Article Cite This: Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Assessment on CO2 Utilization through Rigorous Simulation: Converting CO2 to Dimethyl Carbonate Bor-Yih Yu, Meng-Kai Chen, and I-Lung Chien* Department of Chemical Engineering, National Taiwan University, No. 1, Sec. 4, Roosevelt Road, Taipei 10617, Taiwan S Supporting Information *
ABSTRACT: This paper intends to discuss the economical performances and CO2 reduction potential of two CO2-based dimethyl carbonate (DMC) production processes through rigorous process simulation. One of them is the direct production process with addition of butylene oxide (BO) as dehydrating agent (DIRBO porocess), while the other is the indirect production process through ethylene carbonate (EC) as an intermediate (IND-EC process). Both processes are systematically optimized and heatintegrated. From economical evaluation, the IND-EC process exhibits economical attractiveness, while the DIR-BO process does not. We suggest that once the reaction rate of the DIR-BO process can be improved, the overall economic performance of the direct process can be much better. From the aspect of CO2 reduction, the net CO2 emissions throughout both processes are calculated. We found that DIR-EO process is largely carbon positive, with CO2 emission of 2.242 (kg CO2/kg DMC), yet for the IND-EC process, it is near carbon neutral, with CO2 emission of 0.049 (kg CO2/kg DMC). Thus, from the aspect of achieving CO2 reduction, converting it into DMC provides limited benefits.
1. INTRODUCTION To alleviate the great environmental threat caused by CO2 accumulation, the utilization of it as an environmentally benign feed stock to produce valuable chemicals has been regarded as a possible way to make a contribution. In the process systems engineering (PSE) point of view, extensive studies can be done, as suggested in a recent review paper by Roh et al.1 First, CO2 is an important building block in the C1 industry; thus, the potential products can be obtained through a screening step. However, due to the thermodynamically stable structure of CO2, its conversion toward other chemicals is not easy. Hence, detailed studies about process route, process design, and optimization is needed to enhance the production performance. Moreover, the sustainability and feasibility of the designed process should also be investigated to find out whether the current design is economically attractive or not. Last but not the least, processes for CO2 conversion are now at a relative early stage; thus, there are still many uncertainties and bottlenecks need to be overcome. Among the candidate chemicals, methanol and syngas are most widely studied currently,1−7 while dimethyl carbonate (DMC) follows. DMC can be used as the important building block for different organic synthesis with major production of polycarbonates, as methylation agent, as solvent for lithium ion batteries, and as fuel additive. It also has low toxicity and fast biodegradability. Conventionally, DMC is produced through methanol phosgenation. Although the reactivity is high, the application of this method is still limited since the toxic nature © XXXX American Chemical Society
of phosgene. Therefore, other nonphosgene routes of DMC synthesis, majorly through oxidative carbonylation of methanol (i.e., Bayer process, Enichem process, and UBE process), have long been regarded as the replacement. In Bayer process,8 methanol, oxygen, and carbon monoxide are reacted in liquid phase, with CuCl and KCl serving as catalyst. The major drawback of this process lies in the low productivity, catalyst deterioration, and the reactor corrosion. In Enichem process, CO, methanol, and oxygen are reacted directly in vapor phase to become DMC.9 In this process, the catalyst deterioration is also the major issue that limits the per-pass conversion to be within 50%. In UBE process, alkyl nitrite is used as an oxidant to convert CO into DMC with the aid of palladium catalyst, and dimethyl oxalate is produced as side product.10 The major concern of this process is that nitric oxide is formed through the reaction, and its emission should be carefully regulated. The production of DMC from CO2 and methanol is also considered promising because it consumed the undesired chemical into another useful one. Many researchers have spent effort in developing catalyst and kinetic expressions,11,12 and in investigating the thermodynamic behavior for this direct synthesis route.13−15 However, the direct conversion of CO2 into DMC is severely limited by its thermodynamically stable Received: Revised: Accepted: Published: A
July 16, 2017 October 13, 2017 December 15, 2017 December 15, 2017 DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
Article
Industrial & Engineering Chemistry Research
evaluation include return on investment (ROI) and internal rate of return (IRR), while the amount of CO2 emission is calculated in environmental evaluation. For a consistent comparison, the DMC product purity is set at 99.5 mol % in both processes.
