1982 and 2011, see equations below [38]. [Eq. 5] I 1982 =119.3 I 2011 =310.11 [Eq. 6] In table 21 the total investment cost for all equipment in the five alternatives are shown in 2011 prices. Additional costs means the cost for the steps before and after the separation and are therefore added to the methods where the original separation method has been switched (Vacuum distillation, Oleyl decane perstraction and Mesitylene perstraction), the HAP- 33

catalysis process and the process from the report 2008 do already include these additional costs. The total costs are the C BM2011 plus the additional costs. For a more detailed table for four alternatives, except for the case presented in the report from 2008, see Appendix E. The investment cost for each separation step is added to the butanol production suggested by the report from 2008. Furthermore the cost for the suggested ethanol plant is added to each alternative since this is not included in the costs for the butanol production plant from the 2008 report (some includes the cost for distillation and some do not depending on if separation is included in the original costs). The cost for the original ethanol plant can be calculated from Per Sassner s thesis. Sassner mentions that a plant using 200 000 tons of raw material per year will cost 1 264 MSEK, the cost can be scaled by Sassner s scaling coefficient (n) using equations 7 and 8.[39] [Eq. 7] [Eq. 8] The new cost must be converted to dollar and brought up to date which results in a cost for the previously mentioned ethanol plant of 209.3 million USD. If the distillation equipment (9%) is subtracted from these costs the resulting expenditure will be 190.19 million US dollar. Table 21: Shows the result from the Investment Expenditure for the four alternatives including the case from the 2008 report. All costs are in million US dollar. C BM2011 Additional Costs for butanol plant (MMUSD) Additional cost for ethanol plant (MMUSD) Total investment cost (MMUSD) Alternative (MMUSD) Vacuum Distillation 16.6 41.9 190.2 248.7 Oleyl/decane Perstraction 13.7 41.9 190.2 245.8 Mesitylene Perstraction 8.0 41.9 190.2 240.1 HAP Catalysis 18.4 0 209.3 227.7 Report 2008 46.1 0 209.3 255.4 As seen above the alternative with a HAP catalyst is the least expensive alternative according to the investment expenditure and the Report 2008 is the most expensive one. The cost from the report 2008 and from the vacuum distillation is almost the same which is reasonable since in the 2008 report the distillation has been optimized to resemble vacuum distillation. In table 22 a comparison of investment costs divided by the butanol production for all alternatives are described. The alternative with a HAP catalyst has by far the highest investment cost/butanol production while the other alternatives are almost the same. This might seem strange since the investment cost for the HAP alternative was the lowest but the low productivity of the condensation reaction results in it having the highest investment cost per unit mass of butanol produced. 34

Table 22: Shows the investment cost divided with the butanol production for each alternative. Alternative Butanol production (m 3 /year) Total investment cost (MMUSD) Investment cost/ butanol production ($*year/(m 3 )) Vacuum Distillation 37 996 248.7 6 546 Oleyl/decane Perstraction 37 365 245.9 6 580 Mesitylene Perstraction 38 754 240.1 6 197 HAP Catalysis 18 535 227.7 12 283 Report 2008 37 156 255.4 6 873 4.2 Operating Costs for the five Alternatives The operating costs are in contrast to the investment expenditure not one-time costs but rather more continuous. Examples are the cost for holding material in stock, raw material costs and staff salaries. In the calculation some assumptions have been made, for example the production is assumed to be running 360 days/year and the membrane and the catalyst have a lifespan of 2 years. In table 23 the operating cost including all additional costs are shown. In the case from 2008 they have not included some of the additions costs and therefore it has a lower total annual cost. The report from 2008 has not based the costs on the entire butanol plant but rather on the investments needed in addition to the investments for the ethanol plant which leads to a lower total investment cost and therefore some of the operating costs will be lower. The other four alternatives have almost the same operating cost, a little difference can however be seen for Vacuum distillation which is slightly higher due to the need of many and large distillation columns which require a lot of steam. 35

Table 23: Shows the Operating costs in Million USD for the four alternatives and the case presented in the report from 2008. Operating costs Vacuum Distillation Oleyl/Decan Perstraction Mesitylene perstraction HAP Catalysis Report 2008 Wood 21.64 21.64 21.64 21.64 21.64 SO 2 0.40 0.40 0.40 0.40 0.40 Pellets 6.28 0 0 0 6.4 Enzyme 7.14 7.14 7.14 7.14 7.14 Membrane 0 0.50 0.50 0 0 Catalyst 0 0 0 0.0018 0 Electricity 0.02 0.02 0.02 0.02 0.02 Storage 2.56 2.51 2.50 1.55 0 Spare Parts 0.15 0.15 0.14 0.14 0 Maintenance and Reparation 0.99 0.98 0.96 0.91 0.13 Process Operators 4.61 4.61 4.61 4.61 1.2 Shift Management 0.69 0.69 0.69 0.69 0.18 Laboratory Staff 0.69 0.69 0.69 0.69 0.18 Licenses 1.51 1.34 1.33 1.13 0 Overhead for Personals 4.06 4.06 4.06 4.06 0.97 Administration 1.01 1.01 1.01 1.01 0.24 R&D 0.10 0.10 0.10 0.10 0.8 Total Annual Cost 51.86 45.84 45.80 44.25 39.85 Table 24 shows the butanol production; the cost for producing butanol via HAP Catalysis is much higher than producing it via for example Oleyl decane perstraction. This is most likely due to the very low production (half the size of the other alternatives, see table 24). Table 24: Shows specific operating cost as the total annual operating cost divided with the annual butanol production. Alternatives Butanol production (ton/year) Total Annual Costs (MMUSD) Specific Operating Cost (USD/ton) Vacuum Distillation 30 110.4 51.86 1 722.3 Oleyl/Decan Perstraction 29 479.7 45.84 1 555.0 Mesitylene Perstraction 27 587.5 45.84 1 660.7 HAP Catalysis 14 679.4 44.25 3 014.4 Report 2008 29 427.8 39.85 1 354.2 36

4.3 Revenue for the five Alternatives As seen in table 26 Oleyl/Decan perstraction has the highest income from sold pellets, butanol and acetone. HAP as a catalyst gives the by far lowest income which again depends on its low butanol production, as a matter of fact more than half of the incomes are generated from producing and selling pellets. In the calculation following assumptions have been made, see table 25: Table 25: Price assumptions. Butanol Acetone Pellets Electricity Wood Price 1.64 USD/kg 1.16 USD/kg 739 USD/kWh 0.0166 USD/kWh 120.22 USD/ton Table 26: Shows the incomes from the four alternatives and the case from 2008. Alternatives Pellets (MMUSD) Butanol (MMUSD) Acetone (MMUSD) Total (MMUSD) Vacuum Distillation 0 49.27 8.82 58.09 Oleyl/Decan Perstraction 3.65 48.45 10.04 62.14 Mesitylene Perstraction 6.30 45.14 10.16 61.60 HAP Catalysis 15.30 12.11 0 27.41 Report 2008 0 37.37 8.18 45.55 4.4 Production cost for the five Alternatives The production cost for producing butanol in the five different alternatives can be calculated by adding the depreciation to the operating costs while also removing earnings from acetone and pellets sales. The results for the five alternatives can be seen in table 27. In table 27 it is seen that the perstraction methods are the cheapest production alternatives closely followed by the case from the 2008 report. Table 27; Shows the production cost for producing butanol for the five alternatives Alternatives Annual Production Cost (MMUSD) Production cost (USD/ton) Vacuum Distillation 67.91 2 255 Oleyl/Decan Perstraction 56.74 1 925 Mesitylene Perstraction 53.36 1 934 HAP Catalysis 51.72 3 523 Report 2008 57.21 1 944 4.5 Net Present Value for the five Alternatives The net present values describe an alternatives profitability and is calculated according to equation 9 to 11. a i = Annual net payment [Eq. 9] 37

