Design and Control of an Alternative Process for Biobutanol

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Process Systems Engineering

Design and Control of an Alternative Process for Biobutanol Purification from ABE Fermentation Devrim B. Kaymak Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b03818 • Publication Date (Web): 08 Jan 2019 Downloaded from http://pubs.acs.org on January 9, 2019

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Industrial & Engineering Chemistry Research

Paper submitted to Industrial & Engineering Chemistry Research

Design and Control of an Alternative Process for Biobutanol Purification from ABE Fermentation

Devrim B. Kaymak*

Department of Chemical Engineering Istanbul Technical University 34469, Maslak, Istanbul, Turkey

*To

whom correspondence should be addressed: [email protected]; Phone: +90-212-285-3539; Fax: +90-212-285-2925

August 29, 2018 Revision: December 21, 2018

ACS Paragon Plus Environment

Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Abstract There is an incremental demand for alternative biofuels because of the decreasing oil supply and increasing environmental concerns. In this respect, biobutanol production from acetone-butanolethanol (ABE) fermentation comes forward as a remarkable process alternative. However, the main drawback of this alternative is the production of dilute fermentation broth which results in very energy demanding separation steps. Thus, several hybrid separation techniques have been studied combining extraction and distillation technologies to develop energy efficient separation processes. Although these techniques solve the well-known drawbacks, they also bring their own disadvantages. In this study, an alternative separation process including reactive distillation columns is proposed to purify biobutanol from acetone-butanol-ethanol (ABE) fermentation. As the result of the reaction taking place in the process, ethylene glycol is obtained as an added-value co-product besides the purification of butanol. The steady-state simulations using Aspen Plus indicate that the total annual cost (TAC) decreases significantly when compared with flowsheets given in the literature. In addition, effective control structures are designed using Aspen Dynamics, and their robustness are tested against two different types of load disturbances. It is seen that stable regulatory control is provided using the proposed control structures.

1. Introduction There is an increasing interest to biofuels because of the problems related to economic and environmental concerns like global warming and climate change. Although bioethanol is currently the most used biofuel, biobutanol has several advantages over bioethanol such as higher energy content, lower volatility and less flammability. In addition, it could be blended with gasoline in any percentage without making any modification in the engine. Because of all these properties,


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biobutanol production based on ABE fermentation is considered as a prominent process alternative1. However, ABE fermentation has an important disadvantage in terms of engineering; a dilute product with a very low concentration is obtained because of the butanol toxicity to microorganisms. Thus, biobutanol production with desired purity requires very energy intensive separation and purification steps. Besides, the presence of azeotropes in the mixture (homogeneous ethanol-water and heterogeneous butanol-water) makes the separation more difficult. For this reason, development of energy efficient separation processes for biobutanol comes into question. Although there are several techniques suggested in the literature such as adsorption, gas stripping, liquid-liquid extraction, pervaporation, perstraction and reverse osmosis, most of the recent studies focus on hybrid separation technologies2-4. In case of large scale production of biobutanol, processes consisting of extraction and distillation are the most widely used methods among these hybrid technologies. A hybrid extraction-distillation process using mesitylene as the solvent is suggested by Kraemer et al.5. Another study suggested by van der Merwe et al.6 includes centrifugation, liquid-liquid extraction and steam stripping distillation columns where the extractant agent is 2-ethyl-1-hexanol. Sanchez-Ramirez et al.7 compared different process flowsheets where one of them is consisted of steam stripping distillation columns following a liquid−liquid extraction column. Hexyl acetate is used as extractant agent in this case. Results of these studies showed that the total annualized cost of the suggested processes is considerably lower than the cost of pure distillation processes. Errico et al.8 compared thermally coupled or thermodynamic equivalent configurations of these hybrid processes, and suggested that the flowsheet including two thermal couplings improves the



