Optimal Production of Dimethyl Ether from Switchgrass-Based Syngas

Article pubs.acs.org/IECR

Optimal Production of Dimethyl Ether from Switchgrass-Based Syngas via Direct Synthesis Estela Peral and Mariano Martín* Department of Chemical Engineering, University of Salamanca, Plz. Caídos 1-5 37008, Salamanca, Spain ABSTRACT: A conceptual optimal design for the production of dimethyl ether (DME) and/or power from switchgrass is proposed. A superstructure is formulated embedding two different gasification technologies, direct and indirect, two reforming modes, partial oxidation or steam reforming, gas cleaning, and composition adjustment. Next, DME is produced following direct synthesis, a novel one-step technology. The unreacted gas can be either recycled or used within a Brayton cycle for the simultaneous production of DME and power. The problem is formulated as a mixed integer nonlinear programming problem. The optimal topology involves indirect gasification followed by steam reforming. DME production is favored over power for current electricity prices. Thus, the investment cost for the production of 197 kt/yr of DME adds up to $133 million, at a price of $0.31/kg. process,21 but only a simulation-based study for a fixed flowsheet is available in which the optimization is performed using a sensitivity-based approach.16 Recently, Azizi et al.22 presented a descriptive review on the technologies to produce DME in which the main operating conditions of different reactor configurations can be found. In this paper, mathematical optimization techniques are used to design a process based on single-step synthesis of DME from biomass, switchgrass. We propose a conceptual design by solving the optimization of a superstructure embedding the various process units involved in DME production, considering alternatives for some of the biomass processing technologies as well as for the use of the syngas, to evaluate the trade-offs between DME and power production. The aim is to optimize the structure and the operating conditions minimizing the production cost. The optimization of the superstructure is formulated as a mixed integer nonlinear programming problem (MINLP) in which the model involves a set of constraints representing mass and energy balances for all the units in the system. This particular problem is solved by exhaustive enumeration, fixing the binary variables to the various choices, reducing the problem to nonlinear programming (NLP) subproblems. The variables of the optimization problems are the operating conditions at the gasifier and the DME synthesis reactor and the syngas composition adjustment. We then design a heat exchanger network. Subsequently, a sensitivity analysis is performed to determine the limiting prices for DME and power to produce one or both of them. Finally, a detailed economic evaluation is carried out to determine the production and investment costs.

1. INTRODUCCION Biofuels have become one of the major alternatives to reduce our dependency on fossil fuels. Typically, researchers have focused their efforts on the production of bioethanol, Fischer− Tropsch liquids (FT), and biodiesel from biomass because of their easy implementation in the transportation system.1−3 However, there are a number of alternative fuels that can be produced out of the same raw materials such as 2,5dimethylfuran (DMF) or dimethyl ether (DME)2,4 from lignocelluloses and/or algae, or we can use byproducts of the current biofuel industry, such as glycerol, to obtain methanol,5 ethanol,6 or alternative fuels like glycerol ethers.7 In particular, DME is a multipurpose fuel. It has properties similar to those of propane and butane; therefore, it can replace liquefied petroleum gas, LPG. It can also be used as a diesel substitute, reducing NOx, SOx, and particulate emissions. Furthermore, DME can also replace chlorofluorocarbons as aerosol propellant. The advantage of producing DME is that the cetane number is closer to that of crude based oil, and in fact, most of the production process is similar to the production of bioethanol or FT fuels because it is also based on catalytic synthesis from syngas. It is possible to produce DME using a two-step reaction process comprising methanol synthesis followed by its dehydration. In the literature, a number of simulation studies on the production of DME using process simulators, ASPEN Plus, are available. Some focus on the production of DME8−10 others expand the scope to include the cogeneration of DME and electricity from coal and natural gas11,12 or the production of DME, methanol, and dimethyl carbonate (DMC) from coal.4 Finally, some works report the use of DME for the production of ethanol13 or as an intermediate within a biorefinery.14 Pascall and Adams15 performed a rigorous analysis of the purification of the DME using Aspen Plus. Alternatively, a single-step reaction to synthesize DME directly can be used.16 In this case, a hybrid catalyst is used, avoiding the thermodynamic limitations of the methanol synthesis. The single-step reaction leads to a high conversion of CO.17−20 Besides that, the investment is expected to be lower than the two-step reaction © 2015 American Chemical Society

Received: Revised: Accepted: Published: 7465

March 2, 2015 July 2, 2015 July 15, 2015 July 15, 2015 DOI: 10.1021/acs.iecr.5b00823 Ind. Eng. Chem. Res. 2015, 54, 7465−7475

Article

Industrial & Engineering Chemistry Research Table 1. Correlations for Modeling the Indirect Low-Pressure Gasifier values and correlations26 (TFercogas, °F)

variable mass of gas (kg)

(28.993 − 0.043325(TFercogas) + 0.000020966(TFercogas)2 ) (DryMass/0.454)ρgas

molar fraction C6H6 tar H2S NH3 CO2 CO H2

(1)

0.001 0.002 0.0007 0.003

0.01[− 9.5251 + 0.0377889(TFercogas) − 0.000014927(TFercogas)2 ] 2

0.01[133.46 − 0.1029(TFercogas) + 0.000028792(TFercogas) ]

(3)

2


(4)

CH4

0.01[− 13.82 + 0.044179(TFercogas) − 0.000016167(TFercogas) ]

C2H2

0.01[− 4.3114 + 0.0054499(TFercogas) − 0.000001561(TFercogas)2 ]

