Technoeconomic Study of Biobutanol AB Production. 2. Process Design

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Technoeconomic study of AB biobutanol production: Part 2: Process design Santiago Malmierca, Rebeca Díez-Antolínez, Ana Isabel Paniagua, and Mariano Martín Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.6b02944 • Publication Date (Web): 19 Jan 2017 Downloaded from http://pubs.acs.org on January 23, 2017

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Industrial & Engineering Chemistry Research

Technoeconomic study of biobutanol (AB) production: Part 2: Process design

Santiago Malmiercaa,b, Rebeca Díez-Antolínezb, Ana Isabel Paniaguab, Mariano Martína1 aDepartment

of Chemical Engineering. University of Salamanca. Plz. Caidos 1.5, 37008

Spain bCenter of Biofuels and Bioproducts. Instituto Tecnológico Agrario de Castilla y León (ITACyL), Villarejo de Órbigo, León, Spain.

Abstract In this paper, a process for the production of biobutanol from corn stover is synthesized including the use of novel technologies such as pervaporation and AB fermentation to improve the yield from biomass and reducing the by-products. The processes consists of four sections, pretreatment, hydrolysis, fermentationpervaporation system and product purification. We use the experimental results in Part I of this series to determine the yield of the pretreatment and hydrolysis stages. A hybrid simulation using black boxed kind of models in EXCEL, based on experimental data, and ChemCAD, has been performed. Using the results a detail economic evaluation of the process has been carried out to evaluate the technoeconomic feasibility of this process. The results show an investment of 186 M€ for 27 kt/y of butanol and a butanol production cost of 1.09 €/kg which makes second generation butanol more competitive versus second generation ethanol and fossil based fuels.

Key Words: Biobutanol, Corn stover, Process simulation, Economic analysis

1

Corresponding autor: [email protected]

ACS Paragon Plus Environment


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1. Introduction Over the last 15 years there has been an important effort towards climate change mitigation. Among the different actions, one of the more challenging is reducing emissions in the transport sector. Biofuels offer the potential to provide 27% of the world transportation fuel by 2050.1 Among biofuels, ethanol and biodiesel from different biomass sources represent the current commercial alternatives to substitute crude based gasoline and diesel. However, there are other options such as FT fuels,2 glycerol ethers3 or biobutanol.4 Biobutanol is considered an advanced fuel. It has similar properties to gasoline, it can be directly used in gasoline engines without modification, it has higher energy density and higher boiling point compared to ethanol and, thus, reduced losses in evaporation. Moreover, it is less hydroscopic than ethanol.4 Furthermore, it can also be produced from agricultural residues. However, so far the yield has been low due to the fact that, together with biobutanol, ethanol and acetone are obtained. This type of fermentation is known as ABE, standing for Acetone, Butanol and Ethanol.5 Studies by Mariano et al.6, and Qureshi N. et al.7 have developed economic evaluations of ABE conversion process presenting the challenging economics of such biomass use. In Part 1 of this work, we have worked on the experimental evaluation of biomass processing steps to produce sugars. In the fuels industry, the profit margins are tight. As a result, for a new product to enter the market its economics must be attractive. The production of ethanol as in ABE fermentations represents a yield loss and a challenge for the product separation resulting in economic burdens to the production of biobutanol. Therefore, By using experimental data from the literature on acetone-butanol (AB) fermentations, that which does not produce ethanol,8 together with our results on biomass pretreatment and pervaporation, we propose a flowsheet for the production of biobutanol from lignocellulosic biomass. Thus, in this second part, we present a techno-economic study for the production process of biobutanol from biomass following an AB fermentation which increases the yield to biobutanol and reducing the required investment aiming at generating alternative and competitive biofuels. The rest of the paper is organized as follows. Section two describes the flowsheet. Section three presents the modeling issues for each of the stages to perform the material and energy balances by combining rigorous simulation using CHEMCAD and input – output models based on experimental results. Section four shows the summary of the results of the process simulation reporting the needs of raw


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materials and utilities. Next, section five shows the estimation of the production and investment costs. Finally, section six draws some conclusions.