I i = Annual income U i = Annual operating cost X= Interest rate, 15% in this report N= Lifespan, 10 in this report S n = Return N years forward G= Investment expenditure [Eq. 10] [Eq. 11] The results of these calculations are shown in table 28 below where the lifespan is assumed to be ten years and the interest rate used is 15 %. The reason that all values are negative is that in these calculations investment costs for the ethanol plant has been included which is not the case in the report from 2008. Table 28: Shows the calculations on the net present value. Alternative Annual income Annual operating cost Annual net payment Net Present Value Vacuum Distillation 58.09 51.86 6.23-217.46 Oleyl Decan Perstraction 59.10 45.46 13.34-164.04 Mesitylene perstraction 56.35 45.65 10.70-160.85 HAP Catalysis 14.65 48.87-34.22-312.17 Report 2008 50.6 44.23 6.37-223.39 4.6 Pay-Off Time for the five Alternatives An investments pay-off time describes how long time it takes for the future annual payments to pay off the original investment cost if the interest rate is disregarded. The pay-off time will be calculated according to equation 12. n = Pay-off time G = Investment expenditure a i = Annual net payment [Eq. 12] The pay-off times for four of the alternatives are shown in table 29, the method with a HAP catalyst is not relevant since its annual net payment is negative (table 28) As seen in table 29 Oleyl/Decan and Mesitylene perstraction both have a pay-off time around 15 year which is quite good given the largely negative net present values (table 28). 38

Net present value (MMUSD) Table 29: Shows the time it takes to pay-off the investment costs. Alternative Pay-Off time (years) Vacuum Distillation 39.93 Oleyl Decan Perstraction 15.08 Mesitylene perstraction 15.20 HAP Catalysis N/A Report 2008 40.09 4.7 Sensitivity Analysis for the Five Alternatives Three variables will be analyzed in order to see how sensitive the calculations are to changes in the assumptions. These variables are the price of butanol and pellets and the plant lifespan. The effects of these variables on the net present value and the pay-off time have been studied and can be seen in figures 27-31, for more detailed values see Appendix F. The effects of lifespan changes on pay-off time are non-existing which is why the results are not included in this report. The results from the sensitivity analysis clearly shows that the studied parameters have some effect on the net present value and pay-off times for the alternatives although none of the changes (which were quite large) made caused any specific altercations of the final results other than the alternative became more or less profitable than they already were without making any of them profitable or even close to profitable. 0-50 -100-150 -200-250 -300-350 Sensitivity Analysis of change in Butanol Price Vacuum distillation Oleyl decan perstraction Mesi extraction HAP catalysis 1.43 1.64 1.85 Butanol price (USD/kg) Figure 27: Shows the sensitivity analyses with regard to changes in butanol price from 1.43 USD/L to 1.85 USD/L, effects on the net present values are shown. 39

Net present value (MMUSD) Pay-off time (years, log-scale) Sensitivity Analysis of Change in Butanol Price 1000 100 10 Vacuum distillation Oleyl decan perstraction Mesi extraction 1 1.43 1.64 1.85 Butanol price (USD/kg) Figure 28: Shows the sensitivity analysis with regard to changes in butanol prices from 1.43 USD/L to 1.85 USD/L, effects on the pay-off times are shown. Sensitivity Analysis of Change in Pellets Price 0-50 -100-150 -200-250 -300-350 600 739 900 Pellets price (USD/kW) Vacuum distillation Oleyl decan perstraction Mesi extraction HAP-catalysis Figure 29: Shows the sensitivity analysis with regard to changes in pellets prices between 600 USD/kW and 900USD/kW, effects on the net present values are shown. 40

Net present value (MMUSD) Pay.off time (years) 60 50 40 30 20 10 0 Sensitivity Analysis of Change in Pellets Price Vacuum distillation Oleyl decan perstraction Mesi extraction 600 739 900 Pellets price (USD/kW) Figure 30: Shows the sensitivity analysis with regard to changes in pellets prices between 600 USD/kW and 900USD/kW, effects on the pay-off times are shown. 0-50 -100-150 -200-250 -300-350 Sensitivity Analysis of Change in Lifespan Vacuum distillation Oleyl decan perstraction Mesi extraction HAP catalysis Report 2008 10 15 20 Lifespan (years) Figure 31: Shows the sensitivity analyses of changes in lifespan from 10 to 15 years, effects on the net present values are shown. 4.8 Financial Comparison of the five Alternatives It was hard to compare the four alternatives with the case from 2008 since in the case from 2008 the costs for an ethanol plant were not included. They have made the calculations based on an existing ethanol plant; therefore in order to compare the five alternatives with each other, additions had to be made to the cost from the 2008 report. In the cases based on vacuum distillation, Oleyl decane perstraction and Mesitylene perstraction additional costs for ethanol production plant excluding the distillation costs have been taken into account. To the plant producing butanol with a HAP catalyst the entire cost for the ethanol plant has been added, but no costs from the 2008 report. With this in mind the five alternatives are comparable with each other, but also due to this the result in this report differs from the results in the report 2008. On the other hand in this report, the calculated costs are for an entirely new butanol production plant. Since all the alternatives have a negative net present value it means that no alternative is profitable, but of course the alternatives can be internally ranked. The reason to the fact that 41

all alternatives are non-profitable depends on to low incomes, i.e. low butanol production, and to high investment cost. HAP catalysis has the lowest investment costs and operating costs, but it also produced the lowest amounts of butanol, which leads to low incomes and a negative annual net payment. This makes it irrelevant to calculate the pay-off time. Oleyl decane perstraction and mesitylene perstraction are the two best alternatives, closely followed by the vacuum distillation and the case presented in the report from 2008. Mesitylene perstraction and oleyl decane perstraction are the best alternatives since they have the highest incomes and not too high investment costs. The high income depends on the fact that oleyl decane perstraction and mesitylene perstraction are the only alternatives that have incomes from butanol, acetone and pellets. This leads to the fact that they have the shortest pay-off time and are therefore the least bad alternatives. In table 30 the distances to profitability in million USD are shown, these values represent the annual net payment needed to reach profitability. The case with HAP catalysis has the lowest value which depends on the fact that HAP has the lowest investment cost and therefore needs a lower income to cover this cost. Table 30: Shows the distance to profitability in annual net payment. Distance to profitability in Alternative annual net payment (MMUSD) Vacuum Distillation 49.6 Oleyl Decan Perstraction 49.0 Mesitylene perstraction 47.9 HAP Catalysis 45.4 Report 2008 50.9 Despite table 30 Oleyl/Decane and Mesitylene perstraction are the least bad alternatives thanks to their high income and the low pay-off time. HAP catalysis is the worst alternative since it has a negative net annual payment due to the low butanol productivity. 42