C2H4 C2H6

2

2

0.01[− 38.258 + 0.058435(TFercogas) − 0.000019868(TFercogas) ] 2


2. OVERALL PROCESS DESCRIPTION The process superstructure consists of four different parts: gasification, gas cleanup and H2-to-CO ratio adjustment, DME synthesis, and power generation. For the first section, two different technologies are considered: (1) indirect low-pressure gasification with steam, in which the combustion of char in a parallel equipment (combustor) provides the energy for the gasification of the biomass and (2) direct high-pressure gasification of the raw material with steam and oxygen to avoid the dilution of the gas. The second section comprises technologies to remove solids from the gas as well as other compounds like hydrocarbons, NH3, CO2, or H2S and to adjust the gas composition. The hydrocarbons are partially removed in the tar reformer where they are either reformed with steam or partially oxidized. In the case of the high-pressure gasifier, the solids are removed in a ceramic filter, and next the gas is expanded, generating energy. If the indirect lower-pressure gasification is used, the solids are removed together with NH3 in a wet scrubber and compressed. In both cases, traces of hydrocarbons (HBC) and H2S are removed in a multibed pressure swing adsorption (PSA) system. Next, the composition of the gas is adjusted. To accomplish this, three alternatives are considered: water gas shift reactor, bypass, and hybrid membrane/PSA for removal of H2 (with a bed of oxides). The selection depends on the performance of the gasifier, the operation of the of the tar reformer, and the operating conditions of the reactor for minimum production cost. Finally, DME is synthesized and purified. The gas phase is adjusted for the optimal operating conditions using a compressor and heat exchangers. DME is produced using a reactor governed by the equilibrium reactions for methanol and DME synthesis, and water gas shift occurs simultaneously. The unreacted gas is first separated containing CO2 and traces of CO and H2. It can be used in a Brayton cycle to produce energy, or we can recycle it to the reactor. Next, DME, water, and methanol are finally separated by distillation.10,12,15,19 There are a number of trade-offs in the gasifier, in the synthesis of the DME governed by a series of chemical equilibria, the possibility of recycling the unreacted gas or its

(2)

(5) (6) (7) (8)

use to produce power, as well as in the yield of the water gas shift reactor that can be systematically studied by means of an MINLP formulation of the model to optimize the production of DME from biomass. The MINLP is decomposed into four NLP subproblems, one for each gasifier and reforming mode. Heat integration and an economic evaluation provide the operating costs for the different alternatives.

3. PROCESS MODEL The different units involved in the production of DME are modeled using short-cut methods or models that consist primarily of mass and energy balances, design equations, and reduced order models. The flowsheet is written in terms of the total mass flows, component mass flows, component mass fractions, and temperatures of the streams in the network. These are the main variables whose values have to be determined from the optimization. The biomass treatment section is similar to previous papers in which the equations for the modeling of each of the units are detailed.23,24 The components in the system correspond to those in the set J = { Wa, C6H6, DME, CH3OH, tars, CO2, CO, O2, N2, H2, H2S, NH3, CH4, C2H2, C2H4, C2H6, SO2, C, H, O, S, N, olivine, char, ash}. 3.1. Biomass Gasification. 3.1.1. Pretreatment. The incoming feed of lignocellulosic biomass is washed with freshwater at room temperature to remove dirt, sand, and dust using 0.5 kg per kg of biomass. One percent of the water remains with the biomass, while the rest is treated and reused. Wastewater treatment is out of the scope of this work. Next, the biomass is partially dried by means of a mechanical press, removing 90% of the water that goes with the grass. Subsequently, the biomass is reduced into 10 mm pieces for the gasification to be effective, consuming 30 kWh/t.25 3.1.2. Gasification. Direct or indirect gasification is included in the superstructure. 3.1.2.1. Indirect Gasification (Ferco Battelle). The gasifier is fed with biomass, steam, 0.4 kg/kgdry biomass from Src 7, and preheated olivine, 27 kg/kgdry biomass, from the combustor. It operates at 1.6 bar. The composition of the gas produced is determined using eqs 1−8 from Phillips et al.,26 see Table 1. 7466

DOI: 10.1021/acs.iecr.5b00823 Ind. Eng. Chem. Res. 2015, 54, 7465−7475

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Industrial & Engineering Chemistry Research Table 2. Correlation Coefficients of High-Pressure Direct Gasifier molar ratio H2/feed H CO/feed C CO2/feed C CH4/feed C C2H4/feed C C2H6/feed C benzene/feed C naphthalene/feed C solids % of feed N in char % of feed S in char % of feed O in char

A −3.830761 × 10−1 −8.130017 × 10−2 7.157172 × 10−2 1.093589 × 10−2 5.301812 × 10−2 1.029750 × 10−1 4.676833 × 10−2 1.827359 × 10−2 3.36 8.45 1.512040

B 1.894350 −3.340050 3.843454 1.388446 −6.740399 −5.440777 −1.937444 −2.328921

C × × × × × × × ×

10−4 10−4 10−4 10−4 10−5 10−6 10−5 10−6

1.582010 × 10−4

2.666675 2.614482 1.286060 8.812765 −1.372749 −5.350103 −1.270868 −5.951746

× × × × × × × ×

10−4 10−4 10−5 10−4 10−5 10−5 10−5 10−6

−6.972612 × 10−4

D

E

1.060088 × 10−1 1.495730 × 10−1 6.124545 × 10−1 −2.274854 × 10−1 −9.076286 × 10−3 −3.377091 × 10−2 −1.046762 × 10−2 −1.936385 × 10−2

7.880955 × 10−2 −5.268367 × 10−2 9.980868 × 10−2 3.427825 × 10−2 −4.854082 × 10−3 −1.915339 × 10−3 −8.459647 × 10−3 −7.678310 × 10−4

0.1573581

−0.142091

Figure 1. DME synthesis from lignocellulosic raw materials.

result, we perform the mass balance for the gasifier. The stream from the gasifier feeds cyclone 3, see Figure 1, where 99.999% of the char generated is separated from the gas.