2. Flow diagram and overall process description Corn stover is the raw material selected for the production of butanol and acetone based on its higher yield to sugars as presented in part 1. The process starts with milling up the biomass to obtain small pieces, from 1 to 2 mm of diameter, for an optimal pretreatment performance. Next, the biomass is mixed with a solution of 0.5% sulfuric acid and water in an autoclaving vessel with 10% w/w of biomass. In the autoclave, the mixture is heated up to 121 ºC and pressurized up to 2.1 bar. Under these conditions, the structure of the hemicellulose is totally broken down into sugars, mainly pentoses such as xylose or arabinose. Even though cellulose and lignin suffer little deterioration, glucose and phenolic compounds appear with the juice. As a consequence of the moderate operating conditions, others compounds including acetic acid and HMF also appear in the liquid stream. After the pretreatment, the solid biomass is neutralized and separated from the liquid phase using a solids washer. The biomass is sent to enzymatic hydrolysis, where enzymes chop-off the long chains of cellulose into glucose. The operating conditions are 50 ºC and atmospheric pressure. Two enzymes kindly provided by Novozymes A/S were selected for scaling up the process, Celullase (N50013) and Glucosidase (N50010). Beside enzymes, it is necessary to add a buffer solution of sodium citrate to maintain the pH at 4.8, required by the enzymes as medium where sugars are released. Similar to the pretreatment stage, the stream that exits the hydrolysis reactor is separated into solid and liquid phases. The first one corresponds to the biomass (lignin) that cannot be digested by the enzyme. This fraction could be burned to obtain energy for the process. The second stream is the sugar juice, which is mixed with the juice coming from pretreatment. The recovered sugar juice containing C5 and C6 sugars is now fed to a nanofiltration unit, where sugars are concentrated up to 72 g/l for subsequent fermentation. Before fermentation, the main stream must be neutralized. Calcium hydroxide is added to neutralize the sulfuric acid remaining after the pretreatment. In this case, laccase9 is not needed because the phenolic compounds do not reach enough



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concentration to inhibit the fermentation. Finally, the stream is cooled down to 34 ºC, the optimal temperature for anaerobic Clostridium spp. Bacteria.8 . The AB fermentation coupled with pervaporation is the novel step proposed in this work. It is a promising alternative to reduce energy consumption in the separation stage and to increase the fermentation yield to butanol. The pervaporation step is a limiting factor within the process due to the stage of development of this technology. In spite of the high yields obtained at laboratory scale,10 there is no industrial scale facility. Therefore, further research is needed to achieve industrial exploitation and validate the results. In our process, we consider this option to evaluate the economic feasibility of the production of butanol. Hyper-producing mutant strain of Clostridium beijerinckii BA101 is selected to transform sugars into organic solvents. It needs a strict anaerobic environment; a temperature of 34 ºC and a low circulation rate of fermentative broth, because it could affect the stability of the C.beijerinckii culture. Under these conditions, the strain is able to consume almost 100% of the sugars. Acetic and butyric acid are also transformed at solventogenesis step.8 Note that C.beijerinckii strain has been genetically altered to improve the production of butanol and acetone with low or negligible production of ethanol. Therefore, the fermentation is assumed to be AB type. The pervaporation step is also responsible for the high production rates. It removes organic solvents, butanol and acetone, from the fermentative broth, which are inhibitors for the bacteria. This process works close to vacuum (10 mbar) and at a temperature of 70 ºC. C.beijerinckii cannot bear the relative high temperature that the pervaporation step achieves, 70 ºC, so it is necessary to maintain the bacteria in the fermentation vessel. Thus, an ultrafiltration step is placed between the bioreactor and the pervaporation unit. The organic solution is condensed and stored once it passes through the membrane at the pervaporation unit. After that, it is pumped to a series of four distillation columns. In the first one, water is obtained from the bottom and a two-phase stream exits from the top. The aqueous phase is recycled to the column as reflux and the organic phase is sent to second column. Here, acetone is recovered at 37.27 ºC and 0.5 bar with a purity of 99.00 %w. Finally, the solution resulting from the bottom of the second column is fed to the third column, butanol column, where butanol is obtained at 117.98 ºC and 1 bar with a


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purity of 99.99 %w at the bottom, while the top stream is sent to a liquid-liquid separator, previously mixed with the top of the last column. The organic phase goes back to butanol column while aqueous phase exit to last column where water is recovered at the bottom. Figure 1 shows a complete flowsheet of the process described in this section.

Figure 1.Flow diagram of second-generation biobutanol process.

3. Modeling approach. The simulation of the biorefinery is performed by integrating ChemCAD and EXCEL to take advantage of the rigorous simulation provided by ChemCAD while integrating the experimental results obtained in previous works. Corn stover, and its components such as cellulose, hemicellulose, lignin, protein, lipid and ash, the enzymes and the bacteria culture are not included in the database of this software package. Thus, the first stage consists of creating these species. Other sugars, as xylose and arabinose, do not appear either, so equivalent glucose is calculated, showing the amount of total sugars in the form of equivalent glucose.