structure. It is reported that the equilibrium conversion of CO2 and methanol to DMC is within 1%, even under extremely high pressure (>100 bar). Thus, the existing CO2-based DMC processes are mostly through an indirect route. The indirect route is basically a trans-esterification reaction, in which a highly active oxygenated reactant (i.e., epoxides) is reacted with CO2 to form an intermediate carbonate, and it is further reacted with methanol to become DMC and the corresponding glycol as a byproduct. Among the trans-esterification routes, the most promising one is that via the generation of ethylene carbonate (EC) as an intermediate. In this method, CO2 is first reacted with ethylene oxide (EO) to form EC, and EC is further reacted with methanol (MeOH) to form DMC and ethylene glycol (EG).16,17 Because of the higher reaction conversion between EC and MeOH, this process route is more promising than other routes using different epoxides as the intermediate. In order to improve the reaction performance in direct conversion process, some intensification strategies can be applied. The first one is to utilize a membrane reactor for this reaction. Inside the reactor, water permeates through the membrane, while DMC, unreacted methanol, and CO2 remain as the retentate. Through this strategy, the reaction equilibrium can shift to the product side by removing the water byproduct. Kuenen et al. investigates the techno-economic evaluation of this process, yet they found that the performance is far from economically attractive due to the low reaction conversion (about 6.3%) even if a membrane reactor is applied.18 Another strategy for enhancing the direct conversion process is to add a dehydrating agent inside the reaction system. This dehydrating agent reacts with water to form a byproduct. Those possible dehydrating agents include acetonitrile (or other nitriles), trimethoxymethane, ionic liquids, or butylene oxide.19−21 Although the conversion can be greatly enhanced, the low to moderate selectivity toward DMC may make the separation more difficult. In a recent paper series, it is discovered that 2-cycaopyridine can be used as a water trap to form 2-pyrrolidone with very high yield and selectivity.22−26 However, the lack of detailed expression of reaction kinetic limited the research progress in conceptual design. In recent contributions, Kongpanna et al. studied the technoeconomic evaluation of different CO2-based DMC production processes, and they reported that the indirect process through EC route is the most economical and environmentally friendly route compared with others.27,28 However, the net carbon emission from ethylene carbonate route is positive (0.452 kg/ kg DMC), even if CO2 is used as feed stock to generate DMC. Hence, the process configuration and the operating condition inside this process still have room for improvement. Thus, in this work, a more intensified process through reactive distillation is provided, with a systematic optimization and heat integration performed to further fortify the process performance. This paper intends to focus on the detailed and rigorous simulation and optimization on two different CO2-based DMC production processes. One is the direct synthesis process with addition of butylene oxide (BO) as water trap (called “DIR-BO process” in the remaining text); the other is an indirect synthesis process through EC route (called “IND-EC process” in the remaining text). In both processes, rigorous design and optimization are investigated. Besides, heat integration is applied in order to make the energy usage more efficient. Finally, these two processes are evaluated economically and environmentally. The performance criteria for economic
2. OVERVIEW OF THE PAPER In this work, simulations for both processes are performed in Aspen Plus V8.8, and UNIQUAC method is selected as the thermodynamic model throughout this work. Components in DIR-BO process includes carbon dioxide (CO2), dimethyl carbonate (DMC), methanol (MeOH), aniline (ANI), methoxybutane (MB), butylene oxide (BO), butylene glycol (BG), and water (H2O), while in the IND-EC process, the components included are CO2, DMC, MeOH, ANI, ethylene oxide (EO), and ethylene glycol (EG). The binary parameter sets required to use in this work are listed in Tables S1 and S2. For the binary pairs not mentioned, the thermodynamic behavior is simulated as ideal. The Henry’s constants used to describe the dissolution behavior of CO2 into others are also listed in Table S3. After the steady-state design is complete, optimization is investigated. The objective function for optimization is to minimize total annual cost (TAC), which can be calculated with TAC (kUSD/yr) =
Capital Cost + Operating Cost Payback Period (1)
where the payback period is assumed to be 8 years, and the total annual operating time is 8000 h. Correlations provided by Turton et al. are applied to calculate the capital and utility costs.29 For evaluating the steam cost, the unit prices are corrected based on the steam price (21 bar and 215 °C) provided by Taiwan Institute of Chemical Engineers (TWIChE). Those steam prices after correction are listed in Table 1. Table 1. Utilities Used in This Work utilities
price (USD/GJ)
CO2 emission (kg/GJ)
steam (6 bar, 160 °C) steam (11 bar, 184 °C) steam (21 bar, 215 °C) steam (42 bar, 254 °C) cooling water electricity
11.80 13.28 14.14 15.73 0.354 16.9
72.86 76.60 82.16 91.14
For evaluating economic performances, CAPCOST, a calculation package provided by Turton et al., is used. Return on investment (ROI) and internal rate of return (IRR) are the major indices to be calculated. For calculating CO2 emission, the correlation from Gadalla et al. is selected and assumes that natural gas is the fuel used to produce different grades of steam.30 The CO2 emission amount per GJ of each grade of steam is also listed in Table 1. The details of the processes are illustrated, and the discussions are included in the following sections. The DIRBO process is discussed in sections 3, and the IND-EC process is discussed in section 4. The economical evaluation and carbon emission analysis is addressed in section 5. Section 6 gives the concluding remarks. B
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Figure 1. Optimal flowsheet of the DIR-BO process.