Together with the gas, olivine and char exit the gasifier. The char is composed of the carbon that is left after the production of the gases as well as other residues such as oxygen, at least 4% of the oxygen of the biomass, the sulfur and nitrogen that do not generate gases (8.3% and 6.6%, respectively), as well as all the ash in the biomass. The hydrogen, carbon, and oxygen global balance must match the products generated and the equations in Table 1. The solids are separated using a cyclone and recycled to the combustor, while the gas is fed to the reforming stage. In the combustor, char is burned with preheated atmospheric air at 200 °C in an excess of 20%. The energy generated, 25 000 kJ/kgchar,27 reheats the olivine. A makeup of olivine is needed because of losses. In the combustor, the char decomposes, generating SO2, CO2, and N2 and liberating the ash. The stream exiting the combustor contains solids, olivine, and gases. Thus, a cyclone is used to recycle the olivine to the gasifier. The gases are sent to an electrostatic precipitator to remove 99% of the ash and the remaining olivine. Then, this gas stream is used to integrate heat within the rest of the process. Further details for the equations that model this gasifier can be found in the supplementary material of a previous paper.23 3.1.2.2. Direct Gasification, Renugas. The gas composition is computed using the correlations from Eggeman28 and Zhu et al.29 The general form of the correlations is given by eq 9, while the coefficients can be seen in Table 2. On the basis of the

⎛ mol O ⎞ ⎛ mol H O ⎞ 2 2 ⎟⎟ + E⎜⎜ ⎟⎟ fi = A + B ·P(psi) + C·T(°F) + D⎜⎜ ⎝ mol C,feed ⎠ ⎝ mol C,feed ⎠ (9)

3.2. Gas Cleanup and Composition Adjustment. 3.2.1. Gas Cleanup. The gas exiting the gasifier is cleaned up in a three-step process. The gases produced contain light and heavier hydrocarbons whose decomposition will increase the concentration of CO and H2 in the stream. Partial oxidation or steam reforming are considered for that purpose. Next cold or hot cleaning are considered for the removal of solids while a pressure swing adsorption system (MS1-MS2 in Figure 1) using a bed of oxides is used to trap the traces of hydrocarbons and H2S left after reforming. For the first stage, the trade-off between the two reforming options is as follows. Partial oxidation is exothermic and generates energy, but it shows lower yield to hydrogen compared to steam reforming. However, steam reforming is endothermic, whose operation requires energy.30 The reformer can work at low or high pressure.31 Thus, the model remains the same no matter if the feed is coming from either gasifier. HX4 is allocated before the reformers in case a large amount of water or oxygen is needed in the process, reducing the 7467


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condenses in HX6 and is withdrawn from the system before the PSA. 3.2.2. Composition Adjustment. We divide this section into two parts. First, the proper H2-to-CO ratio is obtained. Next, the CO2 is removed. The reason is the generation of CO2 in case water gas shift is used to increase the proportion of hydrogen. For modifying the H2-to-CO ratio, we consider the water gas shift reaction (WGSR) if further hydrogen is to be produced; a bypass, in case there is no need to alter the composition; and finally a membrane-PSA system is considered to purge the excess hydrogen. 3.2.2.1. Water Gas Shift. The reaction taking place in the water shift reactor is widely known:36−39

temperature of the stream from mix4 to any of the reformers. Otherwise, the high operating temperatures and the need for a stream that provides energy at that temperature makes this heat exchanger prohibitive in terms of design and operating cost. Thus, under the latest conditions, HX4 does not exist. 3.2.1.1. Steam Reforming. This technology is modeled based on the mass balances given by the stoichiometry of the reactions taking place. The form of these reactions is given by eqs 10 and 11. CnHm + nH 2O → nCO +

NH3 →

1 3 N2 + H 2 2 2

⎛m ⎞ ⎜ + n⎟H 2 ⎝2 ⎠

(10)

CO + H 2O ↔ CO2 + H 2

(11)

The conversions of the species are as follows: ConvCH4 = 0.8, ConvC6H6 = 1, ConvTar = 1, ConvC2H6 = 0.99, ConvC2H2 = 0.9, ConvC2H4 = 0.9, and ConvNH3 = 0.9 based on Phillips et al.26 We consider that the reactor operates adiabatically so that, based on the endothermic reactions, the final temperature is reduced. The temperature at the outlet is constrained. In the case of the direct gasification, the feed to the filter should be above 300 °C. In any other case, the average temperature at the reformer should be around 600 °C. Thus, HX5 is used to generate steam. 3.2.1.2. Partial Oxidation. We use pure oxygen in stoichiometric proportions for the hydrocarbons to follow reactions in the form of that given by eq 12. n m CnHm + O2 → nCO + H 2 (12) 2 2

(13)

We compute the conversion based on the equilibrium constant given by eq 14 as a function of the feed composition and the operating temperature:36 kp = 10[1910/ T + 273.15 − 1,784] =

PCO2·PH2 PCO·PH2O

(14)

The products of the reactor are calculated based on the stoichiometry of the chemical reaction where the atomic balances are as follows: Carbon: Hydrogen: Oxygen:

The HX4 is used only in the case of using low-pressure gasifier because of the availability of energy at high temperature. The conversions of the different hydrocarbons are assumed to be the same as presented above, based on the experimental results by Vernon et al.32 and Deutschmann and Schmidt.33 The process is exothermic. Once the hydrocarbons are decomposed, solids and NH3 are eliminated. Two alternatives are considered. For high-pressure gasification, hot cleaning is selected, while cold cleaning is used when indirect gasification is in place.34 3.2.1.3. Cold Cleaning. A wet scrubber at 40 °C and 1.2 bar is employed to remove solids (char) and the remaining NH3. Water from the incoming stream condenses. The flow of washing water for scrubbing should be 0.25 kg per m3 of gas,35 and the waste generated is sent to treatment, while the gas exits the scrubber saturated at the operating conditions. The gas stream is compressed to 4.5 bar in a polytropic compressor and cooled to 25 °C to be fed to the PSA system. 3.2.1.4. Hot Cleaning. We use ceramic filters operating at least at 300 °C to remove the solids (char, olivine). Next, the gas is expanded to provide energy for the process assuming polytropic behavior. Finally, in the third cleanup step, the traces of hydrocarbons and H2S are removed. 3.2.1.5. Final HBC and H2S Elimination. For modeling purposes, we assume that a double-column multibed PSA retains all of the hydrocarbons left in the gas stream as well as the ammonia, H2S, and nitrogen. MS1 operates while MS2 is being regenerated for continuous operation. The typical working conditions for PSA systems are low temperature (25 °C) and moderate pressure (4.5 bar).34 A 10% pressure loss is assumed across the bed. Because of the low temperature, water

mol CO + mol CO2|in = mol CO + mol CO2|out 2 ·mol H2 + 2 ·mol H2O|in = 2 ·mol H2 + 2·mol H2O|out mol H2O + mol CO + 2mol CO2|in = mol H2O + mol CO