5 Recommendations Based on the knowledge as of today it cannot be stated that synthesis of butanol from ethanol is better (with respect to process economy) than fermentation of butanol, the so called ABE process. The following is recommended: 5.1 Synthesis of Butanol from Ethanol The method of indirect synthesis of butanol from ethanol seemed at first to be a promising way to go. However, it turned out that too much byproduct was produced during the condensation reaction, which kept the butanol productivity low. The low butanol production capability of such a plant therefore kept revenues at a minimum and caused a negative net income. Therefore direct synthesis via butanol producing bacteria is the recommended way to use in for future work. The low productivity of indirect synthesis on the other hand leaves room for future process development which is why it might be interesting to also look into the processes optimization opportunities, in this study however the alternative does not look viable. If further investigation was to be conducted in this field interesting areas of research might include looking at possible commercial opportunities for the byproducts, if these were to be further purified, Statoil could sell them as well as the butanol and thus raising the revenue stream significantly. Another interesting topic is any possibility of enhancing the reactions selectivity towards butanol since this would increase the butanol productivity in the plant and in that way increases revenues. 5.2 Separation Processes Out of the investigated separation processes the two perstraction methods are the most efficient, this is mainly due to their lower energy and steam demand compared to vacuum and regular distillation while maintaining about the same productivity. The low steam demand leads to a pellet surplus which thus can be sold to further increase revenues compared to the two pure distillation processes in which pellets has to be bought. Other than this lower demand for steam and energy, no larger differences exist between the four alternative separation processes. Also the two perstraction methods are very much alike even though it seems at this stage that oleyl/decane perstraction is slightly better than mesitylene perstraction. Thus these are the two most interesting processes to look into in future research, other than indirect synthesis. 5.3 Economy From a purely economic aspects none of the alternatives presented in this report are profitable (negative net present values), they are not even close. Though some of the processes might still be interesting to look into, one can see that the large extra costs which are coupled to butanol production compared to ethanol production needs to be lowered and the productivity of butanol must be greatly increased if this investment is ever going to be profitable. All in all, the economic analysis points to the perstraction methods as the two most profitable out of the five studied alternatives, which is why they are most interesting for a future butanol plant. The two perstraction methods have pay-off times of 15.2 years compared to the process using HAP as an catalyst which has a pay-off time of 40.1 years. 43

7 References [1] Larsson, E.; Max-Hansen, M.; Pålsson, A.; Studeny, R. A feasibility study on conversion of an ethanol plant to a butanol plant 2008. [2] http://www.statoil.com/en/about/inbrief/pages/default.aspx, 2011-02-10 [3] http://www.statoil.com/en/about/worldwide/pages/default.aspx, 2011-02-10 [4] http://www.ne.se/lignocellulosa, 2011-02-19 [5] Shamsudin, S.; Sahaid Kalil Hj, M.; Mohtar Wan Yusoff, W. Production of Accetone, Butanol and Ethanol (ABE) by Clostridium saccharoperbutylaccetonicum, Pakistan Journal of Biological Sciences 2006, 9(10), 1923-1928. [6] http://www.vardhandboken.se/texter/infektionerimagtarmkanalen/clostridiumdifficile/ 2011-02-23 [7] Tsuchida, T.; Sakuma, S.; Takeguchi, T.; Ueda, W. Direct Synthesis of n-butanol from Ethanol over Nonstoichiometric Hydroxyapatite. Ind. Eng. Chem. Res. 2006, 45, 8634-8642. [8] Ndou, A.S.; Plint, N.; Coville N.J. Dimerisation of ethanol to butanol over solid-base catalysts. Applied Catalysis A. 2003, 251, 337-345. [9] Yang, C.; Meng, Z. Bimolecular Condensation of Ethanol to 1-Butanol Catalyzed by Alkali Cation Zeolites. Journal of Catalysis. 1993, 142, 37-44. [10] e-book Ullmann 2007. Acetaldehyde, 2011-02-11 [11] Neramittagapong, A.; Attaphaiboon, W.; Neramittagapong, S. Acetaldehyde production from ethanol over Ni-based catalysts. Chiang Mai J. Sci. 2008, 35(1), 171-177. [12] Deng, J.; Cao, Z.; Zhou, B. A study of ethanol dehydrogenation reaction in a palladium membrane reactor. Applied Catalysis A. 1995, 132, 9-20. [13] Weissermel, Klaus and Arpe, Hans-Jürgen (2003), Industrial organic chemistry, 4 th edition, Weinheim, WILEY-VCH verlag GmbH. [14] e-book Ullmann 2007. Crotonaldehyde, 2011-02-11 [15] Campo, B.; Santori G.; Petit, C.; Volpe, M. Liquid phase hydrogenation of crotonaldehyde over Au/CeO 2 catalysts. Applied Catalysis A. 2009, 359, 79 83. [16] Hutchings, G.J.; King, F.; Okoye, I.P.; Rochester, C.H. Influence of sulphur poisoning of copper/alumina catalyst on the selective hydrogenation of crotonaldehyde. Applied Catalysis A, 1992, 83, L7 L13. [17] Qureshi, N.; Blaschek, H.P.; Recovery of butanol from fermentation broth by gas stripping. Renewable Energy, 2001, 22, 557 564. 44

[18] Ezeji, T. C.; Qureshi, N.; Blaschek, H. P. Acetone butanol ethanol(abe) production from concentrated substrate: reduction in substrate inhibition by fed-batch technique and product inhibition by gas stripping, Appl. Microbiol. Biotechnol, 2004, 63, 653 658 [19] Ezeji, T.C.; Karcher, P.M.; Qureshi, N.; Blaschek, H.P. Improving performance of a gas stripping-based recovery system to remove butanol from Clostridium beijerinckii fermentation, Bioprocess Biosyst. Eng. 2005, 27, 207 214. [20] Qureshi, N.; Maddox, I. S. Reduction in butanol inhibition by perstraction: Utilization of Concentrated Lactose/Whey Permeate by Clostridium acetobutylicum to Enhance Butanol Fermentation Economics. Food and Bioproducts Processing, 2005, 83(C1), 43 52. [21] Jeon, Y.J.; Lee, Y.Y. In situ product separation in butanol fermentation by membraneassisted extraction. Enzyme Microb. Technol. 1989, 11, 575 582. [22] Qureshi, N.; Meagher, M.M.; Huang, J.; Hutkins, R.W. Acetone butanol ethanol (ABE) recovery by pervaporation using silicalite silicone composite membrane from fed-batch reactor of Clostridium acetobutylicum. Journal of Membrane Science 2001, 187, 93 102. [23] Oudshoorn, A.; van der Wielen, L. A.M.; Straatho, A.J.J. Adsorption equilibria of biobased butanol solutions using zeolite. Biochemical Engineering Journal 2009, 48, 99 103. [24] Groot, W. J.; van der Lans, R. G. J. M.; Luyben, K. Ch. A. M. Technologies for Butanol Recovery Integrated with Fermentations. Process Biochemistry 1992, 27, 61-75. [25] Roffer, S.; Blanch, H.W.; Wilke, C.R. ExtractiveF ermentatioonf Acetone and Butanol: Process Design and Economic Evaluation. Biotechnology Progress. 1987, 3, 131 140. [26] Tava, M.; Ishii, S.; Kobdvashi, T. J. Ferment. Technol., 1985, 63, 181. [27] Matsumura, M.; Kataoka, H.; Sueki, M.; Araki, K. Bioprocess Eng., 1988, 3, 93 [28] Shi, Z.; Zhang, C.; Chen, J.; Mao, Z. Performance evaluation of acetone butanol continuous flash extractive fermentation process. Bioprocess Biosyst Eng. 2005, 27, 175 183 [29] http://www.osha.gov/dts/osta/otm/otm_iv/otm_iv_2.html, 2011-03-21 [30] http://www.s-ohe.com/ethanol.html, 2011-03-21 [31] http://www.s-ohe.com/butanol.html, 2011-03-21 [32] http://www.s-ohe.com/acetone.html, 2011-03-21 [33] Personal contact, prof. Guido Zacchi, department of chemical engineering, Lund University 45