3. DIR-BO PROCESS 3.1. Reaction Pathway and Kinetic Expression. Eta et al. conducted the experiments for direct DMC production with butylene oxide (BO) as the dehydrating agent with the aid of ZrO2−MgO catalyst and developed a kinetic expression in Eley−Rideal mechanism. This serves as a solid basis for process development.21 In this reaction system, CO2 is dissolved into liquid methanol and then reacts with methanol to form the intermediate, monomethyl carbonate (MC). In contrast, the BO is also added into the system, and is reacted with methanol to become another reaction intermediate, methoxybutanol (MB). For these two reactions, the one between CO2 and MeOH is a catalytic reaction, and the other one is noncatalytic. For the intermediates, there are other two reactions accounts for DMC production. The first one is that MC reacts with MeOH to form DMC and water, while the other one is that MB reacts with MC to form DMC and BG. Eta et al. reported that methanol is adsorbed onto the catalyst, and the catalytic reaction between activated methanol and CO2 is the ratelimiting step.21 With this assumption and the quasi-equilibria for all other elementary steps, the reaction network can be simplified as the following three individual reactions: 2MeOH + CO2 → DMC + H 2O (2)
corresponding diol. However, BO does not react with water directly here. Instead, BO affects the reaction mechanism which leads to the inhibition of water production and the enhancement of the DMC production. The detailed mechanism could be found in Scheme 4 in Eta et al.’s work.21 In the kinetic expression provided by Eta et al., only the production (consumption) rates of DMC, MB, BG, BO, and MeOH are given. In order to apply this kinetic set, the reaction rate of CO2 and H2O should first be derived based on atomic balance. After that, the rate expressions are built into Aspen Fortran Subroutine for successful description. All the rate expressions and the parameters included in this study are listed in the following eq 5−18, which are directly referenced from the original work by Eta et al.21
BO + MeOH → MB
(3)
MeOH + CO2 + MB → DMC + BG
(4)
Equation 2 shows the direct reaction route for DMC formation, without involvement of BO. On the contrary, eqs 3 and 4 show the altered reaction mechanism with involvement of BO. As a dehydrating agent, BO is used to react with water to form the
dC DMC = r1ρB dt
(5)
dCMB = (r2 − γr1ρB )ω dt
(6)
dC BG = r2 − (r2 − γr1ρB )ω dt
(7)
dC BO = r2 dt
(8)
dCMeOH = −2r1ρB − (r2 − γr1ρB )ω dt
(9)
dCCO2 dt C
= −r1ρB
(10) DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Industrial & Engineering Chemistry Research Table 2. Information of Important Streams in the DIR-BO Process total flow (kmol/h) T (°C) P (bar) methanol CO2 BO DMC BG MB water aniline
MeOH
BO
CO2
41.30 150.0 45.0
17.84 150.0 45.0
20.62 150.0 45.0
1.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 COL1-LD
CF
0.0000 0.0000 0.0000 1.0000 1.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 COL1-B COL2-D
total flow (kmol/h) T (°C) P (bar)
544.83 67.0 1.5
26.13 167.5 1.8
8.38 116.1 1.0
17.75 211.0 1.1
methanol CO2 BO DMC BG MB water aniline
0.9560 0.0012 0.0047 0.0380 0.0000 0.0000 0.0001 0.0000
0.0211 0.0000 0.0000 0.0000 0.6797 0.1884 0.1106 0.0002
0.0657 0.0000 0.0000 0.0001 0.0021 0.5864 0.3454 0.0003
0.0000 0.0000 0.0000 0.0000 0.9990 0.0009 0.0000 0.0001
dC H 2 O dt
= r1ρB − r2 + (r2 − γr1ρB )ω
3.45 4.93 84.5 156.3 1.0 1.1 Mole Frac 0.1593 0.0000 0.0000 0.0000 0.0000 0.0000 0.0002 0.0000 0.0000 0.0031 0.0014 0.9964 0.8390 0.0001 0.0000 0.0004
(12)
r2 = k 2C BOCMeOH
(13)
⎛ −62000 ⎛ 1 1 ⎞⎞ ⎜ ⎟⎟ k1 = 1.6 × 10−5 exp⎜ − ⎝ R ⎝T 423.15 ⎠⎠
(14)
k 2 = 0.0112
(15)
⎞ ‐1 ⎛ α(y + 1)C DMC ⎟ 1 ω=⎜ + 2 ⎠ ⎝ (α + y) CMeOH
(16)
⎛ y ⎞⎛ ⎛ 2α ⎞ C DMC ⎞ ⎟⎟ γ=⎜ ⎟⎜⎜1 − ⎜ ⎟ ⎝ α + y ⎠⎝ ⎝ α + y ⎠ CMeOH ⎠
(17)
CMB CMeOH
VAP-R
LIQR-2
Flash-V
Flash-L
COL1-VD
436.62 150.0 45.0
595.55 92.7 45.0
14.19 47.0 1.5
581.36 82.8 2.0
10.41 67.0 1.5
0.2765 0.7136 0.0018 0.0070 0.0001 0.0001 0.0010 0.0000 COL4-D
0.8950 0.0215 0.0046 0.0358 0.0298 0.0083 0.0049 0.0000 water
0.2965 0.9096 0.6814 0.0054 0.0045 0.0046 0.0170 0.0363 0.0000 0.0305 0.0000 0.0085 0.0006 0.0050 0.0000 0.0000 MKP ENT EDC-D
0.7130 0.2410 0.0095 0.0365 0.0000 0.0000 0.0000 0.0000 DMC
0.54 64.2 1.0
2.90 98.9 1.0
0.00 47.0 1.5
490.34 47.0 1.5
524.10 54.5 1.0
20.67 89.4 1.0
0.9950 0.0000 0.0000 0.0013 0.0000 0.0000 0.0037 0.0000
0.0033 0.0000 0.0000 0.0000 0.0000 0.0017 0.9950 0.0000
0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 1.0000
0.0000 0.0000 0.0000 0.0001 0.0000 0.0000 0.0000 0.9999
0.9937 0.0012 0.0048 0.0002 0.0000 0.0000 0.0001 0.0000
0.0010 0.0000 0.0040 0.9950 0.0000 0.0000 0.0000 0.0000
