Estimation and Improvement of the 1,3-Butadiene Production Process

Jun 3, 2016 - When the moisture content of lignin as a feedstock was 80 wt % and the capacity of the process was 500 t/day on a wet basis, the effecti...
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Estimation and Improvement of the 1,3-Butadiene Production Process from Lignin through Pinch Analysis Toshiaki Hanaoka and Shinji Fujimoto* Bio-based Materials Chemistry Group, Research Institute for Sustainable Chemistry, National Institute of Advanced Industrial Science and Technology (AIST), 3-11-32 Kagamiyama, Higashihiroshima, Hiroshima 739-0046, Japan S Supporting Information *

ABSTRACT: In this paper, a process for 1,3-butadiene (1,3-BD) production from lignin via phenolic compounds, cyclohexane, and n-butene is proposed. The process comprised four unit operations: depolymerization, hydrodeoxygenation, catalytic cracking/dehydrogenation, and dehydrogenation/isomerization. The process was simulated with regard to two cases: (1) one where it was simply combined with previously reported operations (SC case) and (2) one where it was combined with theoretical unit operations in which the ideal chemical reaction occurred and the equilibrium composition was obtained (TH case). For the SC and TH cases, the 1,3-BD yields on a dry and ash-free basis were 0.1 and 11.4 wt %, respectively. When the moisture content of lignin as a feedstock was 80 wt % and the capacity of the process was 500 t/day on a wet basis, the effective energy utilization was determined through pinch analysis to evaluate the economics of the proposed process. For the SC and TH cases, the minimum external required heats were 230 and 248 MW, respectively, indicating that the required input energy hardly decreased even if the chemical reaction in each step proceeded ideally under the reported reaction conditions. In the depolymerization step, the condenser temperature in the distillation column and amount of MeOH used as a solvent were dominant factors for determining the external required heat. Economic evaluation through pinch analysis suggested that increasing the condenser temperature from 65 to 120 °C using a heat pump and drastically decreasing (by >97%) the amount of MeOH used led to an effective decrease in the required input energy to 22 MW.

1. INTRODUCTION 1,3-Butadiene (1,3-BD) is an important backbone chemical that is employed as an intermediate for synthetic rubbers, synthetic resin polymers, and other chemicals. Polybutadiene rubber (BR) and styrene−butadiene rubber (SBR) are mainly used as raw materials for automobile tires, and acrylonitrile−butadiene−styrene (ABS) is used for various commodities from household electric appliances to miscellaneous goods.1 The demand is expected to increase in the future. 1,3-BD is produced directly as a byproduct when ethylene is produced through steam cracking of naphtha or produced through dehydrogenation/isomerization (Dehyd/Iso) of nbutene, such as 1-butene, cis-2-butene, and trans-2-butene, produced through steam cracking of naphtha.2 Therefore, the produced amount of 1,3-BD is largely dependent upon that of ethylene. Recently, ethylene has been produced from shale gas, which is less expensive than naphtha. Thus, the produced amount of ethylene derived from naphtha is decreasing because of economic disadvantages, leading to decreased 1,3-BD production. Therefore, it is important to develop new processes for producing 1,3-BD from feedstocks other than naphtha. When the amount of fossil fuels consumed and the resultant emission of CO2 into the atmosphere are taken into account, biorefineries have received attention because biomass is the only renewable resource. In this concept, biomass is separated into cellulose, hemicellulose, and lignin and then various chemicals and fuels are produced via some building blocks through biochemical and thermochemical conversions.3−6 Cellulose and hemicellulose are converted into 1,3-BD via ethanol (EtOH) through saccharification and fermentation.7,8 © XXXX American Chemical Society

However, lignin is not used effectively because it is not biochemically converted into 1,3-BD. Thermochemical conversion has the advantage of generally higher productivity compared to biochemical conversion. Some researchers have reported 1,3-BD production through pyrolysis of lignin9 and biomass.10−13 However, the 1,3-BD yield was as low as 0.8%,9 indicating that converting lignin into 1,3-BD directly is difficult. Therefore, the production of 1,3-BD as a desired product from lignin has not been reported. In the present study, the authors propose a process in which lignin is thermochemically converted into 1,3-BD via phenolic compounds, cyclohexane (CH), and n-butene. Figure 1 shows the flow diagram of the proposed process. The process involves depolymerization (DP) of lignin, hydrodeoxygenation (HDO) of phenolic compounds, catalytic cracking/dehydrogenation (CC/Dehyd) of CH, and Dehyd/Iso of n-butene. Although the 1,3-BD yield can be calculated because each unit operation, that is, DP,14,15 HDO,16−23 CC/Dehyd,24,25 and Dehyd/Iso,26 has been reported, the input energy required for the newly combined process has to be estimated. Pinch analysis can be used to design the heat recovery network for minimizing the input heat energy. Thus far, the optimal input heat energy has been estimated with regard to some conversion technologies using biomass as a feedstock.27−30 In the present study, pinch analysis was employed Special Issue: In Honor of Michael J. Antal Received: April 8, 2016 Revised: June 2, 2016

A

DOI: 10.1021/acs.energyfuels.6b00822 Energy Fuels XXXX, XXX, XXX−XXX

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Figure 1. Proposed process of 1,3-BD production from lignin.