[34] Roffler, S.; Blanch, H.W.; Wilke, C.R.; Extractive Fermentation of Acetone and Butanol: Process Design and Economic Evaluation. Biotechnoly Process, 1987, September. [35] Kraemer, K.; Harwardt, A.; Bronneberg, R.; Marquardt, W.; Separation of butanol from acetone-butanol-ethanol fermentation by a hybrid extraction-distillation process. 20 th European Symposium on Computer Aided Process Engineering, 2010. [36] Hans T. Karlsson, ProjekteringsHandboken, 2007, Institutionen för Kemiteknik Lunds Tekniska Högskola [37] Hans T. Karlsson, Projektmetodik, 1982, Institutionen för kemiteknik Lunds Tekniska Högskola [38] http://www.scb.se/pages/tableandchart 272151.aspx, 2011-04-24 [29] Per Sassner, Lignocelluloses ethanol production based on steam pretreatment and SSF, 2007, Institutionen för kemiteknik Lunds universitet 46

Appendix A Data for the Calculations Table A1: Shows data for the substances used in the production of butanol [100], [101]. Substance Boiling point ( C) Cp (kj/kg*k) Evaporation enthalpies (kj/kg) CH 4-161.5 2.18 549 C 2 H 4-104.0 1.51 515 C 2 H 6-88.6 1.73 540 CH 3 CHO 21.0 1.62 584 C 3 H 6-47.6 2.00 437 CH 3 CH 2 CH 2 OH 97.0 1.98 695 C 4 H 8-42.3 1.53 447 CH 2 =CH-CH=CH 2-4.4 2.30 1044 C 2 H 5 OC 2 H 5 34.6 1.50 360 n-c 4 H 9 OH 118.0 2.43 590 C 2 H 5 OH 78.5 2.43 841 C 3 H 6 O 56.2 2.20 515 H 2-252.8 14.20 466 C 2 H 4 O 2 118.0 2.20 413 Table A2: Boiling points for acetone, 1-butanol and ethanol [102], Specific heat for acetone, butanol and ethanol [103] Pressure (atm) Boiling point, acetone ( C) Boiling point, 1-butanol ( C) Boiling point, ethanol ( C) Cp, acetone (kj/kg*k) Cp, 1-butanol (kj/kg*k) Cp, ethanol (kj/kg*k) 1 56.5 117.7 78.4 2295 3223 3087 0.9 53 115 76 2282 3183 3050 0.8 50 111 73 2270 3119 3007 0.7 47 108 69 2259 3071 2954 0.6 42 103 65 2241 3000 2896 0.5 38 98 61 2227 2944 2837 0.4 32 93 56 2209 2880 2774 0.3 25 85 49 2187 2784 2689 0.2 16 76 40 2161 2693 2588 0.1 1 60 26 2121 2557 2450 i

Appendix B Simulations in Aspen Plus Aspen Plus was used to simulate the processes at steady-state, from these simulations, the energy demands and the dimensions for the distillation stages were obtained. Distillation in Aspen Plus In the report from 2008 [1] donut trays was recommended (because of solid substance in the feed). Since donut trays not were available in Aspen Plus, so called bubble caps were instead used in the calculations. The efficiency was set to 100%. The model used to describe to equilibrium between steam-liquid-liquid was an activity factor model (NRTL) [33]. The idea is to get as pure butanol as possible, and at the same time not lose too much. The composition of each stream is shown in tables E1-E7 (Appendix E) and the dimension of each distillation column is shown in tables D15-D18 (Appendix D). ii

Appendix C Process Calculations Heat Exchangers All heat exchangers are counter-current and have an estimated heat transfer coefficient to 1,0 kw/(m 2 K). The calculation of heat transfer area was made with equations C1 and C2. [Eq. C1] [Eq. C2] Decanters The design of the decanters is based on a retention time of 12 minutes. In the report from 2008 [2] a length to diameter ratio of 3:1 was used, since there is no reason to change it, it is kept the same. Equations C3-C5 were used. [Eq. C3] [Eq. C4] [Eq. C5] Fermenters/Reactors Perstraction Mesitylene as solvent Se Perstraction Oleyl Alcohol/Decane as solvent above for calculations. The extra fermenter volume will be 39 m 3. Perstraction Oleyl Alcohol/Decane as solvent The extra volume needed in the fermenter is based on a retention time of 10 minutes for the solvent. Equations C3 to C5 are used. The extra fermenter volume will be 62 m 3. iii

Feed Composition Perstraction Oleyl Alcohol/Decan as solvent From the 2008 report [1] data presented in table C1 was taken. Table C1: Table with data for calculations of Oleyl/Decan Perstraction ABE ratio [wt%] 21.3:74.5:4.2 Production rate [kg/(m 3 h)] 0.320 Inhibition concentration [kg/m 3 ] 20.0 The perstraction was modeled as a continuous stirred tank reactor (the fermenter) with a smaller tank placed inside it. The two tanks are communicating through a membrane. Through this membrane acetone, butanol, ethanol and water diffuse. The group has made the assumption that mass transfer through the membrane not is the rate limiting step (it is probably the production rate). Data (for oleyl alcohol) in table C2 is taken from an article [34]. Table C1: Table with data for Oleyl/Decan Perstraction calculations k D,Acetone 0.3 k D,Butanol 2.6 k D,Ethanol 0.1 Solubility water in solvent 50 ppm Solubility solvent in water 100ppm At steady-state the production of each component, in kg/h must be equal to the amount of each component that diffuses throw the membrane, in kg/h. This statement gives three equations, a fourth equation is given by the fact that the inhibition concentration is 20 kg/m 3 (equation system C1). The values in the equations are shown below. [Eq. syst. C1] getting washed, filled or cleaned) (Number of tanks in use, not When solving equation system C1 the following parameter values are obtained. iv

Ideally the ratio between concentration of any component in water and the concentration of the same component in solvent, k D, will tune in extremely fast, this leads to the flow of organic solvent to be the lowest possible. The mass of the water dissolved in the solvent is: Perstraction Mesitylene as solvent To determine the composition in the feed stream, the same calculations as for perstraction using oleyl alcohol/decane were carried out. Data can be seen in table C4. Table C2: Data for the simulation of Mesitylene Perstraction. [35] k D,Acetone 0.83 k D,Butanol 2.2 k D,Ethanol 0.1 Solubility water in solvent 0.00112 Solubility solvent in water - v

