Efficiency estimation and improvement of the 1,3-butadiene production

These processes consisted of lignin gasification, conversion of syngas to ... 280 °C were favorite conditions for improving the most promising proces...
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Efficiency Estimation and Improvement of the 1,3-Butadiene Production Process from Lignin via Syngas through Process Simulation Toshiaki Hanaoka, Shinji Fujimoto,* and Masaru Yoshida Bio-based Materials Chemistry Group, Research Institute for Sustainable Chemistry, National Institute of Advanced Industrial Science and Technology 3-11-32, Kagamiyama, Higashihiroshima, Hiroshima 739-0046, Japan S Supporting Information *

ABSTRACT: Three processes for the production of 1,3-butadiene (1,3-BD) from lignin via syngas were proposed, and their 1,3BD yields and input energy, such as electric power and heat loads, were estimated through process simulation. These processes consisted of lignin gasification, conversion of syngas to light olefins (LOs) via (1) dimethyl ether (DME), (2) methanol, or (3) direct synthesis, and isomerization/dehydrogenation of n-C4H8. The process capacity was 200 t/d on a wet lignin basis. The electric power was largely dependent on the process (4777−6073 kWe), while the minimum external heat was 97 kW, according to pinch analysis. When each reaction proceeded ideally, the process featuring the conversion of syngas to LOs via DME was the most promising. The high electric power (6008 kWe) for the process was attributed to excess N2 production through a cryogenic air separation method. A decrease in the amount of N2 supplied to the DME-to-LOs unit led to a decrease in the electric power to 5381 kWe, and the 1,3-BD yield increased to 14.2 wt %. In the DME-to-LOs step, the feed gas with >8.7% DME and a reaction temperature of ∼280 °C were favorite conditions for further improving this most promising process.

1. INTRODUCTION 1,3-Butadiene (1,3-BD) is a raw material for various synthetic rubbers and plastics such as polybutadiene (PB), styrene− butadiene rubber (SBR), and acrylonitrile butadiene styrene (ABS), and the demand for it is expected to increase. 1,3-BD has been produced commercially as a byproduct of ethylene (C2H4) production through the steam cracking of naphtha. In recent years, shale gas has been employed as a feedstock, because it is less expensive than naphtha. Therefore, the amount of C2H4 produced from naphtha will decrease, leading to a decrease in 1,3-BD production in the future. In order to solve this dilemma, researchers have recently focused on 1,3BD production from alternative feedstocks to naphtha. Biomass is the only renewable energy resource that can be converted to materials; it mainly consists of cellulose, hemicellulose, and lignin. In biorefineries, the components obtained through pretreatment methods are converted to building block chemicals employed in the petrochemical industry.1−6 Conversions of cellulose and hemicellulose to chemicals such as ethanol,7,8 lactic acid,9,10 levulinic acid,11,12 and alginic acid13 have been reported. In contrast, lignin is converted to mixtures such as gasoline range products and additives for resin,1 and conversion of lignin to pure chemicals has rarely been reported.14,15 The main reason is that most attempts involve the utilization of specific structures in lignin. Torres and de Jong have reported that biomass is gasified to syngas, and then is converted to light olefins (LOs).16 Gasification technology can produce low-molecular-weight compounds such as syngas from feedstocks thermochemically. Therefore, it has the advantage of being less restricted than pretreatment methods, with respect to the feedstock structure. If n-C4H8 such as 1-C4H8, cis-2-C4H8, and trans-2-C4H8 can be obtained from lignin via syngas effectively, then a promising © XXXX American Chemical Society