Figure 2. Schematic flow diagram for the SC case. separation. As a feedstock, lignin mixed with an aqueous solution of Na2SO3 in papermaking was considered. Table 1 shows the analytical results of the feedstock. The moisture content was 80 wt %. The capacity of the process was 500 t/day on a wet basis.

to simulate the proposed process and then the dominant operation affecting total input energy was determined. The effect of the operation conditions on the total input energy was studied to attain the effective energy utilization. Moreover, the economics of the proposed process were evaluated on the basis of results obtained through pinch analysis. On the basis of the findings, the authors suggested priority factors for the development of the process for producing 1,3-BD from lignin.

Table 1. Analytical Results of Lignin as a Feedstock15 elemental analysis (wt %, on a dry basis)

2. PROCESS DESIGN AND ESTIMATION The proposed process was simulated with regard to two cases. In the first case, four unit operations in which high yields of phenolic compounds, CH, n-butene, and 1,3-BD have been reported previously were simply combined. This case is denoted as the SC case. In the other case, four unit operations in which the chemical reaction proceeded ideally and the equilibrium compositions were obtained were combined. The theoretical case is denoted as the TH case. In all unit operations, each chemical reaction was promoted using catalysts.14−26 Therefore, a comparison of the SC and TH cases would reveal the difference between the reported catalytic performance and ideal catalytic performance. The heat energy was reasonably evaluated through pinch analysis, and then the economics of the proposed process were discussed taking into account the required input energy. 2.1. Process Flow. Figure 2 shows the schematic flow diagram of the SC case. That of the TH case (Figure S1 of the Supporting Information) is different from that of the SC case in terms of product

proximate analysis (wt %, on a wet basis)

C

H

O

N

S

asha

moisture

organic matterb

asha

47.7

4.9

24.4

0.1

3.9

19.0

80.0

16.2

3.8

a

With regard to Na2SO3. bMixture of carbon, hydrogen, nitrogen, oxygen, and sulfur.

2.1.1. DP of Lignin. The DP step was designed on the basis of the previous report by Shu et al.15 The lignin feedstock was dried at 105 °C using a heat exchanger (HX) and then mixed with methanol (MeOH) as a solvent to form a slurry in a tank. It was compressed to 4 MPaG using a pump. As a reactant gas, H2 was compressed to 4 MPaG using a compressor. The mixture of slurry and H2 was heated to 280 °C using a HX and then supplied to a reactor. The reaction pressure, temperature, and reaction time were 4 MPaG, 280 °C, and 5 h, respectively. The product mixture leaving the reactor was cooled and depressurized to atmospheric pressure using an expander and then supplied to a tank. The product mixture spontaneously separated into gas, liquid, and solid phases in the tank. The gas phase contained NH3, B

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Energy & Fuels H2S, and unreacted H2. The liquid phase contained phenolic compounds, byproduct, and MeOH. The solid phase was ash. Unreacted H2 was separated from NH3 and H2S using some adsorbents and then recycled (not shown in Figure 2). The liquid mixture in the tank was supplied to a distillation column. MeOH was collected as a distillate and recycled as a solvent. Phenolic compounds as bottom liquid were used in the following step after cooling using a HX. The configuration in the DP step of the TH case was the same as that of the SC case (Figure S1 of the Supporting Information). 2.1.2. HDO of Phenolic Compounds. The HDO step was designed on the basis of the previous report by Zhang et al.17 Phenolic compounds as a feedstock were compressed to 5 MPaG using a pump. As a reactant, H2 was compressed to 5 MPaG using a compressor and then mixed with the compressed feedstock. The mixture was heated to 300 °C using a HX and then supplied to a reactor. The reaction pressure, temperature, and reaction time were 5 MPaG, 300 °C, and 8 h, respectively. The product mixture leaving the reactor was cooled using a HX and then supplied to a tank after depressurizing to atmospheric pressure using an expander. The product mixture in the tank comprised unreacted H2 in the gas phase and the liquid mixture containing CH. Unreacted H2 was recycled (not shown in Figure 2). For the SC case, the liquid product in the tank was supplied to a distillation column (Figure 2). The CH collected as a distillate was used in the following step. The byproduct as bottom liquid was combusted at 800 °C after mixing with air. For the TH case, the liquid product in the tank was a mixture of CH and CO2. Therefore, it was depressurized to atmospheric pressure using an expander and then supplied to a tank (Figure S1 of the Supporting Information). The CH in the tank was spontaneously separated from CO2 and then used in the following step. 2.1.3. CC/Dehyd of CH. The CC/Dehyd step was designed on the basis of the previous report by Slagtern et al.24 The CH as a feedstock was heated to 600 °C using a HX and then mixed with N2 heated at 600 °C in a tank. The mixture gas was supplied to a reactor. The reaction pressure and temperature were 0 MPaG and 600 °C. The product mixture leaving the reactor was cooled to −100 °C and compressed to 2 MPaG using a compressor and then cooled again using a HX. For the SC case, the liquid mixture cooled using the HX was supplied to the first distillation column (Figure 2). The distillate contained ethane, ethylene, propane, and propylene, was depressurized to atmospheric pressure using an expander, and then combusted at 800 °C after mixing with air. The bottom liquid was a mixture of organic compounds with a boiling point higher than that of n-butene. They were supplied to the second distillation column. The distillate contained mainly n-butene and was used in the following step after depressurizing to atmospheric pressure using an expander. The bottom liquid was combusted at 800 °C after mixing with air. For the TH case, the liquid mixture cooled using the HX was supplied to the first distillation column (Figure S1 of the Supporting Information). The distillate contained mainly N2 and ethylene. It was depressurized to atmospheric pressure using an expander and combusted at 800 °C after mixing with air. The bottom liquid was a mixture of ethylene and n-butene and was supplied to the second distillation column. The distillate was mainly ethylene; it combusted at 800 °C after depressurizing to atmospheric pressure and mixing with air. The bottom liquid was a mixture of n-butene and unreacted CH and was used in the following step after depressurizing to atmospheric pressure using an expander. 2.1.4. Dehyd/Iso of n-Butene. The Dehyd/Iso step was designed on the basis of a previous patent.26 n-Butene as a feedstock and an additional mixture of gases, such as N2, O2, and steam, were heated to 350 °C in HXs. The mixture of gases was supplied to a reactor. The reaction pressure and temperature were 0 MPaG and 350 °C, respectively. The product mixture leaving the reactor was cooled and compressed to 2 MPaG and then cooled again in a HX. For the SC case, the liquid mixture cooled using the HX was supplied to the first distillation column (Figure 2). The distillate was depressurized using an expander and then combusted at 800 °C after mixing with air. The bottom liquid was supplied to the second