+ 2·mol CO2|out (15)

The optimization of the entire flowsheet determines the addition of steam from the source, Src15 in Figure 1, as well as the temperature of the reaction, adjusted by means of HX8. 3.2.2.2. Bypass. It may be possible that the stream does not need any adjustment in the CO:H2 ratio due to the cost involved, or simply because the composition is already suitable for the synthesis. Thus, a bypass is also allowed. 3.2.2.3. H2 Membrane/PSA System. The stream to be treated in the membrane/PSA system for the recovery of hydrogen40 has to be adjusted in terms of temperature, 25 °C, and pressure, 4.5 bar. The compression is modeled assuming polytropic behavior to determine the final temperature and energy required. As a result of the cooling, water condenses in HX10. The amount condensed is determined by the saturation conditions of the exiting gas. In this PSA it is assumed that only hydrogen is eliminated from the stream with an efficiency of 100%. The other gases pass through. Finally, all the streams are mixed adiabatically. The second step of the process is the removal of sour gases that poison the catalyst. 3.2.2.4. CO2 Removal by PSA System. The removal of CO2 is the last cleaning stage for the preparation of the syngas to produce DME. The removal of CO2 uses Zeolite 5A or 13X, capable of eliminating 95% of CO2 in the stream and oxygen. Figure 2 shows a scheme of the process. The cycle is short, and the absorption capacity is around 0.1 kg of CO2 per kg of zeolite. The system is modeled as two beds, one operating and the second one in regeneration to allow continuous operation of the plant. The operating conditions are 25 °C and 4.5 bar. 7468


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Industrial & Engineering Chemistry Research CO2 + H 2 ↔ CO + H 2O

(18)

2CH3OH ↔ CH3OCH3 + H 2O

(19)

The reactions are limited by equilibrium constraints. Therefore, there exist optimal operating conditions and feed composition to maximize the production of DME. Thus, the model of the reactor involves the elementary balances given by eq 20 and the equilibrium constants shown in eq 21, with T in K and P in bar: Carbon:

mol CO + mol CO2 + 2 ·molDME + mol CH3OH|in

= mol CO + 2· molDME + mol CH3OH + mol CO2|out Hydrogen:

2 ·mol H2 + 2 ·mol H2O + 6·molDME

+ 4·mol CH3OH|in = 2·mol H2 + 2·mol H2O + 6·molDME + 4·mol CH3OH|out Oxygen:

mol H2O + mol CO + 2 ·mol CO2 + molDME

+ mol CH3OH|in = mol H2O + mol CO + 2·mol CO2 Figure 2. PSA system for the removal of CO2.

+ molDME + mol CH3OH|out (20)

3.3. DME Synthesis and Purification. The detailed flowsheet for the synthesis of DME and its purification can be seen in Figure 3. The direct synthesis of DME from synthesis gas proceeds with methanol as an intermediate. Methanol is typically produced from syngas in the presence of copper-based catalysts (eqs 16 and 17). Together with methanol, the water gas shift (WGS) reaction also occurs (eq 18). The methanol dehydration to DME is carried out on acidic catalysts, such as the widely studied γ-alumina (eq 19).41 CO2 + 3H 2 ↔ CH3OH + H 2O

(16)

CO + 2H 2 ↔ CH3OH

(17)

Kp1 = 10[3066/ T − 10.592] =

PCH3OH 2

PCO·(PH2)

=

nCH3OH ·n Total

2

Pt 2 ·nCO·(n H2)2

⎛ 2835.2 Kp2 = exp⎜ + 1.6775 ln T − 2.39 × 10−4T ⎝ T PDME·PH2O nDME ·n H2O ⎞ − 0.21 × 10−6T 2 − 13.36⎟ = = 2 ⎠ (PCH3OH) (nCH3OH)2 Kp3 = 10[2073/ T − 2.029] =

PCO2·PH2 PCO·PH2O

=

nCO2 ·n H2 nCO·n H2O (21)

Figure 3. Synthesis area and polygeneration. 7469


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Industrial & Engineering Chemistry Research The optimal conditions are variables of our problem to maximize the CO conversion and DME selectivity. In the literature, the reaction temperatures are in the range of 235− 280 °C, and the ratio of H2/CO is typically fed is around 1.42 The stream exiting the reactor is cooled to −40 °C without modifying pressure so that we recover the DME formed together with water and methanol. CO2 gets absorbed in the stream; thus, part of the unreacted CO2 is dragged with the products. The remainder CO2, together with traces of CH3OH, H2, and CO, goes saturated with DME at the pressure and temperature of the gas liquid separation. We use the Henry law43 to compute the absorbed CO2 concentration in liquid phase (eqs 22−27) ⎞ ⎛ −8969.7 Henry = 429.52· exp⎜ ⎟ ⎝ 8.314· (273 + T (Sep3)) ⎠

fc(DME, HX33, Col1) = fc(DME, Col1, HX37) + fc(DME, Col1, Col2)

The following constraint, eq 31, includes the limitation of DME concentration. fc(DME, Col1, HX37) ≤ 0.1 ·fc(DME, HX33, Col1) (31)

The liquid phase comes out of the bottoms as saturated liquid at 10 bar containing primarily water and DME, but also traces of methanol. A global energy balance assuming adiabatic operation of the column is used to compute the energy input to the reboiler:

∑ Ent(J, HX33, Col1) + Q(HX29) = ∑ Ent(J, Col1, Col2) j

(22)

j

+

⎡ fc(CO2 , HX28, Sep3) ·Henry CO2_dissolv⎢ ⎢ (MW(CO2 )· P_Sep3) ⎣

∑ Ent(J, Col1, HX37) + dH_v(Col1, HX37) (32)

j

This liquid phase is fed at 10 bar to the second column where we obtain DME out of the top and mainly water from the bottom, both as saturated liquids at 10 bar. In this analysis, we do not include water treatment neither in process design nor within the economic evaluation. 3.4. Power Generation. The stream separated after the condensation in Sep3 can follow two options: (a) The unreacted stream is used in a Brayton cycle to generate energy. We burn the CO and H2 in the gas from Sep 3 and feed it to a turbine to be expanded:

⎛ fc(Wa, HX28, Sep3) fc(MetOH, HX28, Sep3) + ⎜⎜ + ρH O ρCH OH ⎝ 2 3 ⎤ ⎞ fc(DME, HX28, Sep3) ⎟⎥ fc(CO2 , HX28, Sep3) + = ⎟ ⎥ MW(CO2 ) ρDME ⎠⎦ (23)

545

(30)

⎛ fc(CO2 , Sep3, HX33) fc(Wa, Sep3, HX33) = CO2_dissolv⎜⎜ MW(CO2 ) ρH O ⎝

CO +

1 O2 → CO2 2

(33)

H2 +

1 O2 → H 2O 2

(34)

2

⎞ fc(MetOH, HX28, Sep3) fc(DME, Sep3, HX33) ⎟ + + ⎟ ρCH OH ρDME ⎠ 3

(b) It can be recycled. We have to remove the CO2; thus, the recycle is sent to the CO2 capture section of the facility so that the H2 and CO come back as synthesis gas.

(24)

(CO2 , Sep3, HX33) = fc(CO2 , HX28, Sep3) − fc(CO2 , Sep3, Spl2)

4. SOLUTION PROCEDURE To extract the optimal flowsheet from the superstructure as well as to determine the optimal operating conditions, a total enumeration of alternatives has been applied because of the small number of 0−1 variables. Simple enumeration is justified to simplify the complexity of the complete superstructure model, reduce its size, and address the optimal operating conditions at the different units. Thus, a series of NLPs are solved instead of a full MINLP. Therefore, the MINLP is decomposed into four subproblems, one for each gasifier and one for each reforming mode. For each one of the subproblems, we compare the production of DME alone with the simultaneous production of DME and power from the unreacted gas. The optimization of the operating conditions is obtained by solving an NLP with GAMS/CONOPT. This yields the operating temperatures of the gasifier and the combustor, in the case of the indirect gasifier, as well as the H2to-CO ratio before entering the reactor and the operating pressure and temperature of the synthesis. This approach allows consideration of the various variables that affect the gas composition from the gasifier to the reactor in an integrated way, unlike the use of process simulators. The objective function (Z) of each of the subproblems is given by eq 35, which is a simplified operating cost considering the production

(25)

⎛ MW(DME) ⎞ P_vapDME(Sep3) y_DME = ⎜ ⎟ ⎝ MW(CO2 ) ⎠ P_Sep3 − P_vapDME(Sep3) (26)

fc(DME, Sep3, Spl2) = y_DME·fc(CO2 , Sep3, Spl2) (27)

The unreacted gases of the outlet stream of the synthesis reactor are recycled, and the liquid phase is distilled to separate methanol from DME. This liquid phase is expanded and distilled in a stripping column to remove the CO2 from the product stream. Thus, it is fed as saturated liquid at 10 bar. The top of the columns is a gas phase that contains CO2 saturated with DME at 10 bar. This stream is recycled; nevertheless, we do not allow losing more than 10% of the initial amount. y_satDME =

P_vapDME(Col1) MW(DME) MW(CO2 ) (P_Col1 − P_vapDME(Col1)) (28)

y_satDME·fc(CO2 , Col1, HX37) = fc(DME, Col1, HX37) (29) 7470


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Figure 4. Objective function of the different alternatives. (F, Ferco; S, steam reforming; R, Renugas; O, partial oxidation).

Table 3. Major Optimal Operating Conditions: DME Production Alone

of hydrogen and the cost of energy and raw material. CH2 = $1.6/kg, CO2 = $0.021/kg, CElectr = $0.06/kW, Csteam = $0.019/ kg.23,44 For the base case, we assume CDME to be $1/kg. Later in the paper, a sensitivity analysis on the prices of DME and power is performed. Each problem involves 3000 equations and 3500 variables. We use 1555 t/d of switchgrass feed to the system, a typical size of biomass processing plants.1

Ferco gasifier combustor gas mixing DME synthesis reactor yield

Z = fc(DME)C DME − C Electr ∑ Wturbines − C Electr ∑ WCompressors − CSteam(

∑

Q−

Consumed

+ fc(H 2)C H2 − C biomass·fc(biomass) −

∑

P (bar)

908 984

1.6 1

H2/CO

kgDME/kgBiomass

kgH2/kgBiomass

0.35

0.017

1 250

50

Q)

5.2.2. DME and Power Production. In case we prefer a polygeneration plant, it is achieved by using the unreacted gases for power production. The optimal topology for the polygeneration system also considers the use of Ferco gasifier followed by steam reforming and the production of DME. The most important operating parameters of the major units are presented in Table 4. In this case, hydrogen is no longer a byproduct of the system.

generated

∑

T (°C)

fc(O2 )CO2

Gasifier,Reformer

(35)

Once the topology is determined, we design a heat exchanger network. The excess of energy within the process is used to produce steam. Finally, a detailed economic evaluation of the optimal process is performed.