process for 1,3-BD production could be realized, because 1,3BD is produced commercially from n-C4H8 through isomerization/dehydrogenation (Iso/Dehyd).17 The possible routes to LOs from syngas are conversion via methanol (MeOH), lower alcohols, wax, and dimethyl ether (DME), and direct synthesis.16 At present, MeOH-to-olefin (MTO) technology is more advanced than that for the conversion of lower alcohols to olefins. In contrast, selective n-C4H8 synthesis through the catalytic cracking of wax is difficult, because it is similar to the steam cracking of naphtha. Therefore, in the future, three processes for the production of 1,3-BD from lignin via syngas are expected: (1) LO synthesis from syngas via DME, (2) LO synthesis from syngas via MeOH, and (3) direct synthesis of LOs from syngas, as shown in Figure 1. However, it is not clear which processes give high 1,3-BD yields and which unit operations are the top priorities for the realization of the processes. Therefore, the product yield and input energy based on mass balance and energy balance must be taken into account when investigating the feasibility of such new processes. Pinch analysis is used when designing a heat recovery network for minimizing the input heat energy.18,19 The minimum external heat (QH,min) can be estimated through the analysis. Therefore, it has been actually used to estimate the heat of the proposed processes using biomass as a feedstock.20−25 In the present study, the three 1,3-BD production processes shown in Figure 1 were simulated to determine which is the most promising and which unit operation to investigate, taking Received: July 30, 2017 Revised: September 26, 2017 Published: October 11, 2017 A

DOI: 10.1021/acs.energyfuels.7b02237 Energy Fuels XXXX, XXX, XXX−XXX

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Figure 1. Proposed route for conversion of lignin to 1,3-BD via syngas.

Figure 2. Unit operations and intermediates of 1,3-BD production processes simulated in the present study: (a) Syn-DME, (b) Syn-MeOH, and (c) Syn-LOs.

because the product selectivity followed the Anderson− Schulz−Flory (ASF) distribution. Therefore, for the byproducts in Syn-LOs, both C2H4 and C3H6 are considered. For each process, both actual and ideal cases were taken into account. For the former case, it was assumed that unit operations reported in previous studies could be simply combined (denoted as SC). For the latter case, it was assumed that chemical reactions in each unit operation proceeded ideally. For the ideal cases for Syn-DME and Syn-MeOH, it was assumed that unit operations in which chemical reactions reached the equilibrium were combined theoretically (denoted as TH). In contrast, in the direct synthesis of LOs from syngas, the product selectivity generally followed the ASF distribution, which is different from the equilibrium composition. The nC4H8 yields were extremely low in some studies, although the product selectivity did not follow the ASF distribution.36−38 Therefore, for the ideal case for Syn-LOs, it was assumed that the process featured the “improved” direct synthesis of LOs step in which the maximum n-C4H8 yield was obtained according to ASF distribution (denoted as Imp). Therefore, the processes are named differently based on the case. For example, the SC case of Syn-DME is denoted as SC-Syn-DME. The ideal cases of Syn-DME and Syn-LOs are denoted as TH-Syn-DME and Imp-Syn-LOs, respectively. A catalyst is required in some unit operations: DME synthesis, conversion of DME to LOs (DME-to-LOs shown in Figure 2a), MeOH synthesis, conversion of MeOH to LOs (MeOH-to-LOs shown in Figure 2b), direct synthesis of LOs from syngas (Figure 2c), and Iso/Dehyd (Figures 2a−2c). Comparison of the yields of the intermediates and 1,3-BD for the SC and TH (or Imp) cases can reveal not only the difference between the actual and ideal catalytic performances,

into account the income from selling 1,3-BD and byproducts and the expenses for electric power, and QH,min by pinch analysis. Then, from the point of view of profitability, we improved the most promising process while maintaining the 1,3-BD yield. Finally, desirable reaction conditions in the toppriority unit operation were discussed in order to decrease the electric power of the entire process.