objective functions of the problem which are total annual cost and eco-indicator. However, this hybrid extraction-distillation process still presents relatively high energy cost, since the large content of water should be removed from the process and the solvent used in the system should be recovered. At this point, combination of in situ product recovery techniques and distillation based downstream processes to purify the products could significantly reduce the energy costs. For this reason, several in situ product recovery techniques are proposed in the literature where the products are removed during fermentation9. Besides the well-known single stage in situ product recovery methods such as gas stripping, liquid-liquid extraction, adsorption and pervaporation, new two-stage in situ product recovery methods start to appear in the literature such as two-stage gas stripping and gas stripping-pervaporation10,11. Thus, recently, two versions of a new downstream purification process using the effluent stream of an in situ gas stripping process as the feed are proposed by Patraşcu et al.12. These processes are based on the combination of a decanter and a distillation sequence and their flowsheets are illustrated in Figure 1. For the base case process, decanter is the first unit of the process where butanol rich organic phase is separated from the aqueous phase rich in water. Organic phase is fed to the first stripping column (C1) to purify butanol from the bottoms, while the distillate, which is rich in water, is recycled to the decanter. To remove a large amount of water from the system, the aqueous phase is sent into a second stripping column (C2). Top stream of this column is fed to a distillation column (C3). The distillate stream of this column includes an acetone-ethanol rich mixture, while the bottoms stream including mainly butanol-water mixture is recycled back to the decanter. Finally, the ethanol-acetone mixture of the distillate stream is separated in the fourth


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distillation column (C4). For the enhanced heat-integrated process, C2 and C3 are replaced by a dividing-wall column. In Aspen Plus, a Petlyuk distillation setup including a prefractinator (PF) and a main column (DWC) represents the dividing-wall column. The design parameters of that configuration given in the literature are reported in Table 1. Because the recycle streams including azeotropic mixtures are combined with fresh feed stream and fed to the process, a large amount of liquid, mainly water, cycles in the process, and this results in an increase in the equipment sizing (diameter) and energy requirements. The aim of this study is to design and control an alternative downstream biobutanol purification process using reactive distillation columns to attain economically efficient biobutanol from fermentation broth. Ethylene oxide-water reaction takes place in these columns to produce ethylene glycol. Thus, the consumption of water in the reaction results in a serious reduction in the amount of water which should be removed from the process or recycled in the system. In addition, removal of water by the reaction breaks the azeotropes, which helps biobutanol to be purified without any need for separation agents and recycle streams. Besides the main products of the process, an added-value co-product, high purity ethylene glycol, is also obtained in the process.

2. Process Studied Figure 2 shows the flowsheet of the alternative separation sequence for the ABE process studied here. Feed stream of this process is the effluent stream of an in situ gas stripping process used also in Patraşcu et al.12. Its flowrate is 26945 kg/h which corresponds to a plant capacity of 40 kt/yr butanol, and its composition properties are given in Table 2. The first separation unit is a decanter which operates at 40oC. The organic phase which is rich in butanol is fed to the first reactive



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distillation column (C1). This column has a second feed stream including pure ethylene oxide, which reacts in the column with water to produce ethylene glycol. Ethylene glycol is taken as a product from the bottoms, while ABE mixture is removed from the top of the column. Design specification target for this column is 99.5 mole % purity of ethylene glycol in the bottoms stream, and it can be achieved by varying the reflux ratio. Meanwhile, the bottoms flowrate is kept constant to satisfy the recovery of acetone, butanol and ethanol. At the same time, the aqueous phase which is rich in water is fed into a stripping column (C2). This column aims to separate a very big fraction of water as bottoms product. Distillate flowrate is adjusted to satisfy the design specification of water (99.95 mole % purity) in the bottoms stream. Top stream of this column is sent to the second reactive distillation column (C3) together with ethylene oxide, where the rest of the water is consumed to produce ethylene glycol. The design specification is obtaining ethylene glycol product from the bottoms with 99.7 mole % purity by adjusting the reflux ratio. ABE mixture leaving the column from the top merges with the top stream of C1. The combined stream is fed into a distillation column (C4) which purifies butanol as bottoms product with a 99.5 mole % purity by varying the reflux ratio to achieve this design specification. Finally, acetone-ethanol rich top stream is sent into a distillation column (C5) to separate acetone with a 99.5 mole % purity as the top product. The bottoms product is ethanol with a purity bigger than 98 mole %. High purity ethylene glycol is produced as the result of irreversible water-ethylene oxide reaction. The kinetic parameters for the uncatalyzed reaction is obtained from Tavan and Hosseini13, C2H4O + H2O → C2H6O2

(1)

r(kmol m-3 s-1) = 3.15 x 1012 exp[-9547/T] XEO XWater

(2)


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Besides this reaction, an ethoxylation reaction, where ethanol and ethylene oxide react, can take place in the process. However, the effect of this reaction is neglected, because the kinetic experiments conducted by An et al. show that this reaction takes place at much lower rates compared to the hydration rate of ethylene oxide, if there is no catalyst14. The commercial simulation program Aspen Plus is used to perform the steady-state simulation of the process. NRTL method, which is a suitable thermodynamic model for the components and process conditions, is selected. Component databank of the simulator includes the binary interaction parameters for this thermodynamic model. Reactive distillation columns are designed using the rigorous RadFrac model in Aspen Plus. The conventional distillation columns are designed starting with DSTWU shortcut model, and then using rigorous RadFrac model.