Table 4. Major Optimal Operating Conditions: DME and Electricity Production

5. RESULTS 5.1. Topology Determination. We have eight alternative topologies related to the use of direct or indirect gasification, steam or partial oxidation, and the use of the syngas for the production of DME or to produce electricity. In Figure 4 we present the objective function, Z ($/s), for each option to compare instead of providing just an optimal value and topology. We can see that the use of indirect gasification is suggested as well as the use of steam reforming. Thus, we based further comparisons and sensitivity analysis based on a topology that uses the Ferco gasifier and steam reforming. 5.2. Operating Results. 5.2.1. DME Production. The topology optimization reveals that the best configuration involves the use of Ferco gasifier followed by steam reforming and the production of DME recycling the unreacted gases. The most important operating parameters of the major units are presented in Table 3. The optimal results validate the experimental studies by Ohno that claimed an optimal ratio of H2 to CO of around 1.42 Moreover, an excess of hydrogen is also produced. These results are interesting because the production of bioethanol from syngas also uses the same H2to-CO ratio;20 therefore, our plant can be flexible to produce either DME or bioethanol with the purified syngas.

Ferco gasifier combustor gas mixing DME synthesis reactor yield

T (°C)

P (bar)

908 984

1.6 1

H2/ CO

kgDME/kgBiomass

MW/kgBiomass

0.15

2.9

1.22 250

50

5.2.3. Sensitivity Analysis. The prices for DME and, in particular, power are volatile. Therefore, to determine the mode of operation of the facility, a sensitivity analysis is performed for the optimal gasifier and reforming mode over a range of processes for DME from $0.25/kg to $0.75/kg and for power, from $0.05/kWh to $0.15/kWh. Figure 5 shows the limit in prices for both products. Below the continuous line, the production of DME alone is suggested. For higher prices of electricity a more flexible facility is recommended. However, beyond a certain electricity price, only power should be produced. In Figures 6 and 7 we can see that, as the price of electricity increases, the production of DME is less profitable. Moreover, 7471


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20%, and 10% of the equipment cost, respectively. The fees represent 3% of the fix cost; other administrative expenses and overhead and the plant layout represent 10% of the direct costs (fees plus fix capital) and 5% of the fix cost, respectively. The ground cost is $6 million, and the plant start-up cost represents 15% of the investment. For the production of DME alone, the investment including start-up costs adds up to $133 million for the production of 197 kt/yr of DME. Biomass processing to obtain the syngas represents 58%; compressors are the second highest contribution to the units (see Figure 8a). When electricity is produced alone, the investment cost increases up to $176 million for the production of 105 MW because of the high cost of the turbine. In this case, the compressors and turbine represent more than 55% of the investment in equipment (see Figure 8b). Furthermore, we estimate the production cost of the DME and electricity considering the hydrogen as a credit. For the average annual cost, we consider the labor costs (2.25% of investment for the DME plant, and that value is maintained in the case of the power plant), amortization (linear with time in 20 years), administration and overhead 2% of investment cost, and we pay for the cooling water. Chemicals correspond to $1.2 million, including the need for refrigeration after the reactor to recover the unreacted gases, minus the credit due to the hydrogen. The raw material is considered at $60/t. In the case of producing DME, we also produce 9.6 kt/yr of hydrogen, which provides credit at $1.6/kg. The contribution of the different items to the production cost can be seen in Figure 9a, and the production cost of DME is $0.31/kg. In the case of the power plant, we produce 105 MW at $0.07/kWh, which is competitive with the fossil fuel-based electricity. In this case, no hydrogen is produced; however, we obtain steam, reducing the needs for utilities (Figure 9b).

Figure 5. Major products as a function of DME cost and electricity cost.

for each DME price there is a correspondent price for power beyond which it is better to use the biomass to produce electricity alone. Not even hydrogen will be produced from that point onward. Furthermore, there is a combination of electricity and DME prices at which the production of DME alone and the use of a polygeneration system have the same objective function. In Figures 6 and 7, search for points in the same vertical line. Another criterion should be used to decide on the final products of the process. This point also is shifted to the right with the cost of electricity. Note that in Figure 7, a negative sign for power production indicates generation out of the system. 5.3. Economic Evaluation. In this section, we present the investment and production costs for the production of DME alone, by recycling the unreacted gases, and for the biomassbased power plant with no production of DME as the two extreme cases. For the evaluation of the investment cost,45 we estimate the cost of the units using www.matche.com46 based on the sizing given by the flows through the units. The installed equipment is assumed to represent 1.5 times the equipment cost. Piping, isolation, instrumentation, and utilities represent 20%, 15%,

6. CONCLUSIONS We have evaluated the production of dimethyl ether from biomass following a direct synthesis path using a mathematical optimization approach. We compare the use of two gasifiers, two reforming modes, and the production of DME where the unreacted gas is either recycled or used to produce electricity in a polygeneration system.

Figure 6. DME production levels at different prices of electricity and DME. 7472


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Figure 7. Electricity production levels at different prices of electricity and DME.

Figure 8. Breakdown of equipment costs: (a) DME production and (b) biomass-based power plant.

Figure 9. Production cost breakdown: (a) DME production and (b) biomass-based power plant.

The optimization of the superstructure yields the optimal topology. It involves the use of Ferco gasification and steam reforming. At current prices for DME and electricity, it is more interesting to recycle the unreacted gases for further production of DME. The production of electricity is favored for electricity prices above $0.09/kWh. The economic evaluation reveals that DME can be produced competitively at $0.31/kg when only DME is produced. Apart from DME, hydrogen is also obtained from such a facility, providing an interesting asset. On the other hand, we can

produce electricity alone. The investment cost is 40% higher, and no hydrogen is produced. However, the electricity is generated at a competitive price of $0.07/kWh.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: +34923294479. E-mail: [email protected]. Notes