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3.1. Biomass Degradation

Pretreatment The process starts by milling the biomass. In this step, size reduction takes place (around 1-2 mm), which makes corn stover more accessible to the external agents. The mill consumes 45 kWh/ kg.11 Figure 2 shows the detail of the flowsheet for this task. Once corn stover is inside reactor (R-01), biomass is broken down. Hemicellulose is dissolved in the form of pentoses, mainly xylose. Other compounds appear as acetic acid, HMF, glucose and arabinose, consequence of operating conditions which affect at overall structure. Biomass Cellulose + Hemicellulose Sugars The conversion of the reaction was obtained experimentally (see part 1 12). This process is modelled as an input-output stage using the experimental results presented in Table 1, obtained in the previous part.12 The mass balance is implemented in ChemCAD as an input-output model, since the depolymerization is a complex process.

Table 1 Pretreatment, mass balance

Sugars

Compound Glucose Xylose Arabinose Acetic Acid HMF Biomass reduction

g compound / kg biomass 17.43 182.48 42.71 30.63 1.31 40%

Recovered solution

21.80 l/kg biomass

In terms of energy, the reaction is endothermic and heat needs to be provided, see Table 2 for the results.

= − ( + )


(1)

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Figure 2. Pretreatment stage diagram.

Besides the energy involved in the reaction, the streams are heated or cooled to maintain the operating temperatures. Thus, HX-01 heats up the acid solution to 121ºC while the product stream is cooled down in HX-02 to 50ºC. The energy recovered in this heat exchanger is reused within the process. Heat exchangers are simulated using ChemCAD, with NRTL as the thermodynamic model. Thus, the thermal energy involved in the pretreatment stage is computed using this package. Only the heat of reaction is calculated using equation (1). Energy results reveal that pretreatment is thermally sustainable but it needs a large amount of electrical energy in the physical pretreatment. Table 2 Pretreatment, energy balance

Device

G HX-01 R-01 HX-02

Heat Exchanged (MJ/ kg biomass) 162.45 2.63 24.50 -3.33

Hydrolysis The biomass, coming from pretreatment, is fed to the hydrolysis reactor, R-02, where it is mixed with an enzymatic cocktail and a buffer solution, see part 1 for details. Figure 3 shows the section of the flowsheet which corresponds to the hydrolysis stage. Hydrolysis is set to take place at 50 ºC with a solid loading of 2.5% and an enzymes to biomass ratio of 400 and 440 µl/g of N50013 and N50010,



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respectively. Here, cellulose is split up in monomers, mainly glucose. After the saccharification process, the stream is fed to a centrifuge to remove the solid biomass remaining. Hence, a juice rich in sugars is obtained.

Figure 3. Hydrolysis stage diagram.

The mass balance for the saccharification is based on the experimental results of part 1 and it is summarized in Table 3.12 The reactor and centrifuge are modeled using the experimental results obtained in the lab. Table 3 Hydrolysis, mass balance

Sugars

Compound Glucose Xylose

Arabinose Biomass reduction Recovered solution

g compound / kg pretreated biomass 392.65 44.24 217.01 60% 18.75 l/ kg biomass

Regarding the energy balance, the buffer stream is heated up in HX-03 to the operating conditions, 50ºC. Moreover, it is necessary to provide energy in order to maintain the optimal hydrolysis temperature inside R-02. The energy required for the saccharification section of the process is 2.52 and 19.82 MJ / kg of biomass, respectively, adding up to 23.34 MJ/ kg of biomass. Fermentation Set up The rich sugar streams from previous stages are mixed. Then, a two stage procedure is used to prepare them for the fermentation; in Figure1, this stage corresponds to units NF, T-02 and HX-04. The


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first step deals with sugars concentration; nanofiltration is used to remove mainly water in order to concentrate sugars up to optimal conditions for the operation of the bacteria C.beijerinckii. To compute the yield of this step, its performance was evaluated in a membrane filtration pilot plant at ITACyL (Model with LM20 Lab Module™ by GEA with NF flat membranes by Alfa Laval). A sample of five liters of solution coming from pretreatment and hydrolysis was nanofiltrated at ambient temperature and 16 bar of pressure. Water and ions as SO42- can pass through the membrane; however, sugars are kept in the retentate.13 Therefore, retentate will continue to fermentation, while permeate, mainly water, is reused in pretreatment step. The obtained results are presented in Table 4. We develop a black box model within ChemCAD using the species separator to model this unit. Table 4 Nanofiltration yield

Compound Glucose Xylose

Arabinose Acetic Acid HMF

%w retentate 100 97 100 27 100

The second step aims at neutralizing the sulfuric acid from the pretreatment. Thus, calcium hydroxide is added in DP-01 in order to obtain a pH of 5.4.