3. Economic Basis The objective function of economic evaluation is the total annual cost (TAC). It involves the capital and energy costs of the process, and assumes a payback period (βpay) for capital investment. TAC is given by:

TAC  Energy Cost 

Capital Investment

 pay

(3)

The parameters used in the equipment sizing and cost calculations are summarized in Table 315. Capital costs of individual equipment are calculated using the following equations: CShell ($) = (M&S/280) . (957.9 . D1.066 . H0.82) . (2.18 + (Fm . Fp))

(4)

CTray ($) = (M&S/280) . (97.2 . D1.55) . (NT – 2) . (Ft + Fm)

(5)



CHEX ($) = (M&S/280) . (474.7 . A0.65) . (2.29 + Fm (Fd + Fp))

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(6)

In these equations, Fi represents different cost factors related to material (Fm), pressure (Fp), tray type (Ft) and design type for heat exchanger (Fd). Distillation column diameter (D) is obtained from tray sizing feature of Aspen Plus, while column height (H) is calculated as H = (1.2) . (0.61) . (NT – 2). It should be noticed that the number of trays, NT, includes condenser and reboiler.

4. Results and Discussion 4.1. Steady-State Design First, the effects of reflux ratio and number of reactive trays on the conversion of water and purity of ethylene glycol in bottoms stream have been studied to determine a feasible design of the reactive distillation columns. Several simulations have been carried out by changing the reflux ratio of the first reactive distillation column (C1) operating at 1 atm pressure. For this simulation, the flowrate of bottoms stream, number of reactive and stripping trays, and feed tray locations are kept constant at 59.92 kmol/h, 33, 4 and 29, respectively. Both the target conversion of water and the desired purity of ethylene glycol in the bottoms have been established as 99.5 mole %. The results in Figure 3 show that the conversion of water has a maximum of ~0.97 mole % at a reflux ratio of 2. Since this maximum value of conversion is far from the target conversion, a big amount of reactants, which are water and ethylene oxide, leaves the system from the distillate stream. Thus, a part of the butanol moves down the column and leaves the system from the bottoms, which results in a decrease in the purity of ethylene glycol. Next, the number of reactive trays are increased and several simulations have been performed on the first reactive distillation column to observe its effect on the conversion of water and purity of


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ethylene glycol in the bottoms stream. In this case, the flowrate of bottoms stream, number of stripping trays, and molar reflux ratio are fixed at 59.92 kmol/h, 4 and 2, respectively, while the column operates at 1 atm. The results in Figure 4 show that the increase in the number of reactive trays increases both the conversion and purity. However, the magnitude of changes in conversion and purity is smaller than the changes in the number of reactive trays which brings a big capital cost to the system. It is easily seen that even with 59 reactive trays, the conversion is far from the 99.5 mole% conversion. These results indicate that reactive distillation columns operating at 1 atm cannot totally convert the reactants. As given in Figure 5, the low concentration of reactants in the reactive distillation column directly effects the reaction rate and thus the conversion. The purities of butanol, acetone and ethanol, which are purified in the following steps of the distillation sequence, are also affected besides the low purity of ethylene glycol removed from the bottoms of reactive distillation columns, because of the unconverted reactants leaving the reactive distillation columns from the distillate stream. Reactive distillation columns operating at higher pressures could help to overcome this problem and obtain higher conversion. For that reason, simulations at different operating pressures have been conducted, and the results showing the effect of pressure on the conversion of water and purity of ethylene glycol in the bottoms are illustrated in Figure 6. Flowrate of bottoms stream, molar reflux ratio, number of reactive and stripping trays, and feed tray locations are fixed at 59.92 kmol/h, 2, 33, 4 and 29, respectively. It is seen that the increase in the pressure from 1 atm to 2 atm dramatically improves the conversion, and the conversion of reactants exceeds 99.5 mole% at the operating pressure of 3 atm. As a result of the high conversion, the amount of reactants running away from the distillate