3.2. Process Development. The optimized flowsheet of DIR-BO process is illustrated in Figure 1, with the information of important streams listed in Table 2. First, fresh CO2, MeOH, and BO are mixed with several recycled streams and then sent to the reactor. Those recycled streams include the vapor outlet from several later units (reactor, flash unit, and later column 1), the unreacted liquid methanol, and MB. The optimal operating condition for reactor mainly follows Eta et al.’s work,21 and the reactor is modeled as a continuous stirred-tank reactor (RCSTR module). For the composition of combined feed sent into reactor, CO2 and MeOH are set at near stoichiometric ratio, and the molar ratio between MeOH and BO is set at 30. The molar ratio between MeOH and BO leads to a trade-off between reaction conversion and selectivity. With less BO added, the reaction conversion is barely enhanced, while with more BO, it may lead to the generation of the undesired product, MB, as more reaction intermediate accumulates in the reactor. For operating temperature, 150 °C is also selected based on the trade-off between reaction conversion and selectivity. Another undesired side product, dimethoxymethane, would be generated easily at temperature higher than 150 °C. For operating pressure, it seems that the reaction conversion is not sensitive to the operating pressure. Thus, the lowest one, 45 bar, mentioned in Eta et al.’s work21 was selected for the sake of easier industrial application. The reactor holdup is designed to have 90% liquid and the remaining vapor. On the basis of the experimental results provide by Eta et al.,21 the conversion of each species becomes much slower after 12 h. Hence, the liquid-phase residence time is reasonably set at this value. The liquid reactor outlet is first heat-exchanged with the liquid stream from later flash unit, and is further cooled to 47 °C and decompressed to 1.5 bar. The reason to cool the outlet stream to 47 °C is because this temperature is considered as the
(11)
⎛ C C ⎞ r1 = k1⎜CMeOH − DMC BG ⎟PCO2 KCMB ⎠ ⎝
y=
LIQ-R
1070.55 595.55 123.2 150.0 45.0 45.0 Mole Frac 0.6492 0.8950 0.3223 0.0215 0.0199 0.0046 0.0035 0.0358 0.0000 0.0298 0.0046 0.0083 0.0004 0.0049 0.0000 0.0000 BG COL3-D COL3-B
(18)
where the concentrations (Ci) are in unit of mol/liter, and PCO2 is in bar. For the parameters, the equilibrium constant (K) is 0.0036 (bar), the bulk catalyst density (ρB) is 33.3 (kg/m3), and the value of α in eq 17 is 0.0014. Because the values of K and α are not specified in the original work,21 these two values are obtained by re-regression from the experimental data at the temperature of 150 °C, which is the most suitable reaction temperature suggested by the authors. Note that k2 is a constant. This is potentially because Eta et al. focused more at T = 150 °C and only obtain the kinetic parameter under this temperature. D
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Industrial & Engineering Chemistry Research
by a preconcentrator column may not be economical. Thus, this stream is directly sent to the extractive distillation section. In a typical extractive distillation process, two columns are required, including an extractive distillation column (EDC) and a solvent recovery column (SRC). For the MeOH and DMC separation system, MeOH is obtained at the top of EDC and DMC comes out from top of SRC. The MeOH separated from EDC is pumped and then recycled to the reactor inlet. Note that there is a 0.01% purge of EDC distillation to avoid the water accumulation. In this work, three different extractive distillation configurations are discussed and compared. The configuration of traditional extractive distillation configuration is illustrated in Figure 3a. In order to be more energy-saving, a simple heat-integration can be achieved by adopting a feed− effluent heat-exchanger (FEHE), as shown in Figure 3b. This FEHE performs heat exchange between the SRC bottom stream and the feed stream to EDC. Through this design configuration, the reboiler duty of EDC can be greatly reduced. A further heat-integration strategy is to configure it into an extractive dividing-wall column system, as shown in Figure 3c. The design of a dividing-wall column was widely discussed in literature. From previous literature, it is known that it has the advantage to reduce the remixing effect in a typical distillation column and can improve the separation efficiency. There are also different configurations for a dividing-wall column. In this work, as the products are obtained from the distillate streams in different columns, a dividing-wall configuration that integrates two different reboilers is used. Figure 3d shows the thermodynamically equivalent configuration of an EDWC configuration. For simulation purpose, the whole column can be divided into three parts, and each part is designed separately. In this subfigure, the side rectifier C1 represents the original EDC, while C2 and C3 performed as the original SRC. In operation, the vapor traffic of C1 is provided by a portion of the overhead vapor from C3. Besides, the diameter of the upper part of EDWC can be determined by the equivalent-diameter method as in the previous work by Wu et al.32 The brief optimization results for these three configurations can also be found in Figure 3a,c. When operating at entrainerto-feed molar ratio (E/F) equal to 0.9, a 29.8% TAC reduction is achieved with adoption of FEHE. When EDWC configuration is adopted, a further 13.4% reduction in TAC is obtained. 