Appendix D Process Design Heat Exchangers Indirect synthesis Hydroxyapatite The sizes (and other relevant data) of the heat exchangers in the process for direct synthesis of butanol from ethanol are reported in table D1. Table D1: Heat exchangers specifications for the HAP catalysis alternative Heat exchanger Heat duty [MW] T in [K] (process) T out [K] (process) T in [K] (external) T out [K] (external) Area [m 2 ] HE1 0.27 519 573 623 623 3.6 HE2 2.3 325 380 389 389 84 Perstraction Mesitylene as solvent The size (and other relevant data) of the heat exchangers in the process using perstraction with mesitylene as solvent are shown in table D2. Table D2: Heat exchangers specifications for the mesitylene perstraction alternative Heat Heat duty T in [K] T out [K] T in [K] T out [K] exchanger [MW] (process) (process) (external) (external) Area [m 2 ] HE1 8.3 423 353 293 343 117 Perstraction Oleyl Alcohol/Decane as solvent The sizes (and other relevant data) of the heat exchangers in the process using perstraction with oleyl alcohol/decane as solvent are reported in table D3. Table D3: Heat exchangers specifications for the oleyl alcohol/decane perstraction alternative Heat exchanger Heat duty [MW] T in [K] (process) T out [K] (process) T in [K] (ext./pro.) T out [K] (ext./pro.) Area [m 2 ] HE1 8.6 392 346 310 353 230 HE2 0.11 401 401 358 390 4.5 HE3 0.43 445 313 293 363 9.7 HE4 6.0 346 310 293 336 454 Vacuum distillation The sizes (and other relevant data) of the heat exchangers in the process using vacuum distillation are reported in table D4. Table D4: Heat exchangers specifications for the vacuum distillation alternative Heat exchanger Heat duty [MW] T in [K] (process) T out [K] (process) T in [K] (external) T out [K] (external) Area [m 2 ] HE1 2.4 348 388 420 420 49 HE2 0.44 350 358 365 365 4 vi

Decanters Indirect synthesis Hydroxyapatite The dimensions and operating conditions for the decanter in the process for direct synthesis of butanol from ethanol are reported in table D5. Table D5: Decanters specifications for the HAP catalysis alternative Height [m] Diameter [m] Pressure [bar] Temp. [K] Decanter 1 2.25 0.75 4.0 390 Vacuum Distillation The dimensions and operating conditions for the decanters in the process using vacuum distillation are reported in table D6. Table D6: Decanters specifications for the vacuum distillation alternative Height [m] Diameter [m] Pressure [bar] Temp. [K] Decanter 1 5.0 1.7 3.4 388 Decanter 2 3.0 1.0 1.0 358 Pumps/Compressors Indirect synthesis Hydroxyapatite The in- and outlet pressure for the compressor in the process for indirect synthesis of butanol from ethanol are reported in table D7. Table D7: Compressor specifications for the HAP catalysis alternative Inlet pressure [bar] Outlet pressure [bar] Compressor 1.0 10 Perstraction Mesitylene as Solvent The in- and outlet pressure for the pumps in the process using perstraction with mesitylene as solvent are reported in table D8. Table D8: Pumps specifications for the mesitylene perstraction alternative Inlet pressure [bar] Outlet pressure [bar] Pump 1 0.70 1.0 Pump 2 0.70 1.0 Perstraction Oleyl Alcohol/Decane as Solvent The in- and outlet pressure for the pump in the process using perstraction with oleyl alcohol/decane as solvent are reported in table D9. Table D9: Pumps specifications for the oleyl alcohol/decane perstraction alternative Inlet pressure [bar] Outlet pressure [bar] Pump 1 1.0 5.0 vii

Vacuum Distillation The in- and outlet pressure for the pumps in the process using vacuum distillation are reported in table D10. Table D10: Pumps specifications for the vacuum distillation alternative Inlet pressure [bar] Outlet pressure [bar] Pump 1 1.0 3.4 Fermenters/Reactors In the processes using perstraction or vacuum distillation fermenters are needed, in the case of indirect synthesis of butanol from ethanol a tubular reactor is needed. Indirect Synthesis Hydroxyapatite The design and operating conditions of the reactor in the process for indirect synthesis of butanol from ethanol is reported in table D11. Table D11: Fermenters specifications for the HAP catalysis alternative Pressure Temp. Length Diameter Catalyst [bar] [K] [m] [m] Volume [m 3 ] Mass [kg] Reactor 10 573 1.4 0.47 0.12 385 Perstraction Mesitylene as Solvent The design of the fermenters in the process using perstraction with mesitylene as solvent is reported in table D12. Table D12: Fermenters specifications for the mesitylene perstraction alternative Number Total volume [m 3 ] Effective fermentation volume [m 3 ] Solvent volume [m 3 ] Fermenter 20 1039 800 39 Perstraction Oleyl Alcohol/Decane as Solvent Number Total volume [m 3 ] Effective fermentation volume [m 3 ] Solvent volume [m 3 ] Fermenter 20 1062 800 62 Vacuum Distillation The design of the fermenters in the process using vacuum distillation is the same as in the 2008 report [1]. It is reported in table D14. Table D14: Fermenters specifications for the vacuum distillation alternative Number Total volume [m 3 ] Effective fermentation volume [m 3 ] Fermenter 23 1000 800 viii

Distillation columns Indirect Synthesis Hydroxyapatite The design of the distillation columns in the process for indirect synthesis of butanol from ethanol is reported in table D15. The energy to be cooled away in column 1 s condenser is used in column 5 s boiler. The rest of the energy to be cooled away is used to preheat the ingoing stream. A total of 6.4 MW is to be added to columns 1,2,3 and 4. Table D15: Distillation columns specifications for the HAP catalysis alternative Column 1 2 3 4 5 Heat duty boiler [MW] 3.8 1.6 0.83 0.37 4.0 T Boiler [K] 430 429 391 351 371 T Heating medium [K] 452 452 452 452 380 k Boiler [kw/(m 2 ºC)] 1.0 1.0 1.0 1.0 1.0 Area boiler [m 2 ] 59 69 14 26 224 Heat duty condenser [MW] 6.4 1.8 0.89 0.51 4.0 T Condenser [K] 380 365 365 305 351 T Cooling medium in [K] - 290 290 290 290 T Cooling medium out [K] - 360 360 300 346 k Condenser [kw/(m 2 ºC)] 1.0 1.0 1.0 1.0 1.0 Area Condenser [m 2 ] - 56 34 56 180 Diameter [m] 1.2 1.0 0.70 0.50 1.3 Height [m] 21.6 13.6 13.6 13.6 13.6 Number of trays [ ] 30 20 20 20 20 Reflux [wt/wt] 6.0 0.80 1.0 10 4.0 Pressure (condenser) [bar] 7.0 1.0 1.0 1.0 1.0 ix