distillation column. The distillate contained 1,3-BD as a main component, was depressurized to atmospheric pressure, and heated. Accordingly, the mixture contained 1,3-BD and n-butane, and nbutene as a final product was collected at 25 °C and 0 MPaG. The bottom liquid was combusted at 800 °C after depressurizing to atmospheric pressure and mixing with air. For the TH case, the liquid mixture cooled using the HX was supplied to a distillation column (Figure S1 of the Supporting Information). The distillate contained N2, unreacted O2, and 1,3-BD, was depressurized to atmospheric pressure using an expander, and then combusted at 800 °C after mixing with air. The bottom liquid was depressurized and heated, and 1,3-BD was collected as a final product at 25 °C and 0 MPaG.

3. ASSUMPTION AND METHODOLOGY The proposed process was simulated to calculate 1,3-BD yield, mechanical energy, and heat loads. The initial conditions with regard to lignin, H2, N2, O2, air, and H2O were 25 °C and 0 MPaG. A steadystate process simulator (PRO/II, Invensys Systems Japan, Inc.) was used to estimate not only heat loads, such as heating and cooling, but also mechanical energy, such as compression power. The equilibrium compositions were calculated using equilibrium software (HSC Chemistry for Windows, Outokumpu Research Oy). Pinch analysis was performed using a spreadsheet (Microsoft Excel, Microsoft Corporation) on the basis of heat loads calculated in sections 3.1 and 3.2. To simplify the calculation, heat loads less than 0.25% on a maximum heat basis were considered negligible. 3.1. SC Case. 3.1.1. DP of Lignin. The process simulator did not have any thermodynamic data with regard to lignin. Therefore, the compression power with regard to the slurry of lignin and MeOH using a pump was calculated using the Bernoulli equation:

w = (P2 − P1)/ρη

(1)

where w (kg/h), P1 (PaG), P2 (PaG), ρ (kg/m ), and η (%) are the slurry feed rate, pressure before compression (0 MPaG), pressure after compression (4 MPaG), feedstock density (788 kg/m3), and pump efficiency (50%), respectively. As a reactant gas, H2 was compressed using three compressors (Figure 2). The efficiency of the compressors was assumed to be 85%. The compression induced heating of H2. Two HXs were set between the compressors to cool compressed H2 to 65 °C to estimate the recovered heat. The outlet pressures of the first and second compressors were controlled to minimize the electricity consumption of the three compressors. The enthalpy change in the DP of lignin at 280 °C and 4 MPaG was calculated as follows. It was assumed that lignin as a feedstock had the molecular structure proposed by Nimz,31 and the mass of the molecular structure was equal to that of 1 kmol of lignin. When 1 kmol of lignin was depolymerized to produce guaiacol, the amounts of H2 consumed to cleave C−O and C−C bonds were estimated to be 34.0 and 30.6 kmol, respectively. Shu et al. reported a 13% yield of phenolic compounds in the DP of lignin.15 It was assumed that only guaiacol was produced. Therefore, 4.4 and 4.0 kmol of H2 were consumed to cleave C−O and C−C bonds, respectively. The enthalpy changes in the cleavage of C−O and C−C bonds by H2 were considered to be those in the following reactions and calculated using the equilibrium software: 3

CH3OH(g) + H 2(g) = CH4(g) + H 2O(g)

C2H6(g) + H 2(g) = 2CH4(g)

−116.9 kJ/mol (2)

−67.8 kJ/mol

(3) 15

Taking into account that the reaction time was 5 h, the enthalpy change in DP was 98 MJ/h (27 kW). NH3 and H2S as product gases were generated because the lignin as a feedstock contained nitrogen and sulfur (Table 1). It was not clear how these atoms were bound to other atoms, and the nitrogen and sulfur contents were very low. Therefore, the enthalpy changes in NH3 and H2S production were negligible. C

DOI: 10.1021/acs.energyfuels.6b00822 Energy Fuels XXXX, XXX, XXX−XXX

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was cooled to 25 °C using a HX, the released and recovered heat were calculated. 3.1.4. Dehyd/Iso of n-Butene. The enthalpy change in the Dehyd/ Iso of n-butene at 350 °C and 0 MPaG was calculated using the equilibrium software as follows:

In the DP step, the product mixture leaving the reactor contained gas, liquid, and solid phases in the tank. The gas phase contained NH3, H2S, and unreacted H2. The liquid components were phenol, guaiacol, p-tert-amylphenol, MeOH, and a heavy phase based on the previous report by Shu et al.15 The mixture of phenol, guaiacol, p-tertamylphenol, and MeOH spontaneously separated from the heavy phase in the tank. Then, the mixture was supplied to the fifth tray of a distillation column to separate the distillate and bottom liquid. The column had 10 trays. The reflux ratio was 1. When 100% MeOH as the distillate and the mixture of phenol, guaiacol, and p-tert-amylphenol as the bottom liquid were obtained, the heat collected in a condenser and heat required in a reboiler were calculated using the process simulator. The heavy phase leaving the tank was mixed with the theoretical amount of air for complete combustion and then was combusted at 800 °C. The exhaust was cooled to 25 °C using a HX. The released and recovered heat were calculated assuming that the heavy phase was naphthalene. 3.1.2. HDO of Phenolic Compounds. As a matter of convenience, it was assumed that phenol had the same reactivity as guaiacol. The mixture of guaiacol and p-tert-amylphenol was regarded as a feedstock in the HDO step. The compression power of the mixture up to 5 MPaG was calculated using eq 1. The feedstock density was 1001 kg/ m3. As a reactant gas, H2 was compressed using three compressors (Figure 2). The compression power was calculated using the same method mentioned in section 3.1.1. The enthalpy change in HDO of guaiacol at 300 °C and 5 MPaG was calculated using the equilibrium software as follows:

1‐C4 H8(g) + 0.5O2 (g) = 1,3‐C4 H6(g) + H 2O(g) − 196.4 kJ/mol

Taking the 68.4% 1,3-BD yield into account, the enthalpy change in the Dehyd/Iso was 16 MJ/h (5 kW). The product yield was calculated on the basis of a previous report.26 The product mixture compressed to 2 MPaG and cooled to −80 °C was supplied to the 25th tray of the first distillation column to separate the distillate and bottom liquid. The column had 50 trays. The reflux ratio was 1. When the distillate mainly contained organic compounds with a boiling point lower than that of 1,3-BD and the bottom liquid contained those with a boiling point higher than that of 1,3-BD, the heat collected in a condenser and heat required in a reboiler were calculated. When the distillate mixed theoretically with air was combusted at 800 °C and the exhaust was cooled to 25 °C using a HX, the released and recovered heat were calculated. The bottom liquid was supplied to the 40th tray of the second distillation column. The column had 50 trays. The reflux ratio was 1. When 1,3-BD was mainly collected as the distillate and the mixture of CH and benzene was mainly collected as the bottom liquid, the heat collected in a condenser and heat required in a reboiler were calculated. When the distillate as a final product was collected at 25 °C and 0 MPaG, the required heat was calculated. When the bottom liquid mixed theoretically with air was combusted at 800 °C and the exhaust was cooled to 25 °C using a HX, the released and recovered heat were calculated. 3.2. TH Case. The assumption and method for simulating the TH case were the same as those for the SC case, except for the descriptions below. 3.2.1. DP of Lignin. The enthalpy change in the DP of lignin at 280 °C and 4 MPaG was calculated as follows. It was assumed that the lignin as a feedstock had the molecular structure proposed by Nimz.31 When 1 kmol of lignin was depolymerized to produce guaiacol in 100% yield, the amounts of H2 consumed to cleave C−O and C−C bonds were estimated to be 34.0 and 30.6 kmol, respectively. It was assumed that only guaiacol was produced. Taking into account that the reaction time was 5 h,15 the enthalpy change in the DP was 751 MJ/h (209 kW). The enthalpy change in NH3 and H2S production was negligible. In the DP step of the TH case, the product mixture leaving the reactor contained gas, liquid, and solid phases in the tank. The gas phase consisted of NH3, H2S, and unreacted H2. The liquid components were guaiacol and EtOH. The liquid mixture was supplied to the fifth tray of a distillation column to separate the distillate and bottom liquid. The column had 10 trays. The reflux ratio was 1. When the distillate was a mixture of MeOH and EtOH and the bottom liquid contained 99% guaiacol, the heat collected in a condenser and heat required in a reboiler were calculated using the process simulator. 3.2.2. HDO of Guaiacol. The compression power of the feedstock up to 5 MPaG was calculated using eq 1. The density of guaiacol as a feedstock was 1110 kg/m3. As a reactant gas, H2 was compressed using three compressors (Figure S1 of the Supporting Information). The compression power was calculated using the same method mentioned in section 3.1.1. Assuming that eq 4 attained the equilibrium condition, the enthalpy change in the HDO of guaiacol at 300 °C and 5 MPaG was calculated. Taking into account that the reaction time was 8 h, the enthalpy change in the DP was 582 MJ/h (162 kW). 3.2.3. CC/Dehyd of CH. The enthalpy change in the CC/Dehyd of CH at 600 °C and 0 MPaG was calculated when eq 5 attained the equilibrium condition after 1 h of reaction. The enthalpy change in the CC/Dehyd was 1215 MJ/h (337 kW). The gas leaving the reactor was a mixture of 1-butene, ethylene, N2, and unreacted CH. The mixture gas compressed to 2 MPaG was