+ () → +

(2)

Finally, before fermentation, the stream is cooled down to 34 ºC in HX-04 to get optimal conditions in fermentation process. The heat recovered in HX-04 is 0.55 MJ/ kg biomass.

3.2. Fermentation coupled with Pervaporation The fermentation is coupled with a pervaporation stage for an increased utilization of sugars; see Figure (4) for a detail of the flowsheet.



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Figure 4. Fermentation coupled with pervaporation stage diagram.

The fermentation of the sugars produces, mainly, butanol and acetone. Qureshi et al.8 evaluated the AB fermentation for the hyper-producing mutant strain of Clostridium beijerinkii BA101. Based on the information of the metabolism of this culture, we propose the following chemical equation to perform the mass balance, see eq (3). 100 C6H12O682.68 C4H10O + 23.15 C3H6O + 2.03 C2H6O + 195.76 CO2 + 10.44 H2 + 100.60 H2O (3) The mass balance is based on the stoichiometry of the previous reaction which has been validated experimentally in terms of the main species, butanol, acetone and ethanol. It is assumed that 100% sugars are transformed mainly into butanol and acetone due to the pervaporation stage. No intermediate acids remain. As long as the fermentation progresses, the pervaporation removes organic solvents and water, avoiding strain inhibition and concentrating the sugars. It is known that both processes have to be run at same time to reach the mentioned yield in Qureshi et al.8. However, for the mass balance to hold, fermentation and pervaporation have been modeled individually as different black boxes, following equation 3 for fermentation and eqs. (4) – (5) for the pervaporation step. The flux for permeate and retentate are computed using the information in Huang et al10 and eqs. (4)(5). A silicon based mambrane is selected. Thus, the fluxes through membrane for butanol and acetone are 907 g/m2 h and 414.4 g/m2 h, and a selectivity,, of 49 and 13.7, respectively is used assuming a butanol yield in permeate side of 98.65%w. The selectivity to ethanol is 4.3. Therefore, together with the small amount of ethanol produced in the fermentation, at the level of other minor species such as organic acids, the flow of ethanol across the membrane was negligible, 1.510-5 kg/h. Thus, we assume that no ethanol transpases the membrane.


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=

!$ ("#!)% & $("#&)%

' =

() *∙

(4) (5)

Where: x: Concentration of i compound in permeate side. y: Concentration of i compound in retentate side. W: Weight of i compound in permeate side. A: Area. t: Time. Regarding the energy balance, proper operating pressures and temperatures have to be maintained at both, fermentation and pervaporation unit. The energy balance to keep the temperature constant in the fermentor, FERM, is written as eq (6).

. = ∆0 − ∆

(6)

Furthermore, the stream fed to the pervaporation is heated up to 70ºC in HX-05 and cool down in HX06 to 34ºC before recycling the retentate to the fermentor. Simple energy balances allow computing the heat load. The results are shown in Table 5. On the permeate side, a condenser (HX-07) is used to get the solution as liquid at -3 ºC, so cooling stream will be a 25%v mixture of glycol and water. Since we need refrigerants for this operation, we do not consider this unit for the regular net energy balance that involves steam and cooling water as utilities alone. Table 5 Fermentation, energy balance

Device

Heat Exchanged (MJ/ kg biomass)

Bioreactor jacket HX-06 HX -07 HX -08

-0.46 1.24 -0.49 -12.94



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3.3. Simulation of Distillation Columns The distillation column section is entirely simulated with ChemCAD. To account for the complex liquidliquid-vapor equilibrium, NRTL is selected as the thermodynamic model to perform the simulation. Figure 5 shows a detail of the purification section of the flowsheet. The separation of the butanol-acetone mixture starts with a distillation column of 20 trays that operates at 1 bar, TC-01. Feed is introduced at the second tray. Most of this stream is water so that it is necessary to remove as much water as possible. Hence, reboiler temperature is fixed at almost 100 ºC to obtain close to pure water at the bottom. At the top, butanol-water azeotrope and acetone are obtained. After condensation, a liquid-liquid separator operating at 82 ºC is used. The aqueous phase goes back to the first tray of the column as reflux, while organic phase progresses to the second column, TC-02.