Perstraction Mesitylene as Solvent The design in the distillation columns in the process using perstraction with mesitylene as solvent is reported in table D16. The bottom stream from Col-1 contains enough energy to heat Col-2 and Col-3. Therefore energy is only to be added in Col-1 s boiler. A total of 12.6 MW is to be added. Table D16: Distillation columns specifications for the mesitylene perstraction alternative Column Col-1 Col-2 Col-3 Heat duty boiler [MW] 12.6 1.3 0.63 T Boiler [K] 437 390 350 T Heating medium in [K] 485 437 428 T Heating medium out [K] 485 428 423 k Boiler [kw/(m 2 ºC)] 1.0 1.0 1.0 Area boiler [m 2 ] 262 30.5 8.35 Heat duty condenser [MW] 2.4 1.2 0.60 T Condenser [K] 351 335 319 T Cooling medium in [K] 293 293 293 T Cooling medium out [K] 340 325 310 k Condenser [kw/(m 2 ºC)] 1.0 1.0 1.0 Area Condenser [m 2 ] 86 51 37 Diameter [m] 4.2 0.91 0.72 Height [m] 25.6 9.6 13.6 Number of trays [ ] 35 15 20 Reflux [wt/wt] 1.5 2 3.3 Pressure (condenser) [bar] 1.0 1.0 0.70 x

Perstraction Oleyl Alcohol/Decane as Solvent The design in the distillation columns in the process using perstraction with oleyl alcohol/decane as solvent is reported in table D17. The bottom stream from Col-1 contains enough energy to heat Col-2, Col-4, and Col-5 and preheat the stream to Col-3. Energy is to be added to Col-1 and Col-3. A total of 16,2 MW is to be added. Table D17: Distillation columns specifications for the oleyl alcohol/decane perstraction alternative Column Col-1 Col-2 Col-3 Col-4 Col-5 Heat duty boiler [MW] 13.2 3.0 3.0 0.77 0.92 T Boiler [K] 416 388 488 390 348 T Heating medium [K] 507-507 - - k Boiler [kw/(m 2 ºC)] 1.0 1.0 1.0 1.0 1.0 Area boiler [m 2 ] 144.8 152.3 110 98.7 19.7 Heat duty condenser [MW] 13.0 3.1 2.8 0.49 0.92 T Condenser [K] 320 358 445 331 329 T Cooling medium in [K] 293 293 293 293 293 T Cooling medium out [K] 315 348 363 326 324 k Condenser [kw/(m 2 ºC)] 1.0 1.0 1.0 1.0 1.0 Area Condenser [m 2 ] 100 105 24.5 30.1 58.5 Diameter [m] 5.4 1.9 1.1 0.71 0.78 Height [m] 14.4 16 22.4 14.4 16 Number of trays [ ] 20 22 30 20 22 Reflux [wt/wt] 0.3 4 5 1.5 5.5 Pressure (condenser) [bar] 0.28 0.28 5 1 1 Vacuum Distillation The design in the distillation columns in the process using perstraction with mesitylene as solvent is reported in table D18. The energy to be cooled away in HP-S-1 s and HP-S-2 s condensers is the same as the energy to be supplied in LP-S-1 s and LP-S-2 s boilers. The split ratio between the high and low pressure columns is set in such a way so the amount of energy to be cooled away in the condensers is the same as the amount of energy to be supplied in the boilers. The split ratio is (2:1). The energy to be cooled away in LP-R-2 s condenser is greater than the energy to be supplied in LP-R-1 and R-3, therefore this energy is used in the two boilers. The rest of the energy to is cooled away in HE2. A total of 29.6 MW is to be added to columns HP-S-1, HP-S-2 and LP- R-2 xi

Table D18: Distillation columns specifications for the vacuum distillation alternative Column LP-S-1 LP-S-2 HP-S-1 HP-S-2 LP-R-1 LP-R-2 R-3 0.7 Heat duty boiler [MW] 6.6 6.6 11.3 11.3 2.5 4.4 5 T Boiler [K] 330 330 373 373 348 390 338 T Heating medium [K] 360 360 420 420 365 420 365 k Boiler [kw/(m 2 ºC)] 1.0 1.0 1.0 1.0 1.0 1.0 1.0 Area boiler [m 2 ] 369 369 241 241 104 150 29 Heat duty condenser [MW] 6.0 6.0 6.6 6.6 1.8 4.5 0.7 7 T Condenser [K] 342 342 360 360 313 365 330 T Cooling medium in [K] 290 290 330 330 290-290 T Cooling medium out [K] 332 332 330 330 303-315 k Condenser [kw/(m 2 ºC)] 1.0 1.0 1.0 1.0 1.0 1.0 1.0 Area Condenser [m 2 ] 359 359 - - 160-30 Diameter [m] 2.0 2.0 2.0 2.0 1.3 1.7 1.0 Height [m] 13.6 13.6 13.6 13.6 13.6 5.6 Number of trays [ ] 20 20 20 20 20 10 20 Reflux [wt/wt] 7.0 7.0 7.0 7.0 10 4 5 Pressure (condenser) [bar] 0.10 0.10 1.0 1.0 0.50 1.0 1.0 13. 6 xii

Appendix E Process Streams All streams named in the flow sheets for vacuum distillation, perstraction and indirect synthesis of butanol can be found in tables E1 to E7. Indirect synthesis Hydroxyapatite Table E1: Specifications of the flows in the Hydroxyapatite step Mass fraction [wt/wt] Feed Inlet Outlet Gas1 Dest1 Bot1 Dest2 Bot2 Water 0.04 0.05 0.19 0.085 0.10 0.25 0.39 0 Ethanol 0.96 0.95 0.38 0.81 0.89 1.3e-4 0 0 Butanol 0 0 0.22 0 1.7e-4 0.39 0.60 0.025 Hexanol 0 0 0.021 0 0 0.037 0 0.10 2-Etyl- 1Butanol 0 0 0.086 0 0 0.15 0 0.41 Octanol 0 0 3.7e-3 0 0 6.5e-3 0 0.018 2-Etyl- 1Hexanol 0 0 0.095 0 0 0.17 0 0.45 Ethen 0 0 1.0e-4 9.7e-3 1.9e-4 0 0 0 Buten 0 0 5.3e-4 5.6e-3 1.2e-3 0 0 0 Hexen 0 0 1.2e-3 0.028 2.7e-3 0 0 0 Acetaldehyde 0 0 6.4e-4 8.6e-3 1.4e-3 0 0 0 Hydrogen (g) 0 0 9.0e-5 0.036 1.3e-5 0 0 0 Octene 0 0 6.8e-4 0 0 1.2e-3 1.9e-3 0 1.3- Butadiene 0 0 1.6e-3 0.016 3.7e-3 0 0 0 Mass flow [kg/h] Feed Inlet Outlet Gas1 Dest1 Bot1 Dest2 Bot2 Water 217 406 1494 1.6 355 1138 1137 0 Ethanol 4721 7636 3068 15.4 3052 0.6 0.020 0 Butanol 0 0 1787 0 0.60 1787 1746 43 Hexanol 0 0 170 0 0 170 0 170 2-Etyl- 1Butanol 0 0 687 0 0 687 0 687 Octanol 0 0 30 0 0 30 0 30 2-Etyl- 1Hexanol 0 0 766 0 0 766 0 766 Ethene 0 0 0.84 0.8 0.65 0 0 0 Butene 0 0 4.3 0.11 4.2 0 0 0 Hexene 0 0 9.7 0.53 9.1 0 0 0 Acetaldehyde 0 0 5.1 0.16 5.0 0 0 0 Hydrogen (g) 0 0 0.73 0.68 0.045 0 0 0 Octene 0 0 5.5 0 0 6 5,5 0 1,3- Butadiene 0 0 13 0.31 13 0 0 0 Total flow [kg/h] 4938 8042 8041 19 3440 4585 2889 1696 Temperature [K] 349 573 573 389 389 430 365 429 Pressure [bar] 1 10 10 7 7 7 1 1 xiii