3.3. Optimization Work. In the whole process, COL-1 and the extractive distillation section are needed to be optimized. For COL-2, COL-3, and COL-4, the TAC change with varying design and operating variables may be too slight to take into consideration. This is because the throughputs of these two columns are very small compared with COL-1 and the extractive distillation section. In COL-1, there are three variables to be optimized, including the condenser temperature (Tcond), total number of stages (NT-COL1), and feed locations (NF-COL1). The optimization work can be achieved by a simple sequential iterative method, with the algorithm shown in the left part of Figure S1. The results are recorded in Figure 4a. From this figure, it is shown that the optimal point is at Tcond = 67 °C, NT‑COL1 = 50, and NF‑COL1 = 38. After the optimization of COL1, the heat integration between flash liquid and reactor effluent is adopted to save energy. Next, the extractive distillation section is optimized, with the algorithm shown in the right side of Figure S1. In this section, six variables need to be optimized, including entrainer-to-feed
lowest allowable one by using the cheap cooling water. This operation enables the easier separation for the unreacted CO2 with MeOH and DMC in the flash unit. After that, the liquid stream from the flash unit is sent to column 1 (COL-1), in which MeOH and DMC are obtained from liquid distillate, while H2O, MB, and BG come out at the bottom. Besides, the adoption of a partial vapor−liquid condenser enables the noncondensed gas to be separated. This vapor distillate is compressed and recycled to the reactor inlet. The bottom stream of COL-1 is then sent to column 2 (COL-2), column 3 (COL-3), and column 4 (COL-4) for further separation. In COL-2, the side product BG is obtained at the bottom, while the remaining species comes out as liquid distillate. In COL-3, the unreacted MB is obtained at the bottom and is recycled to reactor inlet, while the liquid distillate contains mostly water. Note that from the boiling point ranking of this system, water lies in the middle, which makes it difficult to separate. However, in order not to retard the reaction equilibrium, the removal of this slight amount of water is indispensable. In COL-4, the unreacted methanol is separated from water and sent back to reactor inlet, in order to fulfill better recovery of it. The liquid distillate from COL-1 contains 95.6 mol % MeOH and 3.8 mol % DMC, and the balance is BO. This stream is sent to the downstream for further separation. However, MeOH forms an azeotrope with DMC, thus the mixture cannot be separated using a single column. The T−xy diagram of MeOH and DMC is shown in Figure 2. There are
Figure 2. T−xy diagram between MeOH and DMC under 1 bar.
many contributions focused on separation of MeOH and DMC in literature, and it can be concluded that extractive distillation is better than pressure-swing distillation because the azeotropic composition of MeOH and DMC does not vary much under different pressures.17,18,31,32 Hsu et al. reported that aniline is an effective entrainer to enhance the relative volatility between MeOH and DMC; thus, it is selected as the entrainer in the further separation section in this work.17 From the T−xy diagram, it is noted that the composition of COL-1 liquid distillate lies in the right-hand side of azeotropic composition. However, due to the narrow vapor−liquid envelope in this region, purifying to the azeotropic composition E
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Figure 3. Different process configurations for MeOH/DMC separation via extractive distillation
molar ratio (E/F)), rectifying section stages of EDC (NREC), extractive section stages in EDC (NEXT), stripping section stages of and EDC (NSTR), total number of stages of SRC (NT-SRC), and feed location in SRC (NF-SRC). After the optimal point under each E/F is found, the FEHE is added in order to save energy. The optimization results without FEHE are collected as in Figure 4b, while the results with FEHE are shown in Figure 4c. From these subfigures, it is shown that a lower E/F leads to lower TAC, but an optimal point at E/F = 0.85 is found in the cases with FEHE. In EDWC, there are also several variables to be determined, including E/F, total number of trays of C1, C2, and C3 (NT1, NT2, and NT3), two feed locations of EDC feed (NF1 and NF2), and the vapor split ratio (FV). Those symbols are all mentioned in Figure 3d. Among them, FV is used to make sure
that BO and methanol left in bottom product from C1 are low enough to provide DMC from distillate of C2. The remaining six variables are optimized. The algorithm for sequential iterative method is shown in Figure S2, and the optimization results are recorded in Figure 4d. From this figure, it is found that the optimal E/F is 0.9. It is also found that E/F changing leads to little variation in TAC. Comparing with the traditional configuration with FEHE, the optimal E/F found in EDWC also shifts somehow to a larger value.
4. IND-EC PROCESS 4.1. Reaction Pathway and Kinetic Expression. In the IND-EC process, the overall reaction is divided into two steps which occur in different reaction sections. The reaction equations are F
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Industrial & Engineering Chemistry Research
Figure 4. Optimization results of the DIR-BO process.
Figure 5. Optimal flowsheet of the IND-EC process.
