Integrated Renewable production of ETBE from Switchgrass

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Integrated Renewable production of ETBE from Switchgrass Guillermo Galán, Mariano Martín, and Ignacio E. Grossmann ACS Sustainable Chem. Eng., Just Accepted Manuscript • DOI: 10.1021/ acssuschemeng.9b01004 • Publication Date (Web): 01 Apr 2019 Downloaded from http://pubs.acs.org on April 1, 2019

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ACS Sustainable Chemistry & Engineering

Integrated Renewable production of ETBE from Switchgrass Guillermo Galána, Mariano Martín1a Ignacio Grossmannb a b

Department of Chemical Engineering. University of Salamanca. Plz. Caídos 1-5, 37008, Spain Department of Chemical Engineering. Carnegie Mellon University. 5000 Forbes Ave. 15213 Pittsburgh PA.

Abstract. In this work, we propose the optimization of a flowsheet for the conceptual design of an integrated renewable production of ETBE from ethanol and i-butene from switchgrass. A superstructure embedding a number of alternatives is proposed. Two technologies are considered for switchgrass pretreatment, dilute acid and a novel ammonia fibre explosion (AFEX), so that the structure of the grass is broken down. Only glucose is fermented to produce i-butene, while xylose is fermented to ethanol. Ethanol is purified using a multieffect column followed by molecular sieves and a PSA-membrane system is used to upgrade the i-butene. The problem is formulated as an MINLP, and solved for each pretreatment as an NLP. Finally, an economic evaluation is performed including a sensitivity analysis. Biomass composition determines the byproduct obtained. Dilute acid is the selected pretreatment due to the largest yield to sugars and the possibility of adjusting the production of both, i-butene and ethanol for the needs to ETBE. For a facility that produces 90 kt/yr of ETBE , the investment adds up to 160 M€ for a production cost of 0.61€/kg, which is above market price. Even though the results look promising,, further experimental and scale up studies are required for validation.

Keywords: Energy; Biofuels; i-butene; Mathematical optimization; Ethanol; Switchgrass

1

Corresponding autor:Email address: [email protected]

1 ACS Paragon Plus Environment

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Introduction Biofuels such as ethanol or biodiesel involve within their chemical structure oxygen atoms. Their presence improves combustion properties. However, gasoline or crude based diesel as well as synthetic fuels such as Fischer – Tropsch are hydrocarbons. For years methyl-tert-butyl ether (MTBE) has been added as additive to improve the combustion efficiency, increasing the gasoline octane number. However, it is difficult to biodegrade and has shown impact on health as well as toxic effects on water and air. ETBE provides three times more energy than ethanol, doubles the octane number and reduces the vapor pressure for the mixture among other properties. Therefore, ETBE has been chosen as a better alternative as fuel oxygenate.1 ETBE can be produced from ethanol and i-butene or tert- butyl alcohol (TBA). The production of ETBE from ethanol and i-butene has been evaluated in the literature. The traditional process consists of a catalyzed reaction taking place at 60-90ºC under pressure.2 Puziy et al.3 noted that by operating at temperatures from 80ºC to 180ºC the selectivity to ETBE is increased. However, the reaction is in equilibrium. In order to improve the conversion, process intensification techniques have been used. Instead of using a reactor and a couple of distillation columns, a reactive distillation tower has been evaluated4-6 and optimized.7,8 Recent studies have compared the traditional and the integrated method resulting in the fact that although the reactive distillation column increased the conversion, the energy consumption is higher.9 However, other studies highlight not only the high conversions but also the low operating costs.5,6,10 However, typically the feed of i-butene contains a number of other chemicals. Ethanol has been obtained from renewable sources in first11 and second generation biorefineries.12,13 However, TBA is a by-product from the production of propylene oxide which reduces the raw material costs14 and i-butene has traditionally been an expensive C4, around 2 $·kg.1. The advantage is that i-butene can be produced from glucose. Recently, the company Global Bioenergies has patented a process from sugars for the production of i-butene.15-17 Martin and Grossmann18 developed an optimization framework for the production of i-butene from switchgrass resulting in promising production costs. Later, within the same group, i-butene was produced from algae starch and used within an integrated facility for the production of biodiesel and glycerol ethers without the need for fossil based intermediates.19 Therefore, it is possible to produce fully renewable ETBE from switchgrass by using C5’s to obtain ethanol and C6’s to produce i-butene.


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In this paper, the production process of ETBE from lignocellulosic raw materials has been evaluated, comparing two pretreatments of the lignocellulosic biomass for the simultaneous production of ethanol and ibutene and the production of ETBE from both in liquid phase. We propose a limited superstructure optimization approach where a flowsheet embedding the various process units involved in i-butene and ethanol production from switchgrass is constructed. The goal is to optimize the production process of both chemicals simultaneously from biomass to assess its competitiveness with current crude based production. The optimization of the system is formulated as a mixed-integer nonlinear programming (MINLP) problem, where the model involves a set of constraints representing mass and energy balances, experimentally based models and rules of thumb for all the units in the system. Finally, an economic evaluation is also performed. The effect of biomass composition on the operation of the integrated biorefinery is also addressed to guide the selection of biomass towards the renewable production of ETBE. The rest of the paper is structured as follows. Section 2 describes the integrated process. In section 3 the main modeling features are described. Section 4 presents the solution procedure. Section 5 summarizes the results including a sensitivity analysis of the biomass cost and the evaluation of the biomass composition for the integrated process to operate at its optimum. Overall Process Description Biomass follows a size reduction step before pretreatment. There are a large number of alternative pretreatments.20-23 However, among them, the ones that have reached commercial exploitation are, (1) dilute acid (H2SO4) pretreatment,24-27 and (2) ammonia fibre explosion (AFEX).21,28,29 Once the physical structure of the switchgrass is broken, cellulose is separated from hemicelluloses sugars. It has been experimentally proved that i-butene can be produced from glucose using Saccharomyces cerevisiae.16-18 Therefore, cellulose is hydrolyzed at 45-50ºC for 3 days to obtain glucose.24,25,30-32 The gas phase consists of i-butene together with CO2 and steam. First, the water vapor accompanying the gas phase is condensed, and then a PSA separation is suggested. In parallel, hemicelluloses are used for the production of ethanol. Xylose is fermented into ethanol using Z mobilis similar to second generation of ethanol production.12,25 Ethanol is dehydrated using a multieffect distillation column to separate the water – ethanol mixture, followed by a system of molecular sieves.12