Table E2: Specifications of the flows in the Hydroxyapatite step Organ Wat Dest3 Bot3 Gas2 Dest4 Bot4 Dest4 Bot5 Mass fraction [wt/wt] Water 0.18 0.96 0.42 0 6.7e-3 0.041 0.11 0.06 0.97 Ethanol 1.0e-3 5.0e-4 2.4e-3 0 0.14 0.78 0.89 0.94 2.2e-2 Butanol 0.82 0.039 0.57 1 0 0 0 0 3.7e-3 Hexanol 0 0 0 0 0 0 0 0 0 2-Etyl- 1Butanol 0 0 0 0 0 0 0 0 0 Octanol 0 0 0 0 0 0 0 0 0 2-Etyl- 1Hexanol 0 0 0 0 0 0 0 0 0 Ethen 0 0 0 0 0.19 1.9e-3 0 0 0 Buten 0 0 0 0 0.11 0.023 0 0 0 Hexen 0 0 0 0 0.17 0.052 0 0 0 Acetaldehyde 0 0 0 0 0.025 0.029 0 0 0 Hydrogen gas 0 0 0 0 0.026 0 0 0 0 Octene 1.0e-3 4.0e-3 2.7e-3 0 0 0 0 0 0 1.3-Butadien 0 0 0 0 0.33 0.072 0 0 0 Mass flow [kg/h] Water 516 1138 516 0 0.011 7.0 348 188 159 Ethanol 3.0 0.6 3.0 0 0.23 133 0.89 2915 3.6 Butanol 2403 47 706 1700 0 0 2919 0 0.60 Hexanol 0 0 0 0 0 0 0 0 0 2-Etyl- 1Butanol 0 0 0 0 0 0 0 0 0 Octanol 0 0 0 0 0 0 0 0 0 2-Etyl- 1Hexanol 0 0 0 0 0 0 0 0 0 Ethen 0 0 0 0 0.33 0.32 0 0 0 Buten 0 0 0 0 0.20 4.0 0 0 0 Hexen 0 0 0 0 0.30 8.8 0 0 0 Acetaldehyde 0 0 0 0 0.040 4.9 0 0 0 Hydrogen 0 0 0 0 0.045 0 0 0 0 Octene 3.3 5.5 3.3 0 0 0 0 0 0 1.3-Butadien 0 0 0 0 0.56 12 0 0 0 Total flow [kg/h] 2925 1191 1228 1700 1.7 170 3268 3103 163 Temperature [K] 390 390 365 391 305 305 351 351 371 Pressure [bar] 4 4 1 1 1 1 1 1 1 xiv

Perstraction Mesitylene as solvent Table E3: Specifications of the flows in the Mesitylene perstraction step Feed Solvent Top-1 Butanol Top-2 Wa/Et/Bu Acetone Mass frac. [wt/wt] Water 0.0011 0 0.049 0 0.14 0.35 0.0036 Butanol 0.016 0 0.71 1.0 0.14 0.35 0 Acetone 0.0046 0 0.20 0 0.60 0.0046 0.99 Ethanol 0.00093 0 0.040 0 0.12 0.29 0.0048 Mesitylene 0.98 1.0 0 0 0 0 0 Mass flow [kg/h] Water 237 0 237 237 233 3.6 Butanol 3433 10 3423 3193 231 231 0 Acetone 982 0 982 0 982 3.0 979 Ethanol 196 0 196 0 196 191 4.8 Mesitylene 206870 206869 0 0 0 0 0 Total flow [kg/h] 211717 206880 4838 3193 1645 658 987 Temperature [K] 353 437 351 391 335 350 319 Pressure [bar] 1 1 1 1 1 1 1 xv

Perstraction Oleyl Alcohol/Decane as solvent Table E4: Specifications of the flows in the Oleyl alcohol/decane perstraction step Feed BEAWD O-A BD-2 D Butanol D-2 BD Mass fraction [wt/wt] Water 0 0.0026 0 0 0 0 0 0 Butanol 0.013 0.66 0 0.95 0.019 0.996 0.040 0.86 Ethanol 0 0.038 0 0 0 0 0 0 Acetone 0.0036 0.19 0 0 0 0 0 0 Oleyl 0.49 0 0.50 0 0 0 0 0 Decane 0.49 0.10 0.50 0.051 0.98 0.0037 0.96 0.14 Mass flow [kg/h] Water 14 14 0 0.011 0 0.011 0 0.011 Butanol 3438 3438 0 3420 6.7 3412 7.3 3426 Ethanol 196 196 0 1.5 0 1.5 0 1.5 Acetone 985 985 0 0.28 0 0.28 0 0.28 Oleyl 133300 0 133300 0 0 0 0 0 Decane 133300 536 132764 186 350 13 17 536 Total flow [kg/h] 271232 5168 266064 3607 357 3427 180 3964 Temperature [K] 303 320 416 358 388 313 489 391 Pressure [bar] 1 0.28 0.28 0.28 0.28 1 5 1 xvi

Table E5: Specifications of the flows in the Oleyl alcohol/decane perstraction step BAEW Acetone Et/ Bu/Wa Decane Solvent Mass fraction [wt/wt] Water 0.011 0.0040 0.042 0 0 Butanol 0.0010 0 0.051 0.026 0 Ethanol 0.16 0.0060 0.82 0 0 Acetone 0.82 0.99 0.081 0 0 Oleyl 0 0 0 0 0.50 Decane 0 0 0 0.97 0.50 Mass flow [kg/h] Water 14 3.9 9.7 0 0 Butanol 12 0 12 14 29 Ethanol 195 5.8 189 0 0 Acetone 984 966 19 0 0.0019 Oleyl 0 0 0 0 133300 Decane 0.15 0 0.15 523 133287 Total flow [kg/h] 1204 975 229 537 266616 Temperature [K] 331 329 348 425 346 Pressure [bar] 1 1 1 1 1 xvii