EO + CO2 ↔ EC
(19)
EC + MeOH ↔ DMC + EG
(20)
For eq 19, the reaction occurs under high-pressure condition, with the conversion toward EC favored at lower temperature. Compared with eq 20, the reaction rate of eq 19 is relatively G
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Industrial & Engineering Chemistry Research Table 3. Information of Important Streams in the IND-EC Process CO2
EO
rout
EC
EC-in
MeOH-F
MeOH-R
RD-D
EG
ED-in
ANI
MKP
SRC-B
DMC
total flow (kmol/h) T (°C) P (bar)
22.01
20.01
22.01
20.00
20.00
40.16
91.85
111.93
20.09
111.93
85.07
0.01
85.07
20.09
50.0 50.5
50.0 50.5
80.0 50.5
202.5 0.3
47.0 1.3
45.0 1.5
63.5 1.0
201.4 1.3
74.4 1.5
54.0 1.1
54.0 1.2
157.1 1.4
89.1 1.0
EO CO2 EC DMC EG MeOH aniline
0.0000 1.0000 0.0000 0.0000 0.0000 0.0000 0.0000
1.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000
0.0001 0.0910 0.9089 0.0000 0.0000 0.0000 0.0000
0.0000 0.0000 1.0000 0.0000 0.0000 0.0000 0.0000
0.0000 0.0000 1.0000 0.0000 0.0000 0.0000 0.0000
0.0000 0.0000 0.0000 0.0000 0.0000 1.0000 0.0000
0.0000 0.0000 0.0000 0.1786 0.0000 0.8214 0.0000
0.0000 0.0000 0.0005 0.0000 0.9950 0.0043 0.0002
0.0000 0.0000 0.0000 0.1786 0.0000 0.8214 0.0000
0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.9999
0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 1.0000
0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.9999
0.0000 0.0000 0.0000 0.9950 0.0000 0.0046 0.0004
64.3 1.5 Mole Frac 0.0000 0.0000 0.0000 0.0001 0.0000 0.9999 0.0000
distillation section, the major target is to keep EC conversion high. Hence, the targeted EC conversion is set at 99.95%, and the reacting condition with excess in MeOH could be helpful. The distillate stream from RDC is then sent to the downstream extractive distillation process to separate DMC and MeOH. The overall process configuration is the same as the one mentioned in previous Figure 3a but with two differences. The first difference is that in this process, the feed sent to extractive distillation section has composition of 82.13 (mol %) MeOH and 17.87 (mol %) DMC, which contains less MeOH in this feed stream. For comparison purposes, the composition of MeOH in the feed stream of Figure 3a is at 95.60 mol %. This difference will reflect on the required E/F ratio in the EDC to drive MeOH up at this column. As a result, the optimal E/F ratio is at 0.90 in Figure 3a, while for this case, E/F is lower at 0.76. The second difference also arises due to the difference of feed compositions in these two cases. With much less DMC (only 3.8 mol %) in the feed stream of Figure 3a, this lead to results that the same grade of steam is required in both EDC and SRC. Thus, further integrate these two columns into an extractive-dividing-wall column (EDWC) can greatly improve the economic performance. However, in this case shown in Figure 5, the cheaper 11 bar steam can be used as the reboiler heat source in EDC, while the 21 bar steam should be used in SRC. If it is integrated to an EDWC configuration, then a higher grade of steam should be used for total required energy. As already concluded in a previous work by Wu et al.,32 this leads to an increase in overall utility cost, even if the total reboiler duty is reduced. Hence, for indirect synthesis route, the EDWC configuration is not utilized. 4.3. Optimization. Inside this process, the RDC and the extractive distillation section are the most important part to be optimized. Four variables in RDC should be determined, namely, excess ratio between feed MeOH and EC (ER), total number of stages (NT-RDC), the rectifying section stages (NREC) and reaction section stages (NRXN). While the variables discussion in extractive distillation section is the same as those mentioned in section 3.3. ER is the most important parameter in this process, and its influences can be addressed in two aspects. First, larger ER enhances the reaction conversion from the thermodynamic viewpoint. The specified reaction conversion of EC is very high (99.95%), thus there should be a minimum ER value to reach the targeted conversion. Second, ER affects the distillate composition of RDC, which leads to different separation performance in extractive distillation section. The process provided by Hsu et al. has ER = 7.8,17 and this made the RDC distillate composition to be 85.29% MeOH, which is pretty
faster. However, to the best of our knowledge, there is no appropriate kinetic expressions for this reaction provide in literature. For simulation purposes, Souza et al. used an imaginary kinetic expression for this reaction which is fast enough.33 In this work, we did not make such an extra assumption but modeled eq 19 as an equilibrium reaction instead. The equilibrium constant of this reaction is predicted directly in Aspen Plus based on Gibbs energy minimization. Similar to the results by Souza et al., our simulation setting can also lead to an almost-complete reaction conversion. Note that in some recent publications, the concept of onepot synthesis was proposed, in which eqs 19 and 20 are designed to take place in the same reactor,34,35 but considering the different reaction condition requirement of these two reactions, the process configuration with two individual reaction sections seems to be more appropriate. For eq 20, the kinetic expression is referenced from Fang and Xiao16 and is illustrated in the following equations: rate = k f C ECCMeOH −kr
C EGC DMC CMeOH
(21)
⎛ −13060 ⎞ ⎟ k f = 1.3246 exp⎜ ⎝ RT ⎠
(22)
⎛ −28600 ⎞ ⎟ kr = 15022 exp⎜ ⎝ RT ⎠
(23)
where the reaction rate is in mol/L·min, concentration is in mol/L, and activation energies are in J/mol. 4.2. Process Development. The optimized flowsheet of the IND-EC process is provided in Figure 5, and the information on streams is collected in Table 3. Streams of fresh CO2 and EO are mixed and sent to the reactor. The reactor is operated at 50.5 bar and 80 °C, and EO conversion reaches over 99.99%. The reactor effluent is decompressed to 1.5 bar, and then sent to a stripper to distill off the remaining CO2 inside the EC-rich liquid. Almost all the CO2 comes out from the stripper overhead vapor, and this stream is not recycled to the reactor in order to avoid the costly compressing step. In the stripper bottom, the high-purity EC stream is obtained. Because of the high boiling point of EC, the stripper is operated at 0.3 bar, in which the 21 bar steam can be utilized. The EC stream obtained from the previous reaction section is cooled and fed into the reactive distillation column (RDC) at the upper section. The fresh and unreacted MeOH stream is also fed to RDC at the lower section. After reaction, the mixture of methanol and DMC comes out as the liquid distillate stream, while EG is obtained at the bottom. In reactive H
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Figure 6. Optimization results of the IND-EC process.