decreases significantly, and ethylene glycol with more than 99.5 mole% purity is obtained. However, no significant improvement is observed in the conversion for operating pressures bigger than 3 atm. Thus, 3 atm is selected as operating pressure for reactive distillation columns. Table 4 gives the design parameters of the novel process. The results show that there is a dramatic decrease in the reboiler heat duty of the novel process compared with the recent configuration proposed by Patraşcu et al.12. The improvement in the total reboiler heat duty (sum of the columns) is 49% when compared with the base-case ABE downstream process, and 33% when compared with the enhanced ABE downstream heat-integrated process including a dividing-wall column (DWC). On the other hand, the novel process includes five distillation columns, which is one column more than the base-case, and two column more than the enhanced ABE downstream heatintegrated process with DWC. Although the number of columns increases in the novel process, the diameters of columns shrink significantly because of the decrease in the molar holdups passing through the process as the result of the consumption of water in the reactive columns. A detailed economic evaluation is provided in the following section. 4.2. Economic Evaluation Table 5 summarizes the capital and operating costs of the novel process. The capital cost considers the decanter, distillation columns, reboilers and condensers, while the operating cost includes heating and cooling utilities in reboilers and condensers. The total annual cost (TAC) of the process is compared with those of two design alternatives given in Patraşcu et. al.12. All the capital and operating costs of these alternatives are recalculated using the economic basis given in our study based on the data given in their study except condenser and reboiler related costs. Since necessary data (condenser heat duty etc.) for cost calculations are not available in the mentioned paper, their


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own capital and operating cost calculations for these equipment are taken into account in the comparison. Based on the recalculation, TAC of the base-case process proposed by Patraşcu et. al. is found 3,754.00 x103$/year, and TAC of the enhanced heat-integrated process is 3,219.33 x103$/year. On the other hand, TAC obtained in the present study is 2,332.70 x103$/year. These results show that there is a 38 % reduction in TAC compared to the base-case configuration, while the decrease in TAC is 27.5 % against the enhanced heat-integrated configuration. 4.3. Process Control Two control structures have been developed using pressure driven method of Aspen Dynamics. The first one includes direct composition control loops, while the second one consists of inferential temperature control. Slope criterion is used for selecting the trays for temperature control. 1 min and 3 min dead times are added to temperature and composition control loops, respectively, to represent the dynamic lags. Closed loop ATV method is implemented to tune these control loops. Tyreus-Luyben settings are used to calculate the controller parameters. On the other hand, all level controllers are proportional-only with KC = 2. Proportional-integral type controllers are used for pressure control with KC = 20 and τI = 12 min. At steady-state conditions, all valves are designed half open. A table including calculated controller tuning parameters are given as supplementary material (Table S1). For both control structures, the production rate handle is the feed stream, and it is flow controlled. The flowrates of related effluent streams are adjusted to control the levels of both organic and aqueous phases. For all distillation columns, levels of reflux drum and column base are controlled by manipulating the distillate and bottoms stream flowrates, respectively. Except C2, reflux flowrates are rationed to the distillate flowrates. In addition, fresh ethylene oxide flowrates in both



reactive distillation columns are rationed to the second feed streams including dilute ABE solution. The ratio is set by a water composition controller on the feed tray. For the first control structure (CS1) shown in Figure 7, the related reboiler heat duties of distillation columns are manipulated to control the ethylene glycol purities in C1 and C3, water purity in C2, butanol purity in C4 and acetone purity in C5. Figure 8A gives the results of CS1 against ±20% change in production rate handle. Solid and dashed lines represent the response to the positive and negative changes, respectively. It is observed that compositions of ethylene glycol in reactive distillation columns and butanol in C4 are easily recovered back to their set points in less than 10 hours. Their transient deviations are relatively small and do not exceed 0.7 mole% in any product composition. Figure 8B shows the results of feed composition change. At time t = 0.5 h, the initial feed composition of butanol (18.5 wt%) is changed by ±2 wt%, while the mass composition of other components are also increased or decreased proportionally to their concentrations. Results for the increase in feed composition of butanol are given by solid line, while the decrease is represented by dashed line. Results illustrate that ethylene glycol and butanol settle down into their specifications in around 10 h without any big oscillation. In this case, the maximum transient deviation of butanol is less than 0.1 mole%. In the case of butanol composition increase, composition of water decreases in the feed stream. As a result, the composition of water in organic phase also decreases. However, there is an increase in the water composition of aqueous phase. Thus, the purity of ethylene glycol in C1 decreases at the beginning, and reboiler heat duty must be increased to bring it back into the desired value, while the opposite takes place in C3.