The authors declare no competing financial interest. 7473


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(4) Li, Z.; Liu, P.; He, F.; Wang, M.; Pistikopoulos, E. N. Simulation and exergoeconomic analysis of a dual-gas sourced polygeneration process with integrated methanol/DME/DMC catalytic synthesis Comp. Comput. Chem. Eng. 2011, 35, 1857−1862. (5) Martín, M.; Grossmann, I. E. ASI: Toward the Optimal Integrated Production of Biodiesel with Internal Recycling of Methanol Produced from Glycerol. Environ. Prog. Sustainable Energy 2013, 32 (4), 891−901. (6) Martín, M.; Grossmann, I. E. Design of an optimal process for enhanced production of bioethanol and biodiesel from algae oil via glycerol fermentation. Appl. Energy 2014, 135, 108−114. (7) Martín, M.; Grossmann, I. E. Simultaneous dynamic optimization and heat integration for the co-production of diesel substitutes: Biodiesel (FAME & FAEE) and glycerol ethers from algae oil. Ind. Eng. Chem. Res. 2014, 53, 11371−11383. (8) Lee, C.-J.; Lim, Y.; Kim, H.-S.; Han, C. Optimal Gas-To-Liquid Product Selection from Natural Gas under Uncertain Price Scenarios. Ind. Eng. Chem. Res. 2009, 48, 794−800. (9) Zhou, L.; Hu, S.; Chen, D.; Li, Y.; Zhu, B.; Jin, Y. Study on Systems Based on Coal and Natural Gas for Producing Dimethyl Ether. Ind. Eng. Chem. Res. 2009, 48, 4101−4108. (10) Clausen, L. R.; Elmegaard, B.; Ahrenfeldt, J.; Henriksen, U. Thermodynamic analysis of small-scale dimethyl ether (DME) and methanol plants based on the efficient two-stage gasifier. Energy 2011, 36, 5805−5814. (11) Zhou, L.; Hu, S.; Li, Y.; Zhou, Q. Study on co-feed and coproduction system based on coal and natural gas for producing DME and electricity. Chem. Eng. J. 2008, 136, 31−40. (12) Li, F.; Zeng, L.; Fan, L. S. Techno-Economic Analysis of CoalBased Hydrogen and Electricity Cogeneration Processes with CO2 Capture. Ind. Eng. Chem. Res. 2010, 49 (21), 11018−11028. (13) Haro, P.; Ollero, P.; Villanueva Perales, A. L.; Reyes Valle, C. Technoeconomic assessment of lignocellulosic ethanol production via DME (dimethyl ether) hydrocarbonylation. Energy 2012, 44, 891− 901. (14) Haro, P.; Ollero, P.; Villanueva Perales, A. L.; Gómez-Barea, A. Thermochemical biorefinery based on dimethyl ether as intermediate: Technoeconomic assessment Applied. Appl. Energy 2013, 102, 950− 961. (15) Pascall, A.; Adams, T. A. Semicontinuous separation of dimethyl ether (DME) produced from Biomass. Can. J. Chem. Eng. 2013, 91, 1001−1021. (16) Ju, F.; Chen, H.; Ding, X.; Yang, H.; Wang, X.; Zhang, S.; Dai, Z. Biotechnol. Adv. 2009, 27 (5), 599−605. (17) Lee, S.; Gogate, M. R. E.; Kulik, C. J. A novel single step Dimethyl ether (DME) synthesis in a three phase slurry reactor from CO-Ric Syngas. Chem. Eng. Sci. 1992, 47 (13/14), 3769−3776. (18) Cavalcanti, F. A. P.; Stakheev, A.; Yu; Sachtler, W. M.H. Direct synthesis of methanol, dimethyl ether and paraffins from syngas over Pf/zeolite Y catalysts. J. Catal. 1992, 134, 226−241. (19) Ogawa, T.; Inoue, N.; Shikada, T.; Ohno, Y. Direct Dimethyl Ether Synthesis. J. Nat. Gas Chem. 2003, 12, 219−227. (20) Yotaro, O.; Masahiro, Y.; Tsutomu, S.; Osamu, I.; Takashi, O.; Norio, I. New direct synthesis technology for DME (Dimethyl ether) and its application technology. JFE GIHO 2004, 6, 70−75. (21) Zhu, Y.; Wang, S.; Ge, X.; Liu, Q.; Luo, Z.; Cen, K. Experimental study of improved two step synthesis for DME production. Fuel Process. Technol. 2010, 91, 424−429. (22) Azizi, Z.; Rezaeimanesh, M.; Tohidian, T.; Rahimpour, M. R. Dimethyl ether: A review of technologies and production challenges. Chem. Eng. Process. 2014, 82, 150−172. (23) Martín, M.; Grossmann, I. E. Energy Optimization of Bioethanol Production via Gasification of Switchgrass. AIChE J. 2011, 57 (12), 3408−3428. (24) Vidal, M.; Martín, M. Optimal coupling of biomass and solar energy for the production of electricity and chemicals. Comput. Chem. Eng. 2015, 72, 273−283.

ACKNOWLEDGMENTS The authors acknowledge the TCUE fellowship from Fundación Universidad − Empresa USAL 2013 to E.P. NOMENCLATURE CO2_dissolved = Amount of CO2 dissolved in the liquid phase (kmol/m3) Ci = Cost of item i ($/kg) or ($/kWh) DryMass = Flow of dry biomass (kg/s) Ent(J, unit, unit) = Flow enthalpy of component J from unit to unit (kJ/s) F (unit, unit) = Total flow from unit to unit (kg/s) fc(J) = Final product J flow rate (kg/s) fc (J, unit, unit) = Flow of component J from unit to unit (kg/s) Kp = Equilibrium constant Henry = Henry constant (atm m3/kmol) Moli = molar flow of component i (kmol/s) MW(J) = Molecular weight of component J (kg/kmol) Pi = Partial pressure (atm) P(i) = Total pressure at unit i (atm) P_vap = Vapor pressure of component (atm) Q(unit) = Thermal energy involved in unit (kW) T(i) = Temperature at unit i (°C unless otherwise specified) TFercogas = Operating temperature of the Ferco gasifier (°F) y_DME = kg of DME per kg of gas W(unit) = Power involved in unit (kW) Z = Objective function ($/s)