2. PROCESS DESIGN AND SIMULATION 2.1. Unit Operations and Intermediates. Figure 2 shows the unit operations and intermediates in the three processes simulated in the present study. These processes had different unit operations in the syngas-to-LOs synthesis step, while the lignin gasification and the Iso/Dehyd of n-C4H8 were the same unit operations. In the first process, in Figure 2a, the syngas obtained in the gasification step is converted to DME, followed by LO synthesis. This process is denoted as Syn-DME. In the second process, in Figure 2b, the H2/CO/CO2 ratio of the syngas is adjusted, and then MeOH is produced, followed by LO synthesis. This process is denoted as Syn-MeOH. In the third process, in Figure 2c, the syngas is converted to LOs directly. This process is denoted as Syn-LOs. These three processes gave different byproducts (Figure 2). In previous studies on the LO synthesis from DME26−31 and MeOH,32−35 the C2H4 yields were very low, although olefins with carbon numbers of 2−4 were obtained as products. In the future, the selling price of C2H4 is expected to decrease, because the amount of C2H4 derived from shale gas will increase. Therefore, for the byproducts in Syn-DME and Syn-MeOH, only C3H6 is considered. In the direct synthesis of LOs, the C2H4 yield was higher than the C3H6 and n-C4H8 yields, B

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Figure 3. Schematic flow diagram for Syn-DME.

plant had the capacity to process 1000 N m3/h of syngas and capture ∼24 t-CO2/d.43 The absorption liquid was reused after CO2 was released in the regeneration tower. In the present study, when CO2 removal was required, CO2 absorption and MDEA regeneration units were adopted (see Figure 3, as well as Figures S1 and S2). 2.2.2. LO Synthesis from Syngas in Syn-DME. The DME synthesis unit was designed based on an international patent44 (Figure 3). A slurry bed reactor that uses Cu−Zn and γ-Al2O3 as catalysts was employed as the reactor. The Cu−Zn promotes not only methanol synthesis (CO + 2H2 → CH3OH), but also the water-gas-shift (WGS) reaction (CO + H2O → CO2 + H2). The dehydration of methanol (2CH3OH → CH3OCH3 + H2O) is promoted by γ-Al2O3. Therefore, the synthesis of DME from syngas (3CO + 3H2 → CH3OCH3 + CO2) is favorable at higher pressures. The syngas was pressurized and heated to the desired pressure and temperature using three compressors and three HXs, and then supplied to a reactor. The product mixture leaving the reactor was cooled using the HX, and then separated into the gas and liquid phases. Both phases were depressurized to the atmospheric pressure using expanders. The liquid phase was a mixture of DME and MeOH. The MeOH content was very low; however, MeOH was also theoretically converted to LOs in the following step. Therefore, the mixture was supplied to the DME-to-LOs unit after being mixed with N2. The gas phase was mixed with air, and then combusted and cooled. Then, it was released to the air as exhaust. The DME-to-LOs unit was designed based on a report by Park et al.30 A downdraft fixed-bed reactor that uses MCM-68 as the catalyst was employed as the reactor in the present study, since it exhibited high n-C4H8 selectivity, although the objective

but also which unit operation should be investigated to decrease the input energy for the entire process. 2.2. Equipment and Configuration. Figure 3 shows the schematic flow diagram for Syn-DME as a typical diagram. Those for Syn-MeOH and Syn-LOs are shown in Figures S1 and S2 in the Supporting Information. For Syn-DME and SynMeOH, the equipment employed in the Iso/Dehyd unit for the SC case was different from that for the TH case (see Figure 3, as well as Figure S1). In contrast, Syn-LOs used the same equipment for both cases (Figure S2). 2.2.1. Gasification. Syngas production has been investigated using various types of gasifiers.39−42 The gasification step in the present study is required to gasify various feedstocks, prevent the generation of unwelcome CH4, and obtain syngas with high CO and H2 contents. Therefore, the gasification unit was designed based on the Thermoselect process for producing syngas from waste biomass.42 The process has been applied commercially to produce electricity at a feedstock capacity of 300 t/d in Chiba Prefecture in Japan.42 The nonpretreated feedstock can be gasified with pure oxygen and steam, and then the product gas is recovered through reforming, quenching, and refining.42 The designed unit was used for all processes (see Figure 3, as well as Figures S1 and S2). Both an O2/N2 mixture gas with a high O2 content and steam were employed as gasifying agents in the present study. The former was produced using a cryogenic air separation method, while the latter was moisture employed in the feedstock without additional steam. The syngas obtained in the gasification unit was cooled using a heat exchanger (HX). CO2 in cleaned syngas was removed by an absorption tower after impurities were removed by a scrubber. The chemical absorption method using MDEA (methyl diethanol amine) as an absorption liquid was employed for CO2 absorption.43 The CO2 separation and capture pilot C