Figure 5. Separation process diagram.

TC-02 has 15 trays; the stream is fed to the seventh tray at 0.5 bar and 62ºC. In this case, the specifications of the column are as follows: the specific weight composition at the top which it will be set at 99%w of acetone; and the temperature of reboiler heater it is established at 74.8ºC, recovering most of acetone with traces of water. Finally, a mixture with water and butanol is obtained as organic single phase. Thus, this stream will be sent to the butanol column (TC-03). Column TC-03 consists of 20 trays and operating at 1 bar. Pure butanol will be recovered at the bottom while azeotrope exits at the top. As well as the first column, there is a liquid-liquid separator which splits


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the two phases. The organic one goes back to TC-03 and, on the other hand, the aqueous phase is led to the last column, TC-04, where pure water is obtained at the bottom. TC-04 has 15 plates, operating at 1 bar. The feed is introduced at second tray. For this particular column, we specify the mass flows. At the bottom of TC-03 for butanol and at the bottom of TC-04 for water, of this way, material balance will be closed. Table 6 collects the results of the energy balances to the various heat exchangers involved in the separation process. All the values have been obtained from the ChemCAD simulation. Table 6 Separation, energy balance

Device

Heat Exchanged (MJ/ kg biomass)

HX -08 TC-01 HX -09 HX -10 TC-02 TC-03 HX -11 TC-04 HX -12

1.94 0.73 -0.53 -0.02 0.01 0.33 -0.30 0.06 -0.06

4. Results We divide the results into several sections to present the main values related to the mass and energy balances, providing the needs for raw materials and the amount of butanol and acetone produced as well as the energy involved. Next, using a technoeconomic approach based on literature, we estimate the investment and production costs for such a facility. Finally, we present comparison of the results of this work with others in the literature. Note that most of the previous studies on this topic use the traditional ABE production process which results it lower yield to butanol and a more complex separation stage.

4.1.Summary of material and energy balances The capacity of the plant is decided based on the availability of the main raw material, corn stover. The simulated plant is designed for a consumption rate of 15 x 103 kg/h corn stover. It represents an amount of 126 kt/ year of corn stover. The production of butanol is 26.7 kt/ year and 5.7 kt/ year of acetone. In



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addition, 25.2 kt/ year of biomass residue are obtained which may be used as fuel, contributing to plant´s utilities, or as cattle feed, due to high percentage of proteins. In Table 7, all raw materials, needed for the process per kg of dry biomass, are presented. The yield to butanol obtained, 0.21kg per kg of biomass, is close to the one for ethanol in second generation facilities14 and far larger to the previously obtained in the production of butanol.15 Table 7 Total mass balance

Raw material Sulfuric acid Calcium hydroxide Enzymes Distilled water Buffer Citric acid Sodium hydroxide Inoculum

Quantity (kg / kg biomass) 0.083 0.010 0.619 24.000 0.216 0.084 0.500

Products Butanol Acetone Residual biomass

0.211 0.045 0.200

The summary of the energy involved in the process in shown in Table 8 by process stages. Note that the use of refrigerant is not considered in this table for simplicity. We address the use of refrigerant separately as part of auxiliary facilities within the investment cost. Biobutanol process needs a contribution of thermal energy from outside. It is important to note that the overall energy balance considers a total energy integration using heat released during whole process. Residual biomass, which is obtained once it suffers digestibility, could be taken to produce energy and so, process would be self-sufficient energetically or, at least, it provides an important fraction of it. Table 8 Net energy balance

Stage Pretreatment Hydrolysis Set up Fermentation Separation Net balance

Net Heat Exchange (MJ/ kg biomass) 23.80 23.34 -0.55 -12.65 2.16 36.10


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4.2.- Economic evaluation The results of the mass and energy balances are used to estimate the production and investment cost of the facility. Factorial method is used to compute the investment required for our plan.16 This method requires the estimation of the equipment cost. We evaluate the equipment cost using Matche17 and the size of the units computed as detailed in the supplementary material of Almena and Martín18 by the mass and energy balances and short cut design procedures presented along the paper. Hence, each one of budget items are known applying the method described by Silla19 see Table 9. Table 9 Direct fixed capital cost

A. Total plant direct cost

(€)

Equipment purchase cost Installation Process piping Instrumentation Insulation Electrical Buildings and yard movements Auxiliary facilities

42,190,000 37,314,000 16,876,000 6,328,000 2,531,000 5,063,000 3,000,000 21,939,000

B. Total plant indirect cost Engineering Construction

13,500,000 14,342,000

C. Contractor´s fee and contingency Contractor´s fee Contingency

7,594,000 15,189,000

D. Total fixed capital cost

185,866,000

The operating cost involves direct, indirect, and general costs. In the same way that total investment, operating cost is calculated. In this case, the raw material is the main item. In Table 10, a breakdown of the cost is presented. The prices of the species involved have been taken from Qureshi, 8 except corn stover, which comes from Cámara Agraria Provincial de León.20 Figure 6 shows the breakdown of the investment costs.