this specification can no more be reached. From our results, the lower limit for ER is found to be 6.6. This reduction only leads to a decrease of MeOH content in EDC distillate from the original 85.29 mol % to 82.13 mol %, and is the major reason for the nonexistence of the trade-off between RDC and the downstream extractive distillation section. As for E/F in extractive distillation section, the optimal point is at 0.76. E/F is an important design parameter in the extractive distillation process which affects the separation efficiencies. Hence, an obvious trade-off in capital cost and operating cost may be well expected by varying E/F. However, from Figure 6b, the TAC variation from changing E/F is not large. The reason is that we start the optimization from the previous optimal result (E/F = 0.883) obtained by Hsu et al., in which the process is optimized based on Douglas’ work with payback period set at 3 years.17,38 During the reinvestigation of optimization based on different cost correlations and payback years, we actually take a deeper look on the effect of varying E/ F; thus, the moderate change in TAC is reasonable. In summary, comparing this process to Hsu et al.’s work, a 16.8% reduction on TAC is observed.17 The detailed information for this comparison can be found in Table 4. Note again that the process proposed by Hsu et al. was optimized based on the Douglas’ work,38 in which the utility cost is much cheaper than the ones suggested by Turton et al. Thus, we suggest that optimization based on Turton et al.’s model leads to an optimal solution with lower utility requirement, and is more suitable in evaluating CO2 reduction.
close to the azeotropic composition (86.5 mol % MeOH) at 1 bar. As ER decreases, the loading of RDC and EDC can be directly reduced; this may lead to better overall economic performance. Note that the process configuration in this work is somewhat similar to the typical preconcentrator column followed by an azeotropic separation section (i.e., extractive or heterogeneous azeotropic distillation system), as the RDC provided a distillate stream with near azeotropic composition. Luyben discussed the effect of preconcentrator distillate composition on the performance of ethanol−water separation based on heterogeneous azeotropic distillation system.36 The effect in the same ethanol−water system was also discussed by Kiss and Ignat based on extractive distillation system. 37 The general conclusion is that the lower distillate composition saves the energy requirement in the preconcentrator column yet increases that in the downstream azeotropic separation section. Thus, the distillate composition plays an important role in these processes. As the feed ratio into RDC apparently affects its distillate composition, it should be carefully determined through optimization. The overall optimization is also performed by a sequential iterative method, with ER as the outermost variable. Other variables are located inside the network. The overall algorithm is illustrated in Figure S3. The optimization results are illustrated in Figure 6, and only the major results are included. The general results show that TAC drops obviously as ER decreases in the region we investigated, and from this we conclude that the trade-off mentioned above does not exist. The reason for this is because there is a stringent requirement in reaction conversion (99.95%) for EC in RDC. Thus, once ER becomes too small, I
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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another widely used correlations provided by Sieder et al. also assumed in a similar way.39 5.1. DIR-BO Process. In this process, the raw materials are CO2, MeOH, and BO, with the unit prices set to be 70 (USD/ MT), 310 (USD/MT), and 1600 (USD/MT). There is no open market data for BO price. Traditionally, BO is synthesized as the byproduct of ethylene oxide (EO) and propylene oxide (PO) through the hypochlorination reaction. Thus, the authors assume that the production cost of BO is similar to the EO price (1510−1700 USD/MT) and PO price (1680−1810 USD/MT),40,41 and in base case calculations, it is set at 1600 (USD/MT). The products are DMC and BG, with the base case prices set at 685 (USD/MT)42 and 2200 (USD/MT).43 However, due to the uncertainties and volatilities in unit prices, the evaluation results may vary greatly as they change. Recent financial news in China stated that the unit price of DMC has dropped from about 900 USD/MT in January to less than 700 USD/MT at the end of April in 2017.44 The reason for this drop was due to the price fluctuation of the upstream propylene oxide and methanol. This phenomenon clearly shows the volatility of DMC market. Also, in the marketing report provided by ICIS.com, there is also a fluctuation region for the prices. Thus, the uncertainties resulted from different BO and BG unit price should also be taken into consideration. In this work, the price uncertainties are investigated through a series of sensitivity analyses. The interest region for each price is stated as below. DMC unit price is varied from 100−150% base case value, BO price is fluctuated from 95−105% of its base case value, while the price of BG is studied in between 90−100% base case value. For CO2 and MeOH, the sensitivity analysis on their price is not