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Energy & Fuels was propylene production from DME.30 The feedstock mixture was heated to the desired temperature and then supplied to the reactor. The product mixture leaving the reactor was cooled, and then separated into the gas and liquid phases. The gas phase, which was a mixture of N2 and CO2, was depressurized, and then released to the air as exhaust 1. The liquid phase was pressurized and cooled, and then supplied to the first distillation tower. The distillate in which C3H6 was dominant was collected in the tower and was depressurized. C3H6 then was obtained as a byproduct after CO2 was removed by the absorption tower. The distillate in which n-C4H8 was dominant was collected in the second distillation tower. It was cooled and depressurized after unreacted MeOH and DME were removed by the absorption towers. The hydrocarbon mixture containing n-C4H8 was supplied to the Iso/Dehyd unit. The bottom liquid collected in the second distillation tower was depressurized, and then mixed with air and combusted. It was then cooled and released to the air as exhaust 2. 2.2.3. LO Synthesis from Syngas in Syn-MeOH. In the H2/ CO/CO2 adjustment unit (Figure S1), the syngas was mixed with H2O and then supplied to a reactor after heating. The WGS reaction is conducted commercially, using a fixed-bed reactor packed with Cu−Zn or Fe−Cr catalysts. The product gas leaving the reactor was cooled. Then, the H2/CO/CO2 ratio was adjusted to the desired ratio for the MeOH synthesis by removing CO2. The MeOH synthesis unit was designed based on the operation conditions for a commercial plant45 (Figure S1). Several types of reactors that use Cu−Zn as the catalyst are generally employed. An isothermal reactor was employed as the reactor to recover reaction heat in the present study. The syngas was pressurized and heated to the desired pressure and temperature using three compressors and three HXs, and then supplied to the reactor. The product mixture leaving the reactor was cooled and separated into the gas and liquid phases. Both phases then were depressurized. The gas phase was mixed with air and then combusted and cooled. Then, it was released to the air as exhaust. The liquid phase was mixed with N2, and then supplied to the MeOH-to-LOs unit. The MeOH-to-LOs unit was designed based on a report by Kumita et al.33 A downdraft fixed-bed reactor that uses ZSM-58 as the catalyst was employed as the reactor in the present study. The feedstock mixture was heated to the desired temperature and then supplied to the reactor. The product mixture leaving the reactor was cooled and separated into the gas and liquid phases. The gas phase in which N2 was dominant was depressurized and then released to the air as exhaust 1. The liquid phase was pressurized and cooled, and then supplied to the first distillation tower. The distillate in which C3H6 was dominant was collected in the tower and was depressurized. C3H6 then was obtained as a byproduct after CO2 was removed by the absorption tower. The distillate in which n-C4H8 was dominant was collected in the second distillation tower. It was cooled and then depressurized after unreacted MeOH was removed using an absorption tower. The hydrocarbon mixture containing n-C4H8 was supplied to the Iso/Dehyd unit. The bottom liquid collected in the second distillation tower was depressurized and mixed with air, and then combusted and cooled. Then, it was released to the air as exhaust 2. 2.2.4. Direct Synthesis of LOs from Syngas in Syn-LOs. The unit for the direct synthesis of LOs from syngas was designed based on a report by Schulte et al.46 (Figure S2). A fixed-bed reactor that uses Fe-loaded carbon nanotubes as the catalyst