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Figure 6. Total capital investment.

Here, the most important items are equipment cost, 22.70% and their installation, 20.08% of total capital investment. The depreciation cost of equipment is considered constant for the life period of each unit and included in operating cost. Thus, 75% of the cost corresponds to units. The other items are engineering, 7.26%, construction, 7.72%, contractor´s fee, 4.09%, and contingency, 8.17%. Therefore, the total capital investment adds up to 186 MM€. Table 10 Materials cost

Raw material

Amount (t/year)

Unit cost (€/t)

Total cost (€/year)

(%)

Corn Stover Sulfuric Acid Water Sodium Hydroxide Citric Acid Calcium Hydroxide Enzymes Total

126,000 10,458 250,000 10,584 27,216 1,260 77,994

36 50 0.27 17.67 50 50 119

4,536,000 523,000 66,000 187,000 1,360,000 63,000 9,281,000 16,000,000

28.32% 3.26% 0.41% 1.17% 8.50% 0.39% 57.95% 100%


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Silla’s method19 is used to estimate the production costs. Amortization is calculated dividing the equipment cost by the useful lifetime and utilities include refrigerant, steam and cooling water and electricity among others. We assume that they represent a 20% of operating cost21 Labor is determined by the needs of the process and salaries from Ministerio de Empleo y Seguridad Social.22 Table 11 summarizes the results. Table 11 Operating cost Cost item Raw materials Utilities Labor dependent Operating supervision Quality control Maintenance labor Maintenance material Direct cost

(€) 16,000,000 9,561,000 3,098,000 620,000 620,000 165,000 110,000 30,174,000

Cost item Depreciation cost Property taxes Insurance Plant overhead cost Indirect cost Administrative Marketing Financing Research and investment General Cost

Production Cost

(€) 3,135,000 123,000 61,000 2,678,000 5,947,000 2,145,000 6,437,000 147,000 2,742,000 11,471,000

47,592,000

Figure 7 shows the breakdown of the operating costs. The three most important contributions to the operating costs come from the raw materials, 34%, the utilities, 20%, and marketing 13%. In addition, there are other important items that, together with the previous three, represent 95% of total operating cost. They are labor, 6.5%, depreciation, 6.5%, overheads, 5.6%, administrative 4.5% and research, 5.7%, costs. Thus, total operating cost adds up to 48 M€ per year.



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Figure 7. Operating cost.

On the other side, the plant´s revenue, showed in Table 12, consists of selling butanol as the main product, and acetone and waste biomass as by-products. The amount of each one of these has been presented in section 2.3. The prices are taken from ICIS web page.23 Prices were updated using Spanish IPQ (Chemistry Price Index) as show eq. 7. 1 2345 = 1 2336 7

89:2345 89:2336

= 2.18

€ @A

(7)

Table 12 Revenue of by-products Revenues €/t

t annual

€/annual

Acetone

2180

5.670,00

12.360.600,00

Biomass

250

25.200,00

6.300.000,00

Total

18.662.600,00

Once operating cost and revenues of by-product are determined, the butanol production price is computed. In this case, for a plant which consumes 126 x 103 t/ year of corn stover and produces 26.7 x 103 t/ year of butanol, the production cost is 1.09 €/kg. The butanol production price obtained is lower than