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For the second control structure (CS2) given in Figure 9, corresponding reboiler heat duties are adjusted to control Tray 37 of C1, Tray 8 of C2, Tray 28 of C3, Tray 12 of C4 and Tray 41 of C5, instead of direct composition control. Other controllers are kept same with those of CS1. Figure 10A shows the results of CS2 against ±20% change in production rate handle. The dashed lines are used to illustrate the results of 20% decrease, while the solid lines represent the results of 20% increase in the throughput. It is seen that the temperature transient deviations for reactive distillation columns are relatively bigger than that of butanol purification column, but the temperatures of controlled trays are recovered back to their set points in less than 10 h for all distillation columns. On the other hand, there are small offset for some product purities. The maximum transient deviations of ethylene glycol purities in reactive distillation columns do not exceed 0.2 mole%. Although ethylene glycol purity in C3 settles down into its set point, there are offset in ethylene glycol purity of C1. However, this offset is less than 0.01 mole%. The maximum transient deviation of butanol purity in C4 is around 0.5 mole %, and it settles down very close to its set point with an offset smaller than 0.1 mole %. As seen in Figure 10B, controlled temperatures return to their set points easily in the case of feed composition changes, similar to the results of disturbance in the production rate handle. On the other hand, there are small offset in the product purities. Maximum transient deviations of ethylene glycol purities in the first (C1) and second (C3) reactive distillation columns are smaller than 0.1 mole%. Similarly, offset of ethylene glycol purities in the first (C1) and second (C3) reactive distillation columns are smaller than 0.001 and 0.01 mole%, respectively. On the other hand, offset of butanol purity in C4 is bigger compared to that of ethylene glycol, but it does not exceed 0.5 mole% in the case of feed composition change in any direction.



5. Conclusions A new separation process is proposed for biobutanol purification from ABE fermentation. Distinctly from the processes widely studied in the literature, this process does not include any extraction unit. Instead, it includes reactive distillation columns to consume water recycling in the process. Thus, butanol-water and ethanol-water azeotropes are broken, and butanol is purified using a conventional distillation column. In addition, ethylene glycol, which is an added-value coproduct for the whole process besides butanol, is obtained as the result of the reaction taking place in reactive distillation columns. Results of the simulation show that the proposed process significantly decreases the TAC (38 % against base-case and 27.5 % against enhanced heatintegrated processes given in the literature). In addition, dynamic behaviour and controllability of the process is studied by designing multi-loop SISO control structures. It is illustrated that the proposed control schemes provides robust control against disturbances in feed flowrate and composition. Associated Content The Supporting Information is available free of charge on the ACS Publications website. Detailed information for the controller tuning parameters. Acknowledgement This research was financially supported by Research Fund of the Istanbul Technical University (Project Number: MGA-2018-41156).


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Table 1: Design parameters of the decanter-distillation process in the literature12 Number of stages Feed stage locations Pressure (atm) Reflux ratio (molar) Diameter (m) Reboiler duty (kW) Total duty (kW)

C1 37 1 1 1.3 2527 11396

C2 8 1 1 1.3 5459

C3 24 19 1 14.18 1.4 1001

C4 30 25 1 12.56 1.3 2409

Table 2: Properties of feed stream Components Acetone Butanol Ethanol Water

Mass fraction 0.045 0.185 0.010 0.760

Table 3: Basis of economics Reboiler Heat transfer coefficient (kW.K-1.m-2) Temperature difference (K) Condenser Heat transfer coefficient (kW.K-1.m-2) Temperature difference (K) Energy costs LP Steam ($/GJ) MP Steam ($/GJ) HP Steam ($/GJ) Cooling Water ($/GJ) Payback period (year) M&S Index (2012)

0.568 34.8 0.852 13.9 7.78 8.22 9.88 0.72 3 1536.5



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Table 4: Design parameters of the novel process C1 Number of stages 39 Feed stage locations 29-29 Pressure (atm) 3 Reflux ratio (molar) 0.89 Diameter (m) 0.66 Reboiler duty (kW) 344 Total duty (kW) 5848

C2 10 2 1 1.22 3399

C3 31 21-21 3 6.69 1.23 1057

C4 20 11 1 1.58 0.61 380

C5 47 33 2.88 0.69 668

Table 5: Summary of economic evaluation for the novel process [103$]