was employed as the reactor in the present study. The syngas was pressurized and heated to the desired pressure and temperature, using three compressors and three HXs, and then supplied to the reactor. The product mixture leaving the reactor was cooled and separated into the gas and liquid phases. The gas phase was depressurized and mixed with air, and then combusted and cooled. It then was released to the air as exhaust 1. The liquid phase was depressurized, and then cooled and separated into the gas and liquid phases. The gas phase was depressurized and mixed with air, and then combusted and cooled. It then was released to the air as exhaust 2. The liquid phase was supplied to the first distillation tower. The distillate in which C2H4 was dominant was collected and depressurized. C2H4 then was obtained as a byproduct after CO2 was removed. The distillate in which C3H6 was dominant was collected in the second distillation tower and was depressurized. C3H6 then was obtained as a byproduct after CO2 was removed. The distillate in which n-C4H8 was dominant was depressurized, and then supplied to the Iso/Dehyd unit. The bottom liquid collected in the third distillation tower was depressurized and mixed with air, and then combusted and cooled. It then was released to the air as exhaust 3. 2.2.5. Iso/Dehyd. For all processes, the n-C4H8 Iso/Dehyd unit was designed based on a Japanese patent47 (recall Figure 3, as well as Figures S1 and S2). This process is used to supply 1,3-BD commercially from n-butene obtained through the steam cracking of naphtha. A fixed-bed reactor that uses Mo− Bi−Co oxide as the catalyst was employed as the reactor in the present study. The hydrocarbon mixture containing n-C4H8 was mixed with N2, air, and steam, and then heated to the desired temperature and supplied to the reactor. The product mixture leaving the reactor was cooled and pressurized and then cooled again and supplied to the first distillation tower. For SC-Syn-DME, two distillation towers were employed (see Figure 3). The distillate that was inert in the first tower was released to the air after being depressurized as exhaust 1. The distillate in which 1,3-BD was dominant in the second tower was depressurized and then collected. The bottom liquid collected in the second tower was mixed with air and then combusted and cooled. It then was released to the air as exhaust 2. For TH-Syn-DME, a distillation tower was employed. The distillate was depressurized and mixed with air, and then combusted and cooled. It then was released to the air as exhaust 1. The bottom liquid in which 1,3-BD was dominant was collected after being depressurized. For both cases of Syn-MeOH, the liquid phase that was separated from the hydrocarbons supplied to the distillation tower was a mixture of N2 and O2 (see Figure S1). Therefore, it was released to the air as exhaust 1 after being depressurized. For SC-Syn-MeOH, the distillate that was inert in the distillation tower was depressurized and then released to the air as exhaust 2. In contrast, for TH-Syn-MeOH, the distillate was depressurized and mixed with air, and then combusted and cooled. It then was released to the air as exhaust 2. For both cases, the bottom liquid in which 1,3-BD was dominant was depressurized and collected. For both cases of Syn-LOs, the same equipment was employed in the Iso/Dehyd unit (see Figure S2). The distillate in the first tower was depressurized and mixed with air, and then combusted and cooled. It then was released to the air as exhaust 1. The distillate in which 1,3-BD was dominant in the second tower was collected after being depressurized. The bottom liquid collected in the second tower was depressurized D

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Energy & Fuels Table 1. Elemental Analysis and Proximate Analysis of Lignin Proximate Analysis (wt %, wet basis)

Elemental Analysis (wt %, dry basis)

a

C

H

O

Na

K

moisture

organic matter

ash

higher heating value, HHVa (MJ/kg-daf)

enthalpy of formation, ΔH°f (MJ/kmol)

71.6

8.35

19.3

0.371

0.415

50.0

49.5

0.50

33.1

−442.3

Dry ash free basis.

⎛ o⎞ LHV (MJ/kg) = 32.8c + 120⎜h − ⎟ + 9.3s − 2.5w ⎝ 8⎠

and mixed with air, and then combusted and cooled. It then was released to the air as exhaust 2.

(1)