C7H8O2 (g) + 2H 2(g) = C6H12(g) + CO2 (g) − 277.8 kJ/mol

(4)

Taking into account that the reaction time was 8 h and the CH yield was 17.1%,17 the enthalpy change in the DP was 21 MJ/h (6 kW). Zhang et al. reported that CH, alkyl-substituted CH, and alkylsubstituted benzene were obtained as products.17 In the present study, the liquid components were CH, dimethyl CH, and guaiacol. The distillation column had 30 trays. The product mixture supplied to the 15th tray was separated into the distillate and bottom liquid. The reflux ratio was 1. When 100% CH as the distillate and a mixture of dimethyl CH and guaiacol as the bottom liquid were obtained, the heat collected in a condenser and heat required in a reboiler were calculated. When the bottom liquid mixed theoretically with air was combusted at 800 °C and the exhaust was cooled to 25 °C using a HX, the released and recovered heat were calculated. 3.1.3. CC/Dehyd of CH. The enthalpy change in the CC/Dehyd of CH at 600 °C and 0 MPaG was calculated using the equilibrium software as follows:

C6H12(g) = 1‐C4H8(g) + C2H4(g)

+173.0 kJ/mol

(6) 26

(5)

Taking into account that Slagtern et al. reported a 9% yield of n-butene using a continuous flow reactor,24 the enthalpy change in the CC/ Dehyd was 14 MJ/h (4 kW). The product gas was 1-butene, ethylene, ethane, propylene, propane, n-butane, benzene, toluene, o-xylene, N2, and unreacted CH based on the previous report by Slagtern et al.24 The mixture compressed to 2 MPaG was supplied to the 50th tray to separate the distillate and bottom liquid. The first distillation column had 100 trays. The reflux ratio was 1.25. When the distillate contained organic compounds with a boiling point lower than that of CH and the bottom liquid contained those with a boiling point higher than that of CH, the heat collected in a condenser and heat required in a reboiler were calculated. When the distillate mixed theoretically with air was combusted at 800 °C and the exhaust was cooled to 25 °C using a HX, the released and recovered heat were calculated. The bottom liquid was supplied to the 50th tray of the second distillation column. The column had 100 trays. The reflux ratio was 1. When n-butene was mainly collected as the distillate and aromatic compounds were mainly collected as the bottom liquid, the heat collected in a condenser and heat required in a reboiler were calculated. When the bottom liquid mixed theoretically with air was combusted at 800 °C and the exhaust D

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Energy & Fuels Table 2. Productivity and Yield of Intermediate in Each Step and 1,3-BD Yield for the SC and TH Cases lignin feed rate

a

DP

case

kg dry/h

kg dafa/h

productivity (kg/h)

SC TH

4166.7 4166.7

3375.0 3375.0

371.4 2056.8

HDO

CC/Dehyd

Dehyd/Iso

yield (wt %)

productivity (kg/h)

yield (wt %)

productivity (kg/h)

yield (wt %)

productivity (kg/h)

yield (wt %)

1,3-BD yield (wt %)a

11.0 60.9

73.4 1408.1

19.8 68.5

6.6 393.6

9.0 28.0

4.0 386.1

60.5 98.1

0.1 11.4

Dry and ash-free basis.

cooled and then supplied to the 90th tray of the first distillation column. The column had 100 trays. The reflux ratio was 1. When the distillate contained a mixture gas of ethylene and N2 and the bottom liquid mainly contained CH, the heat collected in a condenser and the heat required in a reboiler were calculated. When the distillate mixed theoretically with air was combusted at 800 °C and the exhaust was cooled to 25 °C, the released and recovered heat were calculated. The bottom liquid was supplied to the 90th tray of the second distillation column. The column had 100 trays. The reflux ratio was 1. When the distillate mainly contained ethylene and the bottom liquid contained 97.3% n-butene, the heat collected in a condenser and heat required in a reboiler were calculated. When the distillate mixed theoretically with air was combusted at 800 °C and the exhaust was cooled to 25 °C, the released and recovered heat were calculated. 3.2.4. Dehyd/Iso of n-Butene. The enthalpy change in the Dehyd/ Iso of n-butene at 350 °C and 0 MPaG was calculated when eq 6 attained the equilibrium condition after a reaction time of 1 h. The enthalpy change in the Dehyd/Iso was 1424 MJ/h (396 kW). The mixture gas compressed to 2 MPaG was cooled and then supplied to the 25th tray of the first distillation column. The column had 50 trays. The reflux ratio was 1. When the distillate contained N2, O2, and 1,3-BD and the bottom liquid contained 100% 1,3-BD, the heat collected in a condenser and heat required in a reboiler were calculated. When the bottom liquid was depressurized to atmospheric pressure and heated to 25 °C, the required heat was calculated. When the distillate mixed theoretically with air was combusted at 800 °C and the exhaust was cooled to 25 °C, the released and recovered heat were calculated.