Shell Trays [103$] Reboiler [103$] Condenser [103$] Heating [103$/year] Cooling [103$/year] TAC [103$/year]

C1 151.79 10.41 60.71 142.00 98.37 20.42 240.43

C2 85.16 5.80 269.12 206.42 765.48 36.31 990.62

C3 240.95 21.23 125.93 302.53 302.20 65.38 597.80

C4 78.67 4.51 64.77 102.64 85.56 12.39 181.49


C5 Dec. 185.85 70.06 13.53 93.48 110.92 150.47 13.96 299.03 23.35

Total 812.46 55.49 614.02 864.51 1,402.07 148.47 2,332.70

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Figure Caption Figure 1 – Flowsheet of the decanter-distillation process in the literature12: (A) Base case (B) Enhanced heat-integrated Figure 2 – Flowsheet of the alternative separation process Figure 3 – Effect of reflux ratio on conversion of water and purity of ethylene glycol Figure 4 – Effect of reactive tray numbers on conversion of water and purity of ethylene glycol Figure 5 – Composition profile in reactive distillation column (C1) Figure 6 – Effect of operating pressure on conversion of water and purity of ethylene glycol Figure 7 – Control structure 1 (CS1) Figure 8 – Results of CS1: (A) ±20% throughput change, (B) ±2 wt% feed composition change Figure 9 – Control structure 2 (CS2) Figure 10 – Results of CS2: (A) ±20% throughput change, (B) ±2 wt% feed composition change


Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13Feed from 14Fermentation 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 20 of 30

3

TAC [10 $/year] 3,754.00 2,332.70 C1 Blower

Decanter

Decanter Feed from Ethylene Fermentation Oxide

Ethylene Glycol C3

C1 C4 Butanol

Acetone

C3

C2 Ethanol

C2

C5

Water


Ethylene Oxide Water

Acetone

C4 Ethanol Ethylene Glycol

Butanol

Page 21 of 301 Figure 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41

(A) Feed from Fermentation


(B)

Blower Decanter

Feed from Fermentation

C1 C4 Butanol

Blower Decanter

HEX1

C1 C4

Acetone

C3

Butanol HEX2

Ethanol

C2

Acetone

DWC

Water Water


PF

HEX3

Ethanol

Figure 2 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41


Page 22 of 30

C1

Feed from Fermentation

Decanter

Ethylene Oxide

Ethylene Glycol

C5

Acetone

C4 C3 Ethanol C2

Ethylene Oxide

Ethylene Glycol

Water


Butanol

Page 23 of 303 Figure

0.98

0.98

0.97

0.96

0.96

0.94

0.95 0.92 0.94 0.9

0.93 0.92

0.88

0.91

0.86 1

2

3

4

RR


5

6

Purity of Ethylene Glycol

Conversion of Water

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41


Figure 4

Page 24 of 30

0.985

0.985

0.98

0.98

0.975

0.975

0.97

0.97

0.965

0.965

0.96

0.96

0.955

0.955

0.95

0.95

20

25

30

35

40

45

NT


50

55

60


Conversion of Water

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41


Page 25 of 305 Figure 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41



Figure 6

Page 26 of 30

1

0.995

0.995

0.99

0.99

0.985

0.985

0.98

0.98

0.975

0.975

0.97

0.97

0.965

0.965 1

1.5

2

2.5

3

3.5

P (atm)


4

4.5

5


1

Conversion of Water

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41


Page 27 of 307 Figure 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41


LC FC

LC

FT

C1

ratio

C5 FT

FT

FC

ratio

LC

FC

FT

FC

ratio

FT

LC

C4

Dec LC

CC

CC FT

FT

FC

ratio

LC

LC FC

ratio LC

FT

CC

C3

LC FT

FT

FC

ratio

FT

C2 LC

LC

FT

CC

CC

CC


CC

LC

Figure 8 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41



Page 28 of 30

Page 29 of 309 Figure 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41


LC LC

FT

FC

C1

ratio

CC

C5 FT

FT

FC

ratio

LC

FC LC

TC

FT

TC

C4

Dec LC

FT

FT

FC

ratio

LC

TC

LC FC

ratio LC

FT

C3

LC

CC

FT

FT

FC

ratio

TC

FT

C2 LC

TC

LC


LC

FT

FT

FC

ratio

Figure 10 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41



Page 30 of 30

Design and Control of an Alternative Process for Biobutanol

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