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the ones from ABE processes in the literature, due to the higher yield of the AB fermentation coupled with the pervaporation stage. It means that butanol becomes economically feasible, reducing second generation butanol production from Qureshi et al.’s7 work, 1.23 $/kg and, also, it is closer to compete against second generation Ethanol, 1 €/L24 We can find other more recent results in the literature for biobutanol production, but all of them use an ABE route. Tao et al25 presented a study showing values in the order of $3.6/gal. Procentese et al 26 presented an technoeconomic study for the production of butanol using ABE fermentation. The work neglected the cost of the biomass, which typically represents 33% of the production cost. As a result, their process was competitive, but the cost of the raw material needed to be included. Lately, Raj Baral and Shah27 estimated the production costs in $1.5/L. In all cases we can see the savings when using an AB route. 4.3.-Analysis of the results. It is possible to identify a number of factors that make this process economically challenging but promising. First, the cost of the grain corn (171€/t 20) that was used as feedstock for biofuels production. However, nowadays, development of markets for agricultural residues, as corn stover, wheat straw or burley straw, makes easier the future of the process. We also see that enzymes represent a large of the raw materials, see Table 10. For some time the estimation of the cost of the enzymes was oversimplified resulting in lower production costs for biofuels. Finally, the target for lignocelluslosic biomass is 50-80€/t which would reduce the production cost by around 2-3 M€. Second, the improvements on saccharafication and fermentation, which result in larger yield and larger productivity which means increasing of revenues. Finally, new techniques on separation process, such as pervaporation, replace distillation, avoiding an excessive waste of energy and consequently a overspending.

5. Conclusions In this paper we have developed process for the production of butanol from lignocellulosic biomass following the AB fermentation coupled with pervaporation. High yield to butanol and close to complete sugar conversion are achieved by in situ recovery of butanol an acetone. A technoeconomic analysis is carried combining experimental based models and rigorous simulations using ChemCAD.



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Second generation biobutanol process is closer to reach economic profitability. The process described by this work achieves a butanol production cost of 1.09 €/kg butanol. It needs a capital investment of 186 M€ with an operating cost of 48 M€/yr. The production cost is more competitive with current biofuels than the one reported from traditional ABE production processes, but further experimental work on pervaporation and membrane strucutre is needed as well as on the metabolism leading to butanol production.

Nomenclature Q: Thermal energy (kW) H: Enthalpy (kJ/kg) W: weight of i compound in permeate side (kg) A: area (m2) B: concentration of i compound in permeate side C: concentration of i compound in retentate side t: time (s) α: Membrane selectivity Units G: grinder CF: centrifuge M: mixer HX: heat exchanger SW: solid washer R: reactor NF: nanofiltration unit UF: ultrafiltration unit DP: deposit FERM: fermentor PERV: pervaporation unit VP: vacuum pump P: pump V: valve S: separator device PS: phases separator TC: tray column Acknowledgements We are grateful for the funding from Instituto Tecnológico Agrario de Castilla y León in this research and Salamanca Research for software licenses.

References [1] International Energy Agency (IEA). iea.org [accessed Jan 11,.15]


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[2] Martín, M.; Grossmann, I.E. Process optimization of FT- Diesel production from biomass. Accepted Ind. Eng. Chem. Res. 2011, 50 (23),13485–13499 [3] Martín, M.;Grossmann, I.E. Simultaneous dynamic optimization and heat integration for the coproduction of diesel substitutes: Biodiesel (FAME & FAEE) and glycerol ethers from algae oil. Ind. Eng. Chem. Res. 2014, 53, 11371-11383 [4] Durre, P., Biobutanol: an attractive biofuel Biotech. J. 2007, 2, 1525–15234 [5] Madihah, M.S.; Ariff, A.B.; Sahaid, K.M.; Suraini, A.A.; Karim, M.I.A. Direct fermentation of gelatinized sago starch to acetone–butanol–ethanol by Clostridium acetobutylicum. World J. Microbiol. Biotechnol. 2001, 17, 567–576 [6] Mariano, A.P.; Dias, M.O.S.; Junqueira, T.L.; . Cunha, M.P.; Bonomi,A.; Filho, R.M.; Butanol production in a first generation Brazilian sugarcane biorefinery: Technical aspects and economics of greenfield projects. Biores. Technol. 2013, 135, 316-323. [7] Qureshi, N.; Saha B.C.; Cotta, M.A.; Singh, V. An economic evaluation of biological conversion of wheat straw to butanol: A biofuel. Energ. Conv. Manag. 2013, 65, 456 - 462. [8] Qureshi, N.; Blaschek, H.P. Production of Acetone Butanol Ethanol (ABE) by a Hyper-Producing mutant strain of Clostridium beijerinckii BA101 and recovery by Pervaporation. Biotechnol.Prog. 1999, 15, 594-602. [9] Cai, D.; Zhang, T.; Zheng, J.; Chang, Z.; Wang, Z.; Qin, PY.; Tan., TW. Biobutanol from sweet sorghum bagasse hydrolysate by a hybrid pervaporation process. Biores. Technol. 2013. 145, 97–102. [10] Huang, J.; Meagher, M. Pervaporative recovery of n-butanol from aqueous solutions and ABE fermentation broth using thin-film silicalite-filled silicone composite membranes. J. Memb. Sci. 2001, 192 ,231–242. [11] Mani, S., Tabil, L.G.; Sokhansanj, S. Grinding performance amd physical properties wheat and barley straw, corn stover and switchgrass. Biomass Bioenerg. 2004, 27, 339 – 352. [12] Malmierca, S.; Díez, R.; Paniagua, A.; Martín, M. Technoeconomic study of biobutanol production part 1: Biomass pretreatment and hydrolysis. Ind. Eng. Chem. Res. 2016 Submitted [13] Sulzer Chemtech. Membrane Technology. [accessed Feb, 02, 16]. Available on the web: http://www.sulzer.com/es//media/Documents/ProductsAndServices/Separation_Technology/Membrane_T echnology/Brochures/Membrane_Technology.pdf [14] Martín, M.; Grossmann, I.E. Energy optimization of lignocellulosic bioethanol production via Hydrolysis of Switchgrass. AIChE J. 2012 ,58 (5), 1538-1549 [15] Jiang, W.; Zhao, J.; Wang, Z.; Yang, S.T. Stable high-titer n-butanol production from sucrose and sugarcane juice by Clostridium acetobutylicum JB200 in repeated batch fermentations. Biores. Technol. 2014. 163, 172-179. [16] SInnot, R.K.; Towler, G. Chemical Engineering Design, 5th Edition. Elsevier Singapore. 2009 [17] www.matche.com [ accessed Jan 11,15] [18] Almena, A.; Martín, M . Techno-economic analysis of the production of epiclorhidrin from glycerol. Ind. Eng. Chem. Res. 2016, 55 (12). 3226-3238