3. ASSUMPTIONS AND METHODOLOGY The proposed processes were simulated in order to calculate not only the yields of intermediates and 1,3-BD, but also the electric power and heat properties. The initial conditions, with regard to lignin, N2, O2, air, and H2O, were 25 °C and 0 MPaG. A steady-state process simulator (PRO/II, Invensys Systems Japan, Inc.) was used to estimate not only the electric power, such as the compression power, but also the heat properties, such as the energy for heating and recovered heat. The compression efficiency for a compressor was 85%. The temperature for the combustion of the gas phase, using a combustor, was 800 °C. All exhaust was cooled to 25 °C before being released to the air. All distillation towers had 50 trays. The feedstock was supplied to the 25th tray. The reflux ratio was 1, and the pressure in all towers was 2 MPaG. The equilibrium compositions were calculated using an equilibrium software (HSC Chemistry for Windows, Outokumpu Research Oy). Pinch analysis was performed to estimate QH,min, using a spreadsheet (Microsoft Excel, Microsoft Corporation) on the heat loads calculated in sections 3.1−3.4. To simplify the calculation, heat loads of 97%, because of the difficulty in separating Ar.49 In the present study, an O2/N2 (97:3) mixture gas was employed as a gasifying agent, and the electric power for air separation was calculated using an energy consumption of 0.29 kWeh/Nm3-100% O2. The removal of CO2 from syngas required electricity for the regeneration of MDEA as an absorption liquid. The CO2 recovery efficiency was 90%, and the electric power for the regeneration was calculated using 1.93 MJ/kg-CO2.43 The mass balance in the gasification unit was estimated based on those conditions (see Table S1 in the Supporting Information). 3.2. LO Synthesis from Syngas in Syn-DME. The minimum electric power required to increase the syngas pressure from 0 to 3.2 MPaG using three compressors in the DME synthesis step was calculated. The compressed syngas was cooled to 65 °C, using two HXs between the three compressors. The mass balance in the DME synthesis reaction for SC-SynDME was estimated under the condition that the reaction temperature was 280 °C, the reaction pressure was 3.2 MPaG, the CO conversion was 36.1%, and the DME yield was 36.1% on a carbon basis44 (Table S2 in the Supporting Information). E

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The Kexp/Ktheor ratio in the present study was equal to the molar ratio of the experimental amount of n-C4H8 product to the equilibrium amount, because the amounts of DME as feedstock were the same for both cases. Therefore, the rate of progression (RP) was defined as the molar ratio of MC4,exp to the equilibrium amount (denoted as MC4,theor) as follows:

The mass balance in the DME synthesis reaction for TH-SynDME was estimated based on the same reaction pressure and temperature (Table S3 in the Supporting Information). For both cases, the product mixture leaving the reactor was cooled to −50 °C. The mass balance in the LO synthesis reaction from DME for SC-Syn-DME was estimated under the condition that a DME/N2 (4.8:95.2) mixture gas was employed as a feed gas, the reaction temperature was 400 °C, and the reaction pressure was 0 MPaG30 (Table S2). In previous studies on LO synthesis from DME,26−31 the target product was C3H6. Therefore, in the LOs synthesis reaction from DME for TH-Syn-DME, a reasonable reaction temperature that favors n-C4H8 production must be chosen. The equilibrium amount was calculated under the condition that the composition of the DME/N2 feed gas was 4.8:95.2, the reaction pressure was 0 MPaG, and the product gas contained N2(g), H2O(g), CO2(g), C2H4(g), C3H6(g), 1-C4H8(g), cis-2C4H8(g), trans-2-C4H8, n-C5H10(g), MeOH(g), and DME(g). Figure 4 shows the effect of temperature on the equilibrium

RP =

MC4,exp MC4,theor

(2)

Figure 5 shows the relationship between the natural logarithm of RP and the reciprocal of reaction temperature (1/T). The

Figure 5. Relationship between ln(RP) and 1/T in the n-C4H8 production from DME.

correlation coefficient (R) of −0.9295 indicated that there was a strong negative correlation between ln(RP) and 1/T. The relationship can predict the RP value in the synthesis of n-C4H8 from DME. It was similar to that in the Arrhenius plot, indicating that the RP value was related to a reasonable n-C4H8 production rate. Figure 6 shows the effect of temperature on

Figure 4. Effect of temperature on equilibrium amounts of C4H8 and C3H6 in the LO synthesis reaction from DME. Superscript a) indicates the average amount after 125 min (from ref 30).