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Finally, ETBE synthesis is carried out using a reactive distillation column between the i-butene and the ethanol, maximizing the conversion and obtaining a high purity ETBE from the bottoms of the column.4-6 Figure 1 shows the superstructure of the process. The biomass composition and the required feed to ibutene and ethanol determine the excess of any of the two products that are by products of the renewable production of ETBE.


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Fig 1. Superstructure for the renewable production of ETBE



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Mathematical modelling All the operations in the production of ETBE from switchgrass are modelled using mass and energy balances, rules of thumb, experimental yields, thermodynamic and chemical equilibrium33 as well as surrogate models for particular units such as the pretreatments, the ammonia recovery column, and the reactive distillation column based on detailed simulations and/or experimental data. The model for the superstructure is written in terms of total mass flows, component mass flows, component mass fractions, and temperatures of the streams in the network. The components in the system belong to the set J = { Water, i-butene, ethanol, ETBE, H2SO4, CaO, Ammonia, Protein, Cellulose, Hemi-Cellulose, Glucose, Xylose, Lignin, Ash, CO2, O2, Cells, Glycerol, Succinic acid, Acetic acid, Lactic acid, gypsum}. From the list, hemicellulose can be solubilized after pretreatment, while cellulose requires hydrolysis to be solubilized. Lignin remains suspended in small particle size. Gypsum precipitates and can be separated by filtration. Pretreatment In order for the fermentation to be effective, the bacteria must be able to reach the sugars. Any lignocellulosic raw material consists of a matrix of lignocellulose that protects the plant and maintains the structure. Within the lignin structure, the hemicelluloses and the cellulose constitutes the structure of the plant. This structure must be broken so that the polymers of sugar (cellulose and hemicellulose) can be further used. As raw material, switchgrass is considered, a native species in the Eastern part of the United States. We assume its composition to be 18.62 % moisture, 31.98 % Cellulose, 25.15 % Hemicellulose, 18.3 % Lignin and 5.85% Ash. The feedstock is washed and the size of the switchgrass is reduced by grinding34 so that further pretreatments are more effective.20 Both stages, washing and grinding, are considered only in terms of energy consumption (162 MJ·t-1)34 and cost analysis since they do not change the properties of the feedstock. Next, the two alternatives indicated above, dilute acid pretreatment and AFEX, are analyzed due to their high capability to degrade this structure.21,35-38 Ammonia fiber explosion (AFEX): This method consists of treating the lignocellulosic material at mild temperature and high pressure with ammonia to break the physical structure of the crop. In order to reduce the cost, the ammonia remaining in the slurry after the expansion should be recovered, and the slurry of biomass and 6 ACS Paragon Plus Environment

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water is sent to enzymatic treatment to break the polymers containing sugars.21,28,29,39 The pretreatment is modeled using the following assumptions. Garlock et al.39 developed a model based on design of experiments to evaluate the yield of the release of sugars from different switchgrass raw materials as a function of the ammonia (kg · kg-1 of biomass) and the water load, the operating temperature (C) and the contact time (min) at 2.0 MPa. We assume that xylose hydrolysis also takes place in R-01. Table 1.- Range of operating variables for dilute acid pretreatment

T (oC) Ammonia added (g·g-1) Water added (g·g-1) Residence time (min)

Lower bound 90 0.5 0.5 5

Upper bound 180 2 2 30

Yield  0.01·(88.7919  26.5272·amonia _ ratio  13.6733·water _ pret  1.6561·T _ afex  3.6793·time _ pret  4.4631·amonia _ ratio 2  0.0057·T _ afex 2

(1)

0.0270·time _ pret 2  0.4064·amonia _ ratio·time _ pret  0.1239·water _ pret·T _ afex 0.0132·T _ afex·time _ pret);

Next, the pressure is reduced and the content of the reactor is discharged to a blowdown tank. Since the reactor operates in batch mode, at least two reactors in parallel are fed into an intermediate storage tank to ensure continuous operation.21,40 Next, the ammonia remaining in the slurry is recovered by distillation, T-01, at high pressure (1.5 MPa).27,41 The distillation column is modeled using surrogate models from rigorous simulations in CHEMCAD. In particular, surrogate models are developed to compute the feed, the condenser and the reboiler temperatures as well as the purity and the recovery yield as a function of the feed composition in ammonia and the operating pressure. Thus, the column model is as follows:



Q HX2   m J,HX2,Col1 ·c p , j · T HX2,Col1  T Valv1,HX2 j



(2)

where

T HX2,Col1  a( NH3 /Water ) ·P 2  b( NH3 /Water ) ·P  c( NH3 /Water ) a( NH3 /Water )  1.0152· Load Ammonia _ Water   0.3996· Load Ammonia _ Water  - 2.118 2

b( NH3 /Water )   11.344  Load Ammonia _ Water   8.6088· Load Ammonia _ Water   25.159; 2

c( NH3 /Water )  68.912  Load Ammonia _ Water 

2

(3)

 174.79· Load Ammonia _ Water   59.214

And the purity



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Purity  a( NH3 /Water ) ·P 2  b( NH3 /Water ) ·P  c( NH3 /Water ) a( NH3 /Water )  0.5792· Load Ammonia _ Water   0.9987· Load Ammonia _ Water  +1.48 2

b( NH3 /Water )  9.9096  Load Ammonia _ Water   0.3672· Load Ammonia _ Water   12.319 2

(4)

c( NH3 /Water )  81.468  Load Ammonia _ Water   127.72  Load Ammonia _ Water   56.404 2

Purity  99 Purity  80 and the recovery yield of ammonia is given by eq. (5)

Yield  0.1123· Load _ Ammonia _ Water   0.2055· Load _ Ammonia _ Water   99.889; 2

(5)

Yield  99.99 Yield  99.9 The energy balance to the column is modeled using the following surrogate models for the condenser and the reboiler, respectively (kW) :

Qcondenser 



 4.8709· Load

_ Ammonia _ Water

  1.4989 

f c ( water ) [kg / s ]



f c ( water ) [kg / s ]

Q Re boiler  9.4089· Load _ Ammonia _ Water   2.7523

0.028

0.028

(6)

(7)

The exit temperatures of the column are computed as follows:

T(Col1, Mix 2)  a( NH3 /Water ) ·P 2  b( NH3 /Water ) ·P  c( NH3 /Water ) a( NH3 /Water )  15.315· Load Ammonia _ Water   13.426· Load Ammonia _ Water   4.9178 2

b( NH3 /Water )  127.73  Load Ammonia _ Water   103.7· Load Ammonia _ Water   47.523 2

(8)

c( NH3 /Water )  238.48  Load Ammonia _ Water   116.32  Load Ammonia _ Water   25.337 2

T Col1,Mix3  1.8617·P 2  24.16·P  77.491

(9)

The evaporated ammonia is compressed, condensed and mixed with the ammonia recovered in the distillation column in Mix2 and reused again. This is the key point in the economics of this process that has been improved over the years to tackle the disadvantage of recompressing ammonia. Following these stages, we assume that all of the ammonia is recovered. However, the traces that may be left, typically below 0.5%,40 are used as nutrients for the fermentation. Thus, the traces of ammonia in the feed to the second stage of the


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process are not considered. It is assumed that after the pretreatment, the monomer of glucose is generated. It will not be the molecule of glucose until the hydrolysis in which the monomer is hydrated, but for the sake of reducing the number of components, a dehydrated glucose is obtained that will be hydrated later on. Dilute acid: The yield of the pretreatment depends on the operating conditions. In the literature two main approaches have been developed, surface response models,42-44 and mechanistic kinetics.45 For superstructure optimization the first approach is more convenient. Recently, some groups44 have studied the sugars released from lignocellulosic raw materials using dilute sulfuric acid solutions as a function of the operating temperature, the concentration of the acid the residence time and the enzyme amount used, per gram of glucan, in the hydrolysis stage. It is assumed that after the pretreatment the biomass is broken down and we already have a monomer of glucose, which will be hydrated in the hydrolysis stage to obtain the sugars to be fermented. Xylose is solubilized as such after the pretreatment and therefore we consider it as such downstream. Using the experimental data provided in Shi’s paper,44 DOE based models were developed for the yield of the glucose and xylose released.18 Table 2.- Range of operating variables for dilute acid pretreatment

Lower bound 140 0.005 1 0.0048

T (oC) Acid concentration (g·g.1) Residence time (min) Enzyme load (g·g-1)

Upper bound 180 0.02 80 0.0966

The yield of glucose is given by:

yield _ cellu   0.00055171  0.00355819·T _ acid  0.00067402·conc _ acid _ mix  time _ pret·0.00100531  enzyme _ add·0.0394809 0.0186704·T _ acid·conc _ acid _ mix 0.00043556·T _ acid·time _ pret  0.0002265·T _ acid·enzyme _ add 0.0013224·conc _ acid _ mix·time _ pret 0.00083728·time _ pret·enzyme _ add  0.044353·conc _ acid _ mix·enzyme _ add  0.000014412·T _ acid 2 ; (10) The yield of xylose is given by:



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yield _ hemi   0.00015791  0.00056353·T _ acid  0.000694361·conc _ acid _ mix 0.00014507·time _ pret  enzyme _ add·0.01059248 0.02142606·T _ acid·conc _ acid _ mix 0.000694055·T _ acid·time _ pret  0.00013559·T _ acid·enzyme _ add 0.00145712·conc _ acid _ mix·time _ pret  0.04769633·conc _ acid _ mix·enzyme _ add 0.00138362·time _ pret·enzyme _ add  0.0000059419·T _ acid 2 (11) There is a possibility of producing furfurals in the process via dehydration of the sugars. These processes are favoured at high temperatures, 180ºC, a temperature that is not expected to be reached. Furthermore, the overliming is expected to detoxify a part form neutralize.26 In the analysis we do not consider their production following previous technoeconomic analysis.24 Next a flash evaporation of water (Flash 1) reduces the amount of water in the slurry and provides energy for the process. The slurry is separated in a mechanical centrifuge (Mec Sep 1). The liquid stream is treated with lime, CaO, to adjust the pH to the one needed in the hydrolysis (R-03)24-26,46 Lime is the cheapest chemical for this reaction due to the low cost of CaO, and also because the precipitation of gypsum (CaSO4) allows its easy separation from the liquid.47 The residence time in R-03 is 10 min. Neutralization reactions are exothermic, heating up the exiting stream. CaSO4 (gypsum) precipitates, and can be easily recovered from the liquid stream by filtration (Filter 1). Gypsum can be sold to improve the economics of the process. Some reports indicate the possibility of losses of sugars due to over liming up to 13%, but there is not specific data for our biomass and the low yield to xylose predicted by the model would cover for the general yield of the pretreatment including overliming It is assumed that xylose is already in sugar form after the pretreatment, eq (12), while cellulose is not solubilized in the pretreatment and it is after hydrolysis when the actual sugar molecule is available. The neutralized liquid stream is sent to ethanol production and the cellulose for hydrolysis. (C5 H 8O4 ) m  (m  1) H 2 O   mC5 H10 O5

H=79.0m kJ·mol1

(12)


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i-butene production The slurry with typically 20% in solids and at least 50% water,41 is hydrolyzed at atmospheric pressure to generate glucose from cellulose as given by eq. (13).24-26,31,48 The batch process lasts for 72 h in BR-01. The temperature must be adjusted to 50 ºC in HX26 for the reaction to take place. (C6 H10 O5 ) n  (n  1) H 2 O   nC6 H12 O6

(13)

H=22.1n kJ·mol1

It is after the hydration of the liberated monomers when glucose as a molecule is available in the liquid stream. We assume that buffer/storage tanks such as TK-03 are used to ensure the continuous operation of the process. Although they are considered for costing purposes, the dynamics is out of the scope of this paper and the mass balances do not account for them. The liberated glucose is fermented using a bacterium (S. cerevisiae) in BR-02. The reaction time is about 22-72 h at atmospheric pressure. The reaction takes place at 38 ºC

24

and therefore the feed is cooled

down. Furthermore, water is added so that fermentation is performed with 100 g·L−1 fermentable carbohydrates49 The main reaction is given by equation (14).49 yeast C6 H12 O6   C4 H8 + 2CO 2 + 2H 2 O + 2 ATP

H=-41.9 kJ·mol1

(14)

Formation of ATP from ADP and P requires 30.5 kJ·mol-1, and thus this is the energy involved in these reactions. A conversion of 85% based on the experimental results in the literature is assumed (25 g·g-1 vs the theoretical 31 g·g-1).49 Given the low aqueous solubility of i-butene, lower than 0.5% of the produced isobutene, it is assumed that all of it will go with the off gas. From the bottoms lignin is recovered using a centrifuge, Mec Sep 3. I-butene purification. The gas phase produced in the fermentor is saturated with moisture. Only CO2 and moisture are present with the i-butene. Moisture is removed by condensation, V-02, and further dehydration, while CO2 is removed using an adsorbent bed, MS-03 & MS-04. For simplicity the final dehydration of the i-butene takes place in the same pressure swing adsorption unit (PSA) used to capture the CO2 . A recovery of 99% of the ibutene is set.



Page 12 of 37

However, experimental validation is required. Two units operate in parallel to ensure continuous operation. The operating conditions of the PSA system are 25º C and 0.45 MPa. After the condensation of water, a compressor system with cooling, C-02, is used to feed the gas stream at 0.45 MPa to the PSA Ethanol production There is no experimental evidence that xylose can be converted into i-butene. Therefore, it will only be devoted to the production of ethanol via fermentation using Z. Mobilis. The model is similar to the one presented in Martín & Grossmann.12 The xylose is fermented mainly to ethanol at 38 ºC in BR-03, but there are a number of byproducts obtained via secondary reactions as seen in Table 3. The reaction time is about 24-72 h at 0.12 MPa to ensure anaerobic conditions. The maximum concentration of ethanol in the water is 6- 8%,24-26 and water may need to be fed to the fermentor. Using the NREL data base,32,48 the main reaction is given by equation (15) . yeast 3C5 H10 O5   5C2 H 5OH + 5CO 2

(15)

H=-74.986 kJ·mol xylose 1

The energy balance for the fermentor is calculated based only on the main reaction, eqs (15). The energy involved in the other reactions is neglected due to their low conversions. The solids are separated from the liquid stream in a mechanical press and a settling tank, V-01, before the stream is sent to the distillation column.11,26 Thus, the cells, the lignin and other solids are recovered in a two-stage process from the liquid phase so that the lignin can be used to obtain energy and improve the profitability of the process. Reaction

Table 3.- Chemical reactions in fermentor 3 Conversion

3 Xylose  5Ethanol + 5 CO2 Xylose + NH3  5 Z. mobilis + 2 H2O + 0.25 O2 3Xylose + 5 H2O  5Glycerol + 2.5 O2 3 Xylose + 5 CO2  5 Succinic Acid + 2.5 O2 2 Xylose  5 Acetic Acid 3 Xylose  5 Lactic Acid