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[19] Silla, H. Chemical Process Engineering. Design and Economics. Ed. Heinz Heinemann, Estados Unidos, 2003. Chapter 2, Production and capital cost estimation, pp. 56. [20] Cámara Agraria Provincial de León, .camaraagrariadeleon.com/index.php/cereales [accessed Feb, 02, 5] [21] Vian Ortuño, A., El pronóstico económico de en química industrial. 4ª Ed. Madrid: Eudema S. A, 1991 ISBN 84-7754-082-9. [22] Ministerio de Empleo y Seguridad Social. Boletín oficial del estado nº 68. 20 of March of 2014. Available on the web: http://www.boe.es/boe/dias/2014/03/20/pdfs/BOE-A-2014-2971.pdf [23] ICIS. Chemical Industry News & Chemical Market Intelligence. Available on the web: [24] Klein-Marcuschamer, D.; Oleskowicz-Popiel, P.; Simmons, B.A.; Blanch, H.W.Technoeconomics analysis of biofuels: A wiki-based platform for lignocellulosic bioreineries. Biomass Bioenerg. 2010, 34,1914-1921. [25] Tao, L., Tan, E.C.D., McCormick, R., Zhangm M., Aden, A., He, X., Zigler, B.T. (2014) Technoeconomic analysis and life-cycle assessment of cellulosic isobutanol and comparison with cellulosic ethanol and n-butanol. Biofuels. Bioprod. Bioref. 2014, 8(1), 30-48 [26] Procentese, A., Guida, T., Raganati, F., Olivieri, G., Slatino, P., Mazocchella, A., Process Simulation of Biobutanol Production from Lignocellulosic Feedstocks Chem. Eng. Trans. 2014, 38, 343-348 [27] Raj Baral, N., Shah A. Techno-Economic Analysis of Cellulosic Butanol Production from Corn Stover through Acetone–Butanol–Ethanol Fermentation. Energy Fuels. 2016, 30(7), 5779-5790


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Figure 1.Flow diagram of second-generation biobutanol process. Figure 1 180x98mm (300 x 300 DPI)



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Figure 2. Pretreatment stage diagram. Figure 2 119x130mm (300 x 300 DPI)


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Figure 3. Hydrolysis stage diagram. Figure 3 90x70mm (300 x 300 DPI)



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Figure 4. Fermentation coupled with pervaporation stage diagram. Figure 4 199x54mm (300 x 300 DPI)


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Figure 5. Separation process diagram. Figure 5 137x80mm (300 x 300 DPI)



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Figure 6. Total capital investment. Figure 6 99x84mm (300 x 300 DPI)


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Figure 7. Operating cost. Figure 7 99x73mm (300 x 300 DPI)



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36x9mm (300 x 300 DPI)


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Technoeconomic Study of Biobutanol AB Production. 2. Process Design

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