amounts of C3H6 and n-C4H8. We believe that the reaction temperature in the previous study by Park et al. was 400 °C, because the highest equilibrium amount of C3H6 was obtained at 400 °C. In contrast, the equilibrium amount of n-C4H8 increased with decreasing temperature, and exceeded the equilibrium amount of C3H6 at 20%, the MeOH conversion and selectivity for hydrocarbons with a carbon number of 4 decreased, while the selectivity for heavy compounds such as hydrocarbons with a carbon number of >5 increased.52 The generation of such heavy compounds would lead to catalyst deactivation, because of coking on the catalyst surface. Accordingly, in the DME-to-LOs step, the feed gas with a partial pressure of DME as a feed gas of >8.7% and a reaction temperature of ∼280 °C were desirable conditions, from the point of view of decreasing the electric power of the entire process. When the catalyst for n-C4H8 synthesis from DME is investigated in the future, we must keep in mind that there is an appropriate DME content based on the entire process and high DME content in the feed gas may lead to drastic catalyst deactivation. In previous studies on the LO synthesis reaction from DME, the target product was C3H6.26−30 The design of the process for C3H6 production from coal and natural gas was motivated by the possible depletion of crude oil. The process simulation in the present study revealed the importance of the enhancement of the catalyst for the LO synthesis from DME and appropriate reaction conditions, taking the use of lignin as a feedstock and the realization of the entire process into account. In the future, the authors will report the catalyst used for selective n-C4H8 synthesis in the LO synthesis reaction from DME.

Figure 9. Electric power for each process.

electric power for all processes. Gas compression, MDEA regeneration, and air separation contributed to the electric power. For all cases, the regeneration consumed the most electric power, followed by gas compression and then air separation. Thus, in the present study, we attempted to determine which electric power can be reduced. CO2 removal led to a decrease in the volume of gas to be compressed, which corresponded to a decrease in the electric power for gas compression. Because of this being such a close relationship, it would be very difficult to decrease the electric power for the MDEA regeneration and gas compression without affecting the 1,3-BD yield. In Syn-DME, air was separated into an O2/N2 mixture gas (97:3) and N2 using a cryogenic air separation apparatus (recall Figure 3). The former was supplied to the gasification unit as a gasifying agent, and the latter was supplied to the DME-to-LOs and Iso/Dehyd units as an inert gas. Table 4 shows the mass balance for air, the mixture gas, and N2 for TH-Syn-DME. When the mixture gas required for the gasification unit (154 kmol/h) was produced, theoretically, 445 kmol/h of N2 was produced. However, an additional 303 kmol/h would be required, because a total of 748 kmol/h of N2 was required for the DME-to-LOs and Iso/Dehyd units. The feed gas supplied to the Iso/Dehyd unit was a mixture of hydrocarbons, steam, O2, and N2. It was quite difficult to decrease the amount of N2 in order to keep the feed gas from exploding. In contrast, when a sufficient amount of N2 was produced for the DME-to-LOs and Iso/Dehyd units, 105 kmol/h of the mixture gas remained. Therefore, if the catalytic performance is enhanced, it will be possible to decrease the electric power while maintaining the 1,3-BD yield, even if the amount of N2 supplied to the DME-toLOs unit is decreased. Therefore, in order to decrease the electric power of the entire process, the most desirable conditions in DME-to-LOs were studied. The process in which the obtained N2 was supplied to the DME-to-LOs and Iso/Dehyd units, when the

Table 4. Mass Balances for Cryogenic Air Separation Apparatuses for TH-Syn-DME and Imp TH-Syn-DME Required Flow Rate (kmol/h) case TH-Syn-DME Imp TH-Syn-DME

utility

gasification

DME-to-LOs

Iso/Dehyd

flow rate produced by cryogenic air separation (kmol/h)

residue (kmol/h)