Xylose 0.8 Xylose 0.03 Xylose 0.02 Xylose 0.03 Xylose 0.01 Xylose 0.01

Ethanol purification. Once the liquid stream is separated from the one with solids, recovering the lignin for its use as energy source for the process.12,26 Next, the ethanol must be dehydrated to fuel grade. The first stage is a beer column, T02-T-04, to remove a large amount of water. Next molecular sieves are employed, MS-01 & MS-02.12


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Based on previous work,11 the use of multieffect distillation columns is recommended for reducing energy consumption in ethanol dehydration. A model for a three effect distillation column is embedded in the superstructure. The columns are modeled each one using short-cut methods33 validated with rigorous process simulators19 considering a mass balance and recovery ratios, the energy balance to the condenser, and the reboiler and the linking constraints between the columns so that heat is transferred from the reboiler of the lower pressure column to the condenser of the higher pressure columns. In short, the model is as follows:

F   Fk k

Fi ,k  Di ,k  Wi ,k Di ,k  k Fi ,k Tbk  Tck 1  dt , k  COL

(16)

Tbk  Tck  dt , k  COL Qbk  Qwk 1 Pk 1  Pk R  1.5 The recovery of ethanol is fixed to 0.996. The pressure drop across the column is assumed to be 10% of the operating pressure. The distillate is a vapor since the final dehydration step requires this phase while the bottoms involves basically all the byproducts generated in the number if different reactions presented in table 3. The final dehydration of ethanol is carried out in a zeolite bed that is fed with the distillate of the multieffect column. It is assumed that at least a fraction of ethanol of 0.8 by weight is required for this stage to be effective. The stream is heated up to 95ºC . For costing purposes, we consider two beds in parallel so that the second one is regenerated to maintain continuous operation. Atmospheric air, with an assumed relative humidity of 70% at 20 ºC, is heated up to 95ºC. It removes the water that saturates the bed, and is cooled down to 25ºC. The flow of air required is that which allows a final humidity of 70%.

ETBE synthesis. The reaction between i-butene and ethanol is in equilibrium as follows: C4 H 8  C2 H 5OH € (CH 3 )3 COC2 H 5

(17)



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It is carried out in liquid phase from 40 to 90ºC depending on the catalyst and the pressure is in the range of 5-10 bar. The conventional reactor achieves a conversion above 98% by using an excess of i-butene and requires a set of distillation columns to purify and recover the products and recycle the i-butene.50 However, the conversion can be improved and the excess reduced by using a reactive distillation column that allows driving the equilibrium to products. One of the main differences with previous work is that when using nafta C4 fraction, the feed contains other chemicals apart from i-butene. However, in this case the fermentation does no generate any other hydrocarbon. A excess of 5% of ethanol with respect to the stoichiometry ratio is assumed. In this case from the top the azeotrope of ethanol and ETBE is reached and from the bottoms ETBE is obtained.6,10 Therefore, the model of this unit is based on the stoichiometry of the reaction assuming 100% conversion,5 obtaining the azeotrope composition over the top51 and pure ETBE from the bottoms. Note that the CO2 remaining in the feed stream exits the column from the top separately of the distillate. The reflux ratio and the energy involved is obtained as follows4,6,52 Due to the complex energy balance, data from the literature is used to estimate the reboiler energy balance. 8.3kW per 0.485 kmol/s of ETBE are typically used.5 To compute the cooling needs of the tower, a reflux ratio of 5 is assumed based on data from the literature.5 Finally, the exit temperatures are, from the bottom, 164 ºC using the vapour pressure computed from Rarey et al.51 From the top, the azeotrope is obtained. To compute the azeotropic composition and its temperature, a correlation is developed using the information in the literature.51 Unfortunately a extrapolation to the operating pressure is required since most of the data is up to around 2 bar. However, the smooth slope and trend found results in small increases when extrapolating. Azeotrope Ratio (ETBE/Ethanol)= 8.0390·1017 (T(K))-6.9905

(18)

T(Azeotrope) = 236.9053 (P(kPa))0.0786

(19)


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Figure 2.- Profile of the temperature of the azeotrope

Figure 3.- Azeotrope composition with the temperature. Along the process several waste streams are produced including gypsum, a gas stream of oxygen and CO2 from the fermentation of xylose, a liquid stream from the distillation column containing glycerol and other organics such as succinic acid, acetic acid, a stream of lignin and gas streams from the ibutene purification containing CO2 and hydrocarbons. It is out of the scope of this work to design the treatment of these waste streams.



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Solution procedure For simplicity, due to the presence of only one binary variable, the one related to the selection of the pretreatment, two non-linear optimization models (NLP’s) are solved involving 2400 equations and 3500 variables each. The major decision variables are the operating conditions of the pretreatment reactors, feed ratios and temperatures of operation, the split fraction, operating pressures and temperatures at the multieffect distillation column, and the split fraction to adjust the ratio of ethanol to i-butene before the column. This last one depends on the yield of each pretreatment, as well as the biomass composition. The NLP model is solved using a multistart optimization approach in GAMS with CONOPT 3.0 as the preferred solver. The objective function is based on a simplified profit including ETBE production and the thermal energy consume, due to the fact that it is the largest variable cost, eq. (20):

Z  mETBE  PSteam  i

Qi

(20)