Vacuum distillation Table E6: Specifications of the flows in the vacuum distillation step LP- Split HP- Split Feed HPstripper1 Feed HPstripper2 Feed LPstripper1 Feed Mass fraction [wt/wt] Water 0.98 0.97 0.97 0.97 0.97 0.97 0.97 Butanol 0.014 0.016 0.016 0.016 0.016 0.016 0.016 Acetone 0.0040 0.0094 0.0094 0.0094 0.0094 0.0094 0.0094 Ethanol 0.00081 0.0030 0.0030 0.0030 0.0030 0.0030 0.0030 Mass flow [kg/h] Water 240899 92860 186828 93414 93414 46430 46430 Butanol 3487 1497 3012 1506 1506 748 748 Acetone 997 898 1807 903 903 449 449 Ethanol 199 285 573 287 287 142 142 Total flow [kg/h] 245582 95540 192220 96110 96110 47769 47769 Temperature [K] 310 310 310 310 310 310 310 Pressure [bar] 1.0 1.0 1.0 1.0 1.0 0.1 0.1 Table E7: Specifications of the flows in the Oleyl alcohol/decane perstraction step LPbot 1 LP-top 2 LPbot 2 HP-top 1 Feed LPstripper2 HPbot 1 HP-top 2 HPbot 2 LP-top 1 Mass fraction [wt/wt] Water 0.83 1.0 0.83 1.0 0.83 1.0 0.83 1.0 Butanol 0.095 0 0.095 0 0.097 0 0.097 0 Acetone 0.057 0 0.057 0 0.058 0 0.058 0 Ethanol 0.018 0 0.018 0 0.018 0 0.018 0 Mass flow [kg/h] Water 6542 39888 6542 39888 12874 80540 12874 80540 Butanol 748 0 748 0 1506 0 1506 0 Acetone 449 0 449 0 903 0 903 0 Ethanol 142 0 142 0 287 0 287 0 Total flow [kg/h] 7881 39888 7881 39888 15570 80540 15570 80540 Temperature [K] 330 342 330 342 360 373 360 360 Pressure [bar] 0.1 0.1 0.1 0.1 1.0 1.0 1.0 1.0 xviii

Appendix F Detailed Investment Expenditure Tables F1 to F4 describe the investment cost in all the equipment in detail. Vacuum Distillation Table F1: Shows the costs for the equipments in vacuum distillation. Equipment F BM C p (USD) C BM1982 (USD) Decanter 1 10.6 1 500 159 000 Decanter 2 10.6 700 74 200 Pump-1 4.5 3 100 13 950 Pump-3 4.5 1 100 4 950 Heat Excanger-1 4 30 500 122 000 Heat Exchanger-2 4 7 000 28 000 High Pressure Vessel 10.6 50 000 1060 000 Low Pressure Vessel 10.6 50 000 1060 000 High Pressure Trays 2 480 38 016 Low Pressure Trays 2 480 38 016 High Pressure Reboiler 5 36 000 360 000 Low Pressure Reboiler 5 50 000 500 000 Low Pressure Condenser 4.3 70 000 602 000 R-1 Vessel 10.6 31 000 328 600 R-2 Vessel 10.6 20 000 212 000 R-3 Vessel 10.6 28 000 296 800 R-1 Trays 2 320 12 800 R-2 Trays 2 400 12 000 R-3 Trays 2 290 11 600 R-1 Reboiler 5 20 000 100 000 R-2 Reboiler 5 25 000 125 000 R-3 Reboiler 5 10 000 50 000 R-1 Condenser 4.3 55 000 236 500 R-3 Condenser 4.3 25 000 107 500 Total 5 552 932 Perstraction Oleyl Alcohol/Decane as solvent Table F2: Shows the costs for the equipments in oleyl decane perstraction. Equipment F BM C p (USD) C BM1982 (USD) Colum 1 Vessel 10.6 12 000 1 272 000 Colum 2 Vessel 10.6 48 000 508 800 Colum 3 Vessel 10.6 40 000 424 000 Colum 4 Vessel 10.6 27 000 286 200 Colum 5 Vessel 10.6 27 000 286 200 Colum 1 Trays 2 22 000 88 000 Colum 2 Trays 2 460 20 240 Colum 3 Trays 2 300 18 000 Colum 4 Trays 2 250 10 000 Colum 5 Trays 2 260 11 440 Colum 1 Reboiler 2 30 000 60 000 Colum 2 Reboiler 2 31 000 62 000 Colum 3 Reboiler 2 28 000 56 000 Colum 4 Reboiler 2 20 000 40 000 Colum 5 Reboiler 2 8 100 16 200 Colum 1 Condenser 4.3 40 500 174 150 xix

Colum 2 Condenser 4.3 42 000 180 600 Colum 3 Condenser 4.3 23 000 98 900 Colum 4 Condenser 4.3 25 000 107 500 Colum 5 Condenser 4.3 32 000 137 600 Pump-1 4.3 8 000 34 400 Pump-2 4.3 2 000 8 600 Pump-3 4.3 1 100 4 730 Pump-4 4.3 2 800 12 040 Heat Excanger-1 4.3 60 000 258 000 Heat Excanger-2 4.3 7 500 32 250 Heat Excanger-3 4.3 10 300 44 290 Heat Excanger-4 4.3 80 000 344 000 Total 4 596 140 Perstraction Mesitylene as solvent Table F3: Shows the costs for the equipments in Mesitylene Extraction. Equipment F BM C p (USD) C BM1982 (USD) Colum 1 Vessel 10.6 112 000 1 187 200 Colum 2 Vessel 10.6 17 000 180 200 Colum 3 Vessel 10.6 24 000 254 400 Colum 1 Trays 2 15 000 105 000 Colum 2 Trays 2 270 10 125 Colum 3 Trays 2 230 9 200 Colum 1 Reboiler 5 40 000 200 000 Colum 2 Reboiler 5 10 500 52 500 Colum 3 Reboiler 5 6 000 30 000 Colum 1 Condenser 4.3 40 00 172 000 Colum 2 Condenser 4.3 31 000 133 300 Colum 3 Condenser 4.3 28 000 120 400 Pump-1 4.3 1 250 5 375 Pump-2 4.3 1 300 5 590 Heat Excanger-1 4.3 51 000 219 300 Total 2 684 590 Indirect synthesis Hydroxyapatite Table F4: Shows the costs for the equipments in the method for producing butanol based on HAP catalysis. Equipment F BM C p (USD) C BM1982 (USD) Colum 1 Vessel 10.6 41 000 434 600 Colum 2 Vessel 10.6 30 000 318 000 Colum 3 Vessel 10.6 17 000 180 200 Colum 4 Vessel 10.6 15 000 159 000 Colum 5 Vessel 10.6 32 000 339 200 Colum 1 Trays 2 300 18 000 Colum 2 Trays 2 290 11 600 Colum 3 Trays 2 250 10 000 Colum 4 Trays 2 220 8 800 Colum 5 Trays 2 330 13 200 Colum 1 Reboiler 2 12 000 60 000 Colum 2 Reboiler 2 13 000 65 000 Colum 3 Reboiler 2 7 500 37 500 Colum 4 Reboiler 2 9 200 46 000 Colum 5 Reboiler 2 33 000 165 000 xx

Colum 1 Condenser 4.3 72 000 309 600 Colum 2 Condenser 4.3 70 000 301 000 Colum 3 Condenser 4.3 70 000 301 000 Colum 4 Condenser 4.3 69 000 296 700 Colum 5 Condenser 4.3 69 000 296 700 Heat Excanger-1 4.3 40 000 172 000 Heat Excanger-2 4.3 12 000 516 000 Decanter 10.6 5 500 58 300 Compressor 6.3 390 000 2 487 000 Reactor 4.2 900 3 780 Total 6 143 780 xxi

Appendix G - Detailed Sensitivity Analysis A detailed overview of the results from the sensitivity analysis can be seen in table G1. Table G1: Shows the detailed sensitivity analysis values, italic font highlights the values used in the original calculations in this report. xxii