Table 3. Heat Properties for the SC Case

4. RESULTS AND DISCUSSION 4.1. 1,3-BD Yield and Required Energy for the SC and TH Cases. For both cases, simulations were performed. The mass balance results for the SC and TH cases are shown in Tables S1 and S2 of the Supporting Information, respectively. Table 2 shows the 1,3-BD yield on a dry and ash-free basis. For the SC and TH cases, the 1,3-BD yields were 0.1 and 11.4 wt %, respectively. In comparison to the 1,3-BD yield by pyrolysis of lignin (0.8%),9 that for the TH case was much higher, while that for the SC case was lower. This result indicates that improvement of the catalysts employed in the proposed process may lead to higher 1,3-BD yields compared to those by pyrolysis. Table 2 also shows the productivity and yield of intermediates obtained in each step. The difference in the intermediate yields between the SC and TH cases decreased in the order of DP (5.5 times) > HDO (3.5 times) > CC/Hydro (3.1 times) > Dehyd/Iso (1.6 times), taking into account that 1,3-BD is commercially produced through the Dehyd/Iso of nbutene. When the results and situation are taken into account, the catalysts employed in the DP step have to be improved. Tables S3 and S4 of the Supporting Information show the mechanical energy and heat properties for the SC and TH cases, respectively. In each step for both cases, heat loads were much higher than mechanical energy. Therefore, the input heat energy was optimized through pinch analysis. Tables 3 and 4 show the operations and heat properties selected from Tables S3 and S4 of the Supporting Information for pinch analysis.

Some operations, such as air heating, combustion, and exhaust cooling, had extremely low heat load values. However, they were taken into account for pinch analysis because they had the same supply temperature (Ts) and target temperature (Tt) that some operations with large values had. Figures 3 and 4 show the composite curves when the minimum internal temperature approached 10 °C. The pinch points for the SC and TH cases were 65.5 and 65.6 °C, respectively, and were substantially affected by the condenser temperature in the DP step. In the enthalpy from 150 000 to 350 000 kW for both cases, the recovered heat of combustion, cooling of exhaust, product mixture, and H2, and part of the condensing heat in the condenser were used for lignin drying and part of MeOH, heavy phase, air, and H2 was used for heating. However, the minimum external heat (QH min) required for the SC and TH cases was 230 and 248 MW, respectively, because the recovered heat could not supply the heat for the reboiler in the distillation column, etc. Table 5 shows the mechanical energy, QH min, and total required energy for both cases. The mechanical energy for the TH case (6579 kW) was lower than that for the SC case (6776 kW). For the TH case, the gas phase to be compressed had a smaller volume because the yield of liquid intermediates was higher. The total required energies for the SC and TH cases were 237 and 255 MW, respectively. When the chemical reactions proceeded ideally, the total required energy increased by 1.1 times, while the 1,3-BD yield for the TH case was approximately 97 times that for the SC case (Table 2).

step DP

HDO DP, HDO, CC/Dehyd, and Dehyd/Iso

E

operation

type

Ts (°C)

Tt (°C)

heat load (kW)

lignin drying H2 cooling by first HX H2 cooling by second HX MeOH + lignin (including ash) heating H2 heating product mixture cooling reboiling in the distillation column condensation in the distillation column distillate cooling product mixture cooling heavy phase + off gas + air heating combustion exhaust cooling

cold hot

25.0 282.6

105.0 65.0

121013 2378

hot

234.4

65.0

1854

cold

25.0

280.0

120800

cold hot

129.7 280.0

280.0 25.0

1650 123752

cold

208.6

208.8

187812

hot

65.5

65.3

88216

hot hot

65.3 300.0

25.0 25.0

99475 565

cold

25.0

800.0

16935

hot hot

800.2 800.0

800.0 25.0

31928 17158

DOI: 10.1021/acs.energyfuels.6b00822 Energy Fuels XXXX, XXX, XXX−XXX

Article

Energy & Fuels Table 4. Heat Properties for the TH Case step DP

HDO CC/Dehyd

Dehyd/Iso CC/Dehyd and Dehyd/Iso

operation

type

Ts (°C)

Tt (°C)

heat load (kW)

lignin drying H2 cooling by first HX H2 cooling by second HX MeOH + lignin (including ash) heating H2 heating product mixture cooling reboiling in the distillation column condensation in the distillation column distillate cooling product mixture cooling feedstock heating product mixture first cooling to −100 °C first off-gas heating reaction heat product mixture cooling air heating combustion exhaust cooling

cold hot hot cold cold hot cold hot hot hot cold hot cold hot hot cold hot hot

25.0 282.6 234.4 25.0 129.7 280.0 221.0 65.6 65.4 300.0 −59.0 600.0 −31.0 350.2 350.0 25.0 800.2 800.0

105.0 65.0 65.0 280.0 280.0 25.0 221.2 65.4 25.0 25.0 600.0 −100.0 800.0 350.0 −20.0 800.0 800.0 25.0

121013 2378 1854 120800 1650 125042 188545 88513 99819 439 775 923 493 396 654 5777 12462 6623

the DP step have to be improved in the near future. In contrast, improvements to the catalysts in the proposed process hardly led to a decrease in the input energy. This result suggests that reconfiguration of the process is required to enhance the economics of the proposed process. 4.2. Decreasing Input Energy through Pinch Analysis by Introducing a Heat Pump. From Figures 3 and 4, one can see that the condensation heat recovered in the distillation column during the DP step supplied approximately 40% of the heat for lignin drying in the enthalpy range from 150 000 to 200 000 kW. If the condenser temperature is increased, the condensation heat would be expected to be supplied for lignin drying and the QH min values would decrease for both cases. The condenser temperature can be increased practically using a heat pump, while the introduction of the heat pump would increase the mechanical energy. Therefore, the effect of the condenser temperature on total required energy was investigated. The coefficient of performance (COP) is usually defined as follows:

Figure 3. Composite curves as the minimum internal temperature approached 10 °C for the SC case.