Next, the heat exchanger network is developed53 and an economic evaluation is performed.54 A sensitivity analysis for the cost of biomass is also carried out, as well as a study on the biomass composition for the optimal operation of this integrated biorefinery. Results The production cost involves annualized equipment, chemicals (enzymes, sulfuric acid, CaO, ammonia, and the profit from gypsum), labor, utilities, raw material and the credit that can be obtained from ethanol. The cost for the equipment such as heat exchangers, fermentors, tanks, distillation columns, mechanical separation, filters, molecular sieves, is updated from the values calculated using the correlations developed by the authors, see supplementary material of Martín & Grossmann13 and Almena and Martín.55 The costs for utilities are updated from the literature, 19 $·t-1 Steam, 0.057 $·t-1 cooling water, Electricity: 1.7·10-8 $ ·J-1 56,57 and the base price for biomass is 30€/t . Finally, a sensitivity analysis will be performed to evaluate the effect of biomass cost from 30 to 100€/t in the production cost of ETBE. A feed of 18 kg/s of biomass is used, typical from bioethanol production facilities.12


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Facility operation Tables 4-5 summarize the operation of the major units in both processes, the rest can be seen in the supplementary material. One important issue is that since the dilute acid pretreatment exhibits separate yields to cellulose and to hemicellulose, the process involving this pretreatment does not produce any excess of i-butene or ethanol, see Table 5. In terms of the operating conditions, the dilute acid requires a higher temperature, 167ºC vs 132ºC, that together with the feed of acid are used to adjust the ratio of C6 sugars to C5 sugars obtained for the integrated facility to operate optimally. In particular, cellulose production is maximized due to the need for ibutene while the yield to hemicelluloses, the source for xylose and ethanol, is maintained low to reduce the energy consumption in its dehydration. As a result, the operating temperature is maintained in lower limits compared to the production of ethanol or i-butene alone.12,18 Due to the limited information on the performance of the AFEX reactor, an excess of ethanol is obtained so that enough cellulose is available for i-butene production, see Table 4. The excess of ethanol can be later considered as credit for the economic evaluation at the cost of its dehydration. The process using dilute acid pretreatment shows larger yield to ETBE, but no additional product is obtained. In Tables 4 and 5 the operation of the multieffect column is also reported. The operation of the multieffect column is interesting for reducing the consumption of energy within the process.11 Small differences can be found for both examples due to the composition of the feed to the column when fed from either of the pretreatments. The additional information on the yield to C6 and C5 sugars in the dilute acid pretreatment, together with the energy consumption involved in the dehydration of the excess of ethanol produced with using the AFEX pretreatment, results in a clear economic advantage of the process that uses dilute acid as pretreatment.

In both cases lignin is produced. The amount of lignin available would make it possible to generate a fair fraction of the needs, if not all. However, the humidity typically is an issue to process this lignin as energy source. The amount of lignin available would be able to provide up to 3.6 MW/kg of biomass. In the supplementary material, Tables S1 for AFEX and S2 for dilute acid pretreatment respectively, we present the stream data.



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Table 4.- AFEX pretreatment operating conditions Legend: LP: Low pressure: IP: Intermediate pressure: HP: High pressure : fraction of total feed to LP column  fraction of total feed to IP column

Feed Biomass (kg/s) AFEX reactor Ammonia Ratio (kg/kg dry biomass) Water added ratio (kg/kg dry biomass) Yield to sugars T(ºC) Ammonia recovery column P(bar) T (ºC) Tbot(ºC) Recovery Purity Multieffect column  (Fraction of feed to LP column)  (Fraction of feed to MP column) P(LP) mmHg LP/IP IP/HP Products Ethanol (kg/kg biomass) I-butene (kg/kg/biomass) Excess (Ethanol) ETBE (kg/kg/biomass) Utilities Thermal Energy (MW) Cooling (MW) Power (MW)

22 1.6 2 0.812 132 4.5 90 148 0.9998 0.84 0.075 0.226 217 2.56 2.82 0.095 0.077 0.025 0.136 132 90 7 18

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Table 5.- Dilute acid pretreatment operating conditions Legend: LP: Low pressure: IP: Intermediate pressure: HP: High pressure : fraction of total feed to LP column  fraction of total feed to IP column

Feed Biomass (kg/s) ACID reactor Acid concentration Water added ratio (kg/kg dry biomass) T(ºC) Enzyme added Yield to Cellulose Yield to hemicellulose Multieffect column  (Fraction of feed to LP column)  (Fraction of feed to MP column) P(LP) mmHg LP/IP IP/HP Products Ethanol (kg/kg biomass) I-butene (kg/kg/biomass) Excess ETBE (kg/kg/biomass) Utilities Thermal Energy (MW) Cooling (MW) Power (MW)

18 0.5% 1.5 167 44 0.97 0.675 0.084 0.238 150 2.33 2.29 0.078 0.096 0 0.16 21 19 6

Economic evaluation



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To compare both processes, a detailed economic evaluation is carried out computing the production and investment costs. The estimation of the investment is performed using the factor cost method.54 First the equipment cost is estimated using the mass and energy balances resulting from the optimization, and the unit estimation procedure in Martín and Grossmann13 updated by Almena and Martín.55 In the supplementary material we have provided Tables S3 and S4 including the size and the cost of the major units involved both processes. Next, the investment cost is evaluated as a function of the equipment costs assuming factors of 3.15 and 1.4, corresponding to a facility that processes fluids and solids, for the physical and total fixed costs, respectively.54 For the same production capacity of ETBE, the investment for the ACID based facility adds up to 160 M€, while that for the AFEX increases up to 195 M€. It is assumed that no inhibitory species are generated in the ACID based pretreatment. Figure 4 shows the distribution of the cost into the five main sections of the facility, the pretreatment, the production of each species, ethanol and i-butene, the production of ETBE and the heat exchanger network. The ACID based process shows an evenly distributed investment among HX, ethanol and ibutene sections, while the AFEX presents a larger contribution of the HX in the equipment cost.