Units

Col = Columns Compress = Compressor Cyc = Cyclon Expand = Turbine expander Flash = Flash separator HX = Heat Eexchanger Mec Pres = Mechanical press Mix = Mixer MS = Molecular sieve Sep = Phase separator Snk = Sink Spl = Splitter Src = Source Subindexes

O2 = oxygen CO = carbon monoxide CO2 = carbon dioxide MetOH = methanol H2 = hydrogen Wa = water Steam = Steam Electr = electricity

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REFERENCES

(1) Martín, M.; Grossmann, I. E. On the systematic synthesis of sustainable biorefineries. Ind. Eng. Chem. Res. 2013, 52 (9), 3044− 3064. (2) Baliban, R. C.; Elia, J. A.; Floudas, C. A.; Gurau, B.; Weingarten, M. B.; Klotz, S. D. Hardwood Biomass to Gasoline, Diesel, and Jet Fuel: 1. Process Synthesis and Global Optimization of a Thermochemical Refinery. Energy Fuels 2013, 27 (8), 4302−4324. (3) Baliban, R. C.; Elia, J. A.; Floudas, C. A. Biomass to liquid transportation fuels (BTL) systems: process synthesis and global optimization framework. Energy Environ. Sci. 2013, 6 (1), 267−287. 7474


Article

Industrial & Engineering Chemistry Research (25) Mani, S.; Tabil, L. G.; Sokhansanj, S. Grinding performance and physical properties of wheat and barley straws, corn stover and switchgrass. Biomass Bioenergy 2004, 27, 339−352. (26) Phillips, S.; Aden, A.; Jechura, J., Dayton. D.; Eggeman, T. Thermochemical Ethanol via Indirect Gasification and Mixed Alcohol Synthesis of Lignocellulosic Biomass; NREL/TP-510-41168; National Renewable Energy Laboratory: Golden, CO, April 2007. (27) Di Blasi, C. Modeling wood gasification in a countercurrent fixed-bed reactor. AIChE J. 2004, 50 (9), 2306−2319. (28) Eggeman, T. Updated Correlations for GTI Gasifier − WDYLD8. Technical memorandum for Pam Spath; National Renewable Energy Laboratory: Golden, CO. June 27, 2005. (29) Zhu, Y.; Gerber, M. A.; Jones, S. B.; Stevens, D. J. Analysis of the effects of compositional and configurational assumptions on product costs for the thermochemical conversion of lignocellulosic biomass to mixed alcohols- FY 2007 Progress Report; US DOE; PNNL-17949, 2009. (30) Rand, D. A. J.; Dell, R. M. Hydrogen Energy Challenges and Prospects; The Royal Society of Chemistry, Thomas Graham House: Cambridge, UK, 2008; ISBN: 978-0-85404-597-6. (31) Brenes, M. D. Biomass and bioenergy. Nova Science Publishers: New York, 2006; ISBN-13: 9781594548659. (32) Vernon, P. D. F.; Green, M. L. H.; Cheetham, A. K.; Ashcroft, A. T. Partial oxidation of methane to synthesis gas. Catal. Lett. 1990, 6 (2), 181−186. (33) Deutschmann, O.; Schmidt, L. D. Two-dimensional modeling of partial oxidation of methane on rhodium in a short contact time reactor. Symp. (Int.) Combust., [Proc.] 1998, 27 (2), 2283−2291. (34) Olofsson, I.; Nordin, A.; Söderlind, U. Initial Review and Evaluation of Process Technologies and Systems Suitable for Cost-Efficient Medium-Scale Gasification for Biomass to Liquid Fuels; ETPC Report 05-02, 2005; ISSN: 1653-0551. (35) Martelli, E.; Kreutz, T.; Consonni, S. Comparison of coal IGCC with and without CO2 capture and storage: Shell gasification with standard vs. partial water quench. Energy Procedia 2009, 1, 607−614. (36) Roh, H.-S.; Lee, D. K.; Koo, K. Y.; Jung, U. H.; Yoon, W. L. Natural gas steam reforming for hydrogen production over metal monolith catalyst with efficient heat-transfer. Int. J. Hydrogen Energy 2010, 35 (3), 1613−1619. (37) Killmeyer, R.; Rothenberger, K.; Howard, B.; Ciocco, M.; Morreale, B.; Enick,R.; Bustamante, F. Water-Gas Shift Membrane Reactor Studies. Hydrogen, Fuel Cells, and Infrastructure Technologies FY 2003 Progress Report, 2003; http://www1.eere.energy.gov/ hydrogenandfuelcells/pdfs/iia9_killmeyer.pdf (Accessed February 2011). (38) Ji, P.; Feng, W.; Chen, B. Production of ultrapure hydrogen from biomass gasification with air. Chem. Eng. Sci. 2009, 64, 582−592. (39) Ji, P.; Feng, W.; Chen, B. Comprehensive Simulation of an Intensified Process for H2 Production from Steam Gasification of Biomass. Ind. Eng. Chem. Res. 2009, 48, 3909−3920. (40) IST World. http://www.ist-world.org. (41) McBride, K.; Turek, T.; Güttel, R. Direct Dimethyl Ether Synthesis by Spatial Patterned Catalyst Arrangement: A Modeling and Simulation Study. AIChE J. 2012, 58 (11), 3468−3473. (42) Ohno, Y. Proceedings of the Japan DME Forum Workshop, Tokyo, 2002; pp 113−122. (43) Papari, S.; Kazemeini, M.; Fattahi, M. Mathemathical Modeling of a Slurry Reactor for DME direct synthesis from syngas. J. Nat. Gas Chem. 2012, 21 (2), 148−157. (44) Smith, R.; Varbanov, P. What’s the price of Steam? CEP 2005, July, 29−33. (45) Sinnott, R.; Towler, G. Chemical Engineering Design; Elsevier Ltd.: Oxford, 2009; ISBN: 978-84-291-7199-0. (46) Matches (2014). www.matche.com.

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