COP = Q /W

(7)

where Q (kW) and W (kW) are the output of heat and input of work, respectively. Fujimoto et al. reported the improvement of a bioethanol production process using a heat pump.28 Because the COP was 19.65 when saturated steam was compressed, the required energy was calculated assuming that the efficiency of the heat pump was 70% (i.e., COP of 13.755).28 In the present study, the same COP of the heat pump was used taking into account the possibility that saturated MeOH is compressed to increase the condenser temperature. Figure 5 shows the effect of the condenser temperature in the DP step on the QH min value for both cases. The QH min values decreased monotonously as the condenser temperature increased from 65 to 120 °C and became constant at temperatures greater than 120 °C. Figures 3 and 4 indicated that the straight lines from 25 to 105 °C in the cold composite curve were closest to each pinch point (65.5 and 65.6 °C). The straight lines were closest to the pinch point even when the condenser temperature was increased to 110 °C (not shown in the figures). However, when the condenser temperature was 120 °C, the straight lines

Figure 4. Composite curves as the minimum internal temperature approached 10 °C for the TH case.

The process of 1,3-BD production from lignin proposed in the present study was newly combined with previously reported unit operations. In the DP, HDO, and CC/Dehyd steps, the catalysts employed should be improved to increase the intermediate yields. In particular, taking the ideal phenolic compound yield in Table 2 into account, effective catalysts in F

DOI: 10.1021/acs.energyfuels.6b00822 Energy Fuels XXXX, XXX, XXX−XXX

Article

Energy & Fuels Table 5. Mechanical Energy, Condenser Temperature, QH min, and Total Energy case

mechanical energy (kW)

condenser temperature (°C)

SC Imp-SC TH Imp-TH-1 Imp-TH-2 Imp-TH-3

6776 6776 6579 6579 6579 6579

65.3 120.0 65.4 120.0 120.0 120.0

mechanical energy for heat pump (kW) 6413 6435 6435 6435

QH min (kW)

total energy (kW)

229900 159678 247931 179148 191171 8488

236676 172868 254510 192162 204185 21502

4.3. Economic Evaluation and Improvement through Pinch Analysis. As stated in the previous section, the 1,3-BD yield and required energy were optimized on the basis of reported reaction conditions. In this section, the economics of the TH cases were briefly evaluated, taking into account the income obtained by selling products and the expenses for electricity and heat, followed by being improved by changing reaction conditions. Table 6 shows the income, expenses, and balance of payments on a feedstock weight basis. For the Imp-TH-1 case, the balance of payments was negative because the expenses were overwhelming compared to the income obtained by selling 1,3-BD. This result suggests that other products should be sold in addition to 1,3-BD to increase the income. In the CC/Dehyd step of the TH case, ethylene was produced (Table S2 of the Supporting Information) and was combusted at 800 °C after separating from n-butene using the two distillation columns and mixing air (Figure S1 of the Supporting Information). The distillate obtained in the first distillation column was a mixture of ethylene and nitrogen. Therefore, the case in which the distillate was sold without combustion was regarded as the Imp-TH-2 case. Figure 6 shows the composite curves when the minimum internal temperature approached 10 °C. The pinch point was 231 °C. Table 5 also shows the mechanical energy, QH min, and total required energy for this case. The QH min value increased from 179 148 to 191 171 kW because no heat produced by ethylene combustion was used for reboiling in the distillation columns. Table 6 also shows the economic evaluation of this case. The increase in the income caused by selling ethylene led to a decrease in the negative balance of payment by approximately 7% compared to the Imp-TH-1 case. This result suggests that further decreasing the input energy is required to enhance the economics. Figure 6 also indicates that the heat energy for reboiling in the distillation columns was a dominant factor for determining the QH min value. This was attributed to the heat energy required to vaporize a large amount of MeOH (Table 4).

Figure 5. Effect of the condenser temperature in the DP step on the QH min value.

from 65 to 234 °C in the hot composite curve were closest to the left ends at 209 and 221 °C in Figures 3 and 4, respectively. The relationship was maintained at above 120 °C (not shown in the figures). Therefore, increasing the condenser temperature above 120 °C did not contribute to the decrease in the QH min value. These results indicated that increasing the condenser temperature up to 120 °C led to sufficient heat for lignin drying. Taking into account that the temperature is not usually increased beyond 150 °C using a heat pump, the total required energy could be decreased practically by increasing the condenser temperature up to 120 °C. Table 5 also shows the values of QH min, mechanical energy for the heat pump, and total required energy when the condenser temperature was 120 °C for the improved SC and TH cases (Imp-SC and Imp-TH-1). Although introducing the heat pump increased the mechanical energy by approximately 13 MW, the total required energies for the Imp-SC and ImpTH-1 cases were 173 and 192 MW, respectively. For both cases, the introduction of the heat pump led to a decrease in the total required energy by approximately 25%. These results suggest that the introduction of the heat pump is indispensable for improving the proposed process. Table 6. Economic Evaluation for the TH Case income

expenses

case

product

selling price ($/t)

yield kg/h

yield t/day

$/day

energy

required energy (kW)

time (h)

purchase price ($/kWh)

cost ($/day)