Afex Acid Figure 4.- Breakdown of the units contribution to investment The production costs involve maintenance, labour, chemicals (H2SO4, CaO, NH3, Gypsum, Ammonia), raw material, utilities (Steam, electricity and cooling water) and other expenses including taxes, fees and administration.54 The mass and energy balances compute the utilities, raw materials and chemicals needs for the facility. Initially, a biomass price of 30 €/t is used. Without any credit due to mayor byproducts, the production cost of ETBE using an ACID based process adds up to 55 M€/yr, representing 0.61 €/kg of ETBE. However, using the AFEX process the production costs increase up to 97 M€/yr for a production cost of ETBE of 1.08 €/kg. However, in this second process ethanol is also produced. Assuming a price of 0.41 €/kg, a credit of 7 M€/yr can be obtained reducing the production cost of 1€/kg. A large amount of energy is involved in ammonia recovery. Figure 20 ACS Paragon Plus Environment

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5 shows the breakdown of the costs, where it is possible to see that utilities represent almost 45% of the production costs of ETBE using the AFEX process.

Acid Afex Figure 5.- Breakdown of the production costs.

Sensitivity analysis In this section two different topics are covered. On the one hand, the effect of the biomass and ethanol price in the final price of the ETBE. On the other hand, the effect of biomass composition on the operation of the integrated facility. The cost of ETBE depends on the biomass cost, as well as that of the bioethanol in the case of the process that uses the AFEX pretreatment where it is obtained as a by-product. The base price of ethanol was considered at 0.42 €/kg as in recent literature.58 A sensitivity analysis is performed on both variables, bioethanol price and biomass price to evaluate the ETBE production cost, see Figure 6. The ETBE prices in the literature range from values of 0.397 €/kg59 to others as high as $1.3/kg 60 or 1.27€/kg 7 that have been reported over the last few years. Lately prices of $0.2 /kg have also been reported due to the low prices of gasoline.61 Therefore, the production cost shown in the analysis presented values between two and three times the current values. the current values In Figure 6 it is possible to see that the renewable based ETBE can be, under certain raw material and byproduct prices, competitive with market prices. However, the price of ethanol has to be beyond current market price for the AFEX pretreatment to be selected. Note that if larger yield at the reactor is obtained the results may change.



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Figure 6.- Sensitivity analysis for the ETBE production cost

Finally, since the composition of the biomass is responsible for the availability of glucose and xylose, and in the end ethanol and i-butene, it is possible to select or engineer a biomass for the optimal production of ETBE. It can be considered a simultaneous process and product design problem that can allow selecting the proper biomass for the operation of integrated facilities similar to the work on algae design.62 Note that pretreatment efficiency is sensitive to the composition of the biomass. Therefore, it is reasonable to operate within a small range of the original composition (25-40% Cellulose, 20-30% Hemicellulose, 15-20% lignin, 6% and the rest moisture). From the optimization point of view, the model remains similar, but the composition of the biomass is now a variable and the problem is optimized as presented in Section 4. The dilute acid pretreatment is selected due to the availability of different models for glucose and xylose production. Table 6 shows a summary of the results. The solution targets maximum i-butene production, because it is the limiting factor, and improves the yield to sugars while reducing the fraction of hemicellulose. In this case, the freedom in both the composition and the yield from biomass to each of the sugars can be adjusted. The production capacity increases by almost 30% if the proper biomass type is selected. The cost estimation results in an investment of 165M€ for the production of 112kt/y and a production cost of 0.51€/kg. It represents a reduction of 16% in the production cost when using the generic biomass presented in the base case. Note that it is a purely theoretical analysis, and that the results must be validated, especially the possibility of such a biomass to be feasible and the yield of the pretreatment. Table 6.- Summary of operating conditions for “best biomass” case 22 ACS Paragon Plus Environment

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Feed Biomass (kg/s) 18 Wa 15% Cellulose 40% Hemicellulose 21.9% Lignin 17.1% Ash 6% ACID reactor Acid concentration 0.5% Water added ratio (kg/kg dry biomass) 1.5 T(ºC) 151 Enzyme added 44 Yield to Cellulose 0.97 Yield to hemicellulose 0.97 Multieffect column 0.087  0.247  P(LP) mmHg 150 LP/IP 2.33 IP/HP 2.29 Products Ethanol (kg/kg biomass) 1.75/18 I-butene (kg/kg/biomass) 2.064/18 Excess 0 ETBE (kg/kg/biomass) 3.613/18 Utilities Thermal Energy (MW) 21.5 Cooling (MW) 16.5 Power (MW) 8 This facility is a biorefinery capable of producing ethanol, i-butene and ETBE depending on the biomass type available and the demand of each one of the three. Note that the yield of the pretreatments must be validated.