studied, as these two components are very common chemicals around the world. In this work, the authors consider an economically attractive case should have ROI > 15% and IRR > 10%. The analysis result of base case is listed in Table S4, and the sensitivity analysis results are shown in Figure 7. For the base case, it is clear that it just approaches the break-even point, which is not considered economical. However, as BO price drops a bit and the DMC price rises a bit, the economic performance is somewhat improved. Given the DMC price trend and the reason for its dropping mentioned above, it is likely that the DMC unit price can bounce back in the future. However, in order to be economically attractive, the DMC price has to be 40−50% higher than the base case value. Thus, in the current stage, the economical attractiveness is not obvious. Besides, the reactor in this process is very expensive. As listed in Table S4, it accounted for almost 60% of total capital investment (total grass root cost in the table), which is not the typical case of an industrial plant. In the contribution by Eta et al, the kinetic parameters were obtained under a reactor condition with bed void fraction around 97% (assume the catalyst density is 1000 kg/m3). Therefore, the catalyst amount has some room to be increased. By this operation, the residence time required can be greatly reduced and can save the capital investment. However, due to the lack of kinetic parameters, the further improvement on the process is not continued. Instead, we proposed a hypothetical situation that the reactor size can be reduced and investigate the influence of it on the overall economic performance of the base case. The analysis result is shown in Table 5. If the reactor size can be reduced, then the enhancement in economic performance can be apparent, as about 17% net increment in ROI and 14% increment in IRR can be obtained as the reactor size is reduced to 10% of its
Table 4. TAC Comparison between the Optimal Process and the Results of Hsu et al.17 ratio E/F
capital RD operating
RD-vessel RD-tray cond REB steam CW
RD TAC
capital ED
operating
total
ED TAC total TAC
ED-vessel ED-tray ED-cond ED-REB SR-vessel SR-tray SR-cond SR-REB HX1 HX2 steam ED steam SR CW
Hsu et al.
this work
7.8 0.883
6.6 0.760
156.22 50.31 77.11 117.44 740.71 18.36 809.20 149.79 49.48 72.92 89.36 57.13 24.36 50.60 63.15 50.72 53.81 573.13 246.10 22.13 924.02 1733.22
189.41 69.49 73.55 123.03 660.31 16.00 733.25 176.58 79.20 65.99 74.61 54.87 29.60 50.31 61.49 52.03 51.76 410.13 195.14 16.32 708.63 1441.89 (−16.81%)
5. ECONOMICAL EVALUATION AND CARBON EMISSION CALCULATION After design and optimization of the steady-state simulation are finished, economical evaluation is investigated. The economical evaluation is performed by the software package CAPCOST, which is provided by Turton et al. For the indices to be evaluated, both the nondiscounted profitability criteria (ROI) and discounted profitability criteria (IRR) are calculated. In calculating IRR, a 10% commercial interest rate is assumed to discount the value of money as time goes on. The items for economical evaluation can be roughly divided into two parts, namely, capital cost and operating cost. In determination of capital cost, the size, material, and operating condition of each process unit should be known. For evaluating the operating cost, the most important part could be the cost of utilities, operating labor, and raw materials. Also, the analysis results may be highly dependent on the financial model; thus, the parameters inside the model should be carefully set. Turton et al.’s capital cost correlations were regressed from the actual cost information from vendors. Thus, in each correlation, there is a feasible regions and is bounded by the upper and lower limit of the equipment size. However, for some units, the feasible size range is relatively narrow (i.e., reactor, kettle reboiler, etc.), and some units in this work are designed to have their sizes exceed the upper limits under the given category. Therefore, the costs for those units are assumed to be calculated based on the correlations of other units with similar type. The assumptions are stated as the following. First, the reactor cost is calculated by the correlation used for horizontal vessel. Second, the reboiler cost are calculated as floating type heat exchanger. Lastly, the double-pipe category is assumed for all the heat exchangers with exchanging area less than 10 m2. The authors consider these assumptions valid, as J
DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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Figure 7. Sensitivity analysis results of ROI and IRR on the DIR-BO process.
Figure 8. Sensitivity analysis results of ROI and IRR on the IND-EC process.
Table 5. ROI and IRR at Different Reduced Reactor Sizesa reactor ratio
ROI (%)
IRR (%)
100% 20% 10%
−0.67 11.55 16.68
X 11.22 14.56
process, which can be viewed as the total capital investment, and is the basis for calculating ROI and IRR. A common point can be concluded from these two processes, that is, the revenue provided by the side-product (BG or EG) plays an important role in the economic analysis. Thus, a more conservative ROI and IRR target should be suggested in the economical evaluation in order to deal with the price fluctuation. 5.3