Imp-TH-1

1,3-BD

1500

386.1

9.3

13900

0.1 0.04

386.1 884.8

9.3 21.2

24 24

0.1 0.04

1500 1200

386.1 884.8

9.3 21.2

13014 179148 192162 13014 191171 204185 13014 8488 21502

24 24

1500 1200

electricity heat total electricity heat total electricity heat total

24 24

0.1 0.04

31233 171982 203216 31233 183524 214757 31233 8148 39381

Imp-TH-2

Imp-TH-3

total 1,3-BD C2H4 total 1,3-BD C2H4 total

13900 13900 25482 39382 13900 25482 39382

G

balance of payments ($/t of feedstock)

−379

−351

0

DOI: 10.1021/acs.energyfuels.6b00822 Energy Fuels XXXX, XXX, XXX−XXX

Article

Energy & Fuels

In the present study, the simulation of the process through pinch analysis suggested the following valuable means for developing the proposed process: (1) enhancing catalytic performance in each step and (2) increasing the condenser temperature using a heat pump in the DP step. Further efforts to enhance the economics should be made, and investigations on the catalyst when a small amount of reaction solvent is used in the DP step will be conducted experimentally.

5. CONCLUSION A process for 1,3-BD production from lignin via phenolic compounds, CH, and n-butene was proposed. The process was simulated with regard to SC and TH cases with a capacity of 500 t/day. The effective energy utilization was estimated through pinch analysis to evaluate the economics of the proposed process. For the SC and TH cases, the 1,3-BD yields were 0.1 and 11.4 wt %, respectively. The differences in the yields of each step decreased in the order of DP > HDO > CC/ Dehyd > Dehyd/Iso. The required input energies were 230 and 248 MW, respectively, and the heat energy was dominant for both cases. A comparison of the SC and TH cases indicated that the required energy hardly decreased even if the chemical reaction in each step proceeded ideally under the reported reaction conditions. In the DP step, the condenser temperature in the distillation column and the amount of MeOH used as a solvent were dominant factors for determining the QH min value. The economic evaluation for the TH cases suggested that increasing the condenser temperature using a heat pump from 65 to 120 °C and decreasing the amount of MeOH used by more than 97% compared to that reported previously led to the effective decrease in the QH min value to 22 MW, followed by the enhancement of the economics of the proposed process.

Figure 6. Composite curves as the minimum internal temperature approached 10 °C for the Imp-TH-2 case.

Unfortunately, the relationship between the amount of MeOH and the DP reaction is not clear. The required heat energy could be decreased if the reaction proceeded ideally as the amount of MeOH decreased. Therefore, the effect of the amount of MeOH and the economics of the TH cases were investigated through pinch analysis (not shown in the figures). Figure 7 shows the effect of the amount of MeOH used in the

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ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.energyfuels.6b00822. Mass balance for the SC case (Table S1), mass balance for the TH case (Table S2), mechanical energy and heat properties for the SC case (Table S3), mechanical energy and heat properties for the TH case (Table S4), and schematic flow diagram for the TH case (Figure S1) (PDF)

Figure 7. Effect of the amount of MeOH in the DP step on the balance of payments.

DP step and the balance of payments. The x axis represents the ratio of the amount of MeOH used to the amount of MeOH used in the Imp-TH-2 case. Here, the amount of MeOH used in the Imp-TH-2 case was the same as those in the SC, Imp-SC, TH, and Imp-TH-1 cases. Figure 7 and Table 6 indicate that the balance of payments was 0 when the amount of MeOH used was 3.1% on the basis of the Imp-TH-2 case. This was regarded as the Imp-TH-3 case. Table 5 also shows the mechanical energy, QH min value, and total energy of this case. The QH min value decreased drastically compared to that of the TH case, leading to a decrease in the total energy by approximately 92%. Figure 7 indicates that the balance of payments was positive when the amount of MeOH used was less than 3.1% on the basis of the amount used in the Imp-TH-2 case. Instead of the drastic decrease, the amount of MeOH used in the Imp-TH-3 case was reasonable for the DP reaction. For example, when the ratio was 2%, the volume of feedstock was almost the same as that of MeOH. The feedstock would be mixed with MeOH well, and the reaction may proceed under such reaction conditions. For this case, the balance of payments was positive ($4/t).

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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H

NOMENCLATURE 1,3-BD = 1,3-butadiene CH = cyclohexane BR = polybutadiene rubber SBR = styrene−butadiene rubber ABS = acrylonitrile−butadiene−styrene DP = depolymerization HDO = hydrodeoxygenation CC/Dehyd = catalytic cracking/dehydrogenation Dehyd/Iso = dehydrogenation/isomerization MeOH = methanol DOI: 10.1021/acs.energyfuels.6b00822 Energy Fuels XXXX, XXX, XXX−XXX

Article

Energy & Fuels

(30) Morandin, M.; Toffolo, A.; Lazzaretto, A.; Marechal, F.; Ensinas, A. V.; Nebra, S. A. Energy 2011, 36, 3675−3690. (31) Nimz, H. Angew. Chem., Int. Ed. Engl. 1974, 13, 313−321.

EtOH = ethanol Ts = supply temperature (°C) Tt = target temperature (°C) QH min = minimum external heat COP = coefficient of performance Q = output of heat (kW) W = input of work (kW)

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DOI: 10.1021/acs.energyfuels.6b00822 Energy Fuels XXXX, XXX, XXX−XXX