N2 O2/N2 = 97:3 N2 O2/N2 = 97:3

0 154 0 154

693 0 370 0

55 0 55 0

748 258 445 154

0 105 20 0

J

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Energy & Fuels Notes

5. CONCLUSION Three proposed processes for 1,3-BD production from lignin via syngas were simulated to determine which is the most promising, in terms of their respective realizations. (1) The TH-Syn-DME exhibited the maximum 1,3-BD yield (13.4 wt %) and was the most promising when the income from selling products and the expenses for electricity and external heat were taken into account. (2) The electric power was largely dependent on the process (4777−6073 kWe), while the QH,min value was the same (97 kW) for all processes. The high electric power for TH-SynDME (6008 kWe) was attributed to the production of excess N2 through a cryogenic air separation method. A decrease in the amount of N2 supplied to the DME-to-LOs unit led to a decrease in the electric power to 5381 kWe, and the 1,3-BD yield increased to 14.2 wt % because the equilibrium amount of n-C4H8 increased in the LO synthesis reaction from DME. (3) The improvement of the DME-to-LOs step was the top priority for the realization of the proposed process. In the step, the feed gas with >8.7% DME and a reaction temperature of ∼280 °C were favorite conditions for improving the most promising process further. (4) For Syn-MeOH, the volume of syngas available for MeOH synthesis was decreased to adjust the H2/CO/CO2 ratio, while the catalytic performance was very high, leading to poorer realization. (5) In the direct synthesis of LOs from syngas in Syn-LOs, the n-C4H8 yield was lower, because the product selectivity followed the ASF distribution, while the electric power was lower, leading to poorer realization.

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The authors declare no competing financial interest.

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ACKNOWLEDGMENTS This work was supported by the fundamental research fund of AIST. The authors are grateful to valuable comments for Dr. Tadahiro Fujitani, Dr. Tatsuo Yagishita, Dr. Masaru Aoyagi, Dr. Hirohmi Watanabe, and Dr. Shotaro Ito.

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ASSOCIATED CONTENT

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S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.energyfuels.7b02237. Mass balances for gasification (Table S1), SC-Syn-DME (Table S2), TH-Syn-DME (Table S3), SC-Syn-MeOH (Table S4), mass balance for TH-Syn-MeOH (Table S5), SC-Syn-LOs (Table S6), and Imp-Syn-LOs (Table S7); electric power and heat properties for SC-Syn-DME (Table S8), TH-Syn-DME (Table S9), SC-Syn-MeOH (Table S10), TH-Syn-MeOH (Table S11), SC-Syn-LOs (Table S12), Imp-Syn-LOs (Table S13); mass balance for Imp TH-Syn-DME (Table S14); electric power and heat properties for Imp TH-Syn-DME (Table S15); schematic flow diagrams for Syn-MeOH (Figure S1) and Syn-LOs (Figure S2); effect of temperature on the equilibrium amount and experimental n-C4H8 amount in the LO synthesis from MeOH (Figure S3); composite curves for SC-Syn-DME (Figure S4), SC-Syn-MeOH (Figure S5), TH-Syn-MeOH (Figure S6), SC-Syn-LOs (Figure S7), Imp-Syn-LOs (Figure S8), and Imp THSyn-DME (Figure S9) (PDF)

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NOMENCLATURE 1,3-BD = 1,3-butadiene ABS = acrylonitrile butadiene styrene ASF = Anderson−Schulz−Flory DME = dimethyl ether HHV = higher heating value HX = heat exchanger Iso/Dehyd = isomerization/dehydrogenation LHV = lower heating value LOs = light olefins MeOH = methanol O2/C = molar ratio of oxygen as a gasifying agent to carbon in lignin PB = polybutadiene QH,min = minimum external heat R = correlation coefficient RP = rate of progression SBR = styrene−butadiene rubber Ts = supply temperature (°C) Tt = target temperature (°C) α = chain growth probability ΔHf° = enthalpy of formation ΔTmin = minimum internal temperature REFERENCES

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AUTHOR INFORMATION

Corresponding Author

*Tel. & Fax: +81-82-420-8309. E-mail: [email protected]. ORCID

Shinji Fujimoto: 0000-0001-7406-622X K

DOI: 10.1021/acs.energyfuels.7b02237 Energy Fuels XXXX, XXX, XXX−XXX

Article

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DOI: 10.1021/acs.energyfuels.7b02237 Energy Fuels XXXX, XXX, XXX−XXX