Conclusions The production of ETBE from lignocellulosic switchgrass has been evaluated within the concept of an integrated biorefinery at a preliminary stage presenting a conceptual design. Both raw materials, the ethanol and the i- butene, are produced from fermentation of sugars. Since there are only experimental data to prove the conversion of glucose into i-butene, xylose is devoted to ethanol production. We use a superstructure optimization approach to evaluate the renewable production of ETBE. Surrogate models for each one of the operations from the pretreatments, including acid dilution and a new AFEX process that reduces the energy consumption by using ammonia absorption in water, hydrolysis of cellulose and



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sugar fermentation based on experimental data, multieffect distillation column and ethanol dehydration, and ibutene purification and a reduced order model for the reactive distillation column that is used to simultaneously produce and purify ETBE. The model is optimized resulting in the selection of dilute acid as the pretreatment of choice due to the larger yield to sugars, and the possibility of adjusting the production of both, i-butene and ethanol for the needs to ETBE, assuming that no inhibitors are produced. For a facility that produces 90 kt/yr of ETBE, the investment adds up to 160 M€ for a production cost of 0.61€/kg. This price is above current market value, 0.2 €/kg. Integrated facilities operate at their optimum for specific biomass compositions. This framework also allows evaluating the best use of each biomass depending on its composition, as long as the models for the pretreatments are valid. It is important to emphasize that the analysis of the proposed process is at conceptual level based on data and results that are available at laboratory scale. Therefore, further experimental validation and scale up studies are required to validated the proposed process.

Nomenclature a,b,c amonia_ratio conc_acid_mix Ci Di,k Dt: enzyme_add F: LoadAmmonia_water m(J, unit,unit1) Pi Pk Q(unit) Qbk Qwk R T(Unit, Unit1) Tbk Tck time_pret (min) T_acid T_afex time_pret water_pret

Fitting parameters for ammonia recovery column operation Ratio of ammonia added vs. dry biomass to AFEX pretreatment (g·g-1) Acid concentration at pretreatment in weight percentage. Material cost ($ ·g-1 or $·W-1) Flow of component I in distillate of column k Temperature increment Ratio of enzyme added to hydrolysis for acid pretreatment as function of the glucan (g·g-1 ) Feed to multieffect column Mass ratio between ammonia and water mass flow of component J from unit to unit 1(kg/s) Prices of component i (€/kg - €/kwh) Pressure of column k thermal energy involved in unit (W) Thermal flow in boiler of column k Thermal flow in condenser of column k Reflux ration Temperature of the stream from unit to unit 1 (ºC) Temperature in boiler of column k Temperature in condenser of column k Time for acid pretreatment Operating temperature acid pretreatment (oC) Operating temperature AFEX pretreatment (oC) Time for AFEX pretreatment (min) Ratio of water added to AFEX pretreatment function of the dry biomass(g·g-1) 24 ACS Paragon Plus Environment

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Wi,k W(unit) yield Symbols  

Flow of component I in residue of column k electrical power involved in unit (W) yield of the pretreatment / unit Latent heat steam (J·g-1) Separation ratio in column

Appendix Wa: Water Ibut: I-butene EtOH: Ethanol CaO: Lime CaSO4 : gypsum H2SO4 : Sulfuric acid CO2 : Carbon dioxide O2 : Oxygen. Units BR: Bioreactor T: Column D: tanks HX: Heat Exchanger MS: Molecular sieves MecSEp: Mechanical separation P: Pump R: Reactor Snk: Sink Src: Source TK: Storage tank V: Vessel Supporting Information Data on the streams and unit cost data can be found in the supporting information. This material is available free of charge via the Internet at http://pubs.acs.org. Acknowledgments The authors acknowledge JCyL grant SA026G18 and USAL PSEM3 for financial support References. [1] Yee, K.F.; Mohamed, A.R.; Tan, S.H. A review on the evolution of ethyl tert-butyl ether (ETBE) and its future prospects . Renew. Sust. Energy. Revs. 2013, 22, 604-620 DOI: 10.1016/j.rser.2013.02.016 [2] Menezes, E.W.; Cataluña, R. Optimization of ETBE (ethyl ter-butyl ether) production process. Fuel Process.Technol. 2008, 89(11), 1148-1152 DOI: 10.1016/j.fuproc.2008.05.006 [3] Puziy, A.M.; Poddubnaya, O.I.; Kochkin, Y.N.; Vlasenko, N.V.; Tsyba, M.M. Acid properties of phosphoric acid activated carbons and their catalytic behavior in ethyl-tert-butyl ether synthesis Carbon, 2010, 48, 706-713 DOI: 10.1016/j.carbon.2009.10.015 25 ACS Paragon Plus Environment


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For Table of Contents Use Only



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Ibutene ETBE

Ethanol Integrated Renewable production of ETBE from Switchgrass Guillermo Galána, Mariano Martín2a Ignacio Grossmannb Renewable ETBE is produced from lignocellulosic raw materials via i-butene and ethanol production from glucose and xylose and its further reaction.

2

Corresponding autor:Email address: [email protected]


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Figure 1. Superstructure for the renewable production of ETBE 287x102mm (300 x 300 DPI)



Figure 2.- Profile of the temperature of the azeotrope 90x52mm (300 x 300 DPI)


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Figure 3.- Azeotrope composition with the temperature. 90x53mm (300 x 300 DPI)



Figure 4.- Breakdown of the units contribution to investment 180x68mm (300 x 300 DPI)


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Figure 5.- Breakdown of the production costs. 180x54mm (300 x 300 DPI)



Figure 6.- Sensitivity analysis for the ETBE production cost 90x52mm (300 x 300 DPI)


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TOC 180x53mm (300 x 300 DPI)


Integrated Renewable production of ETBE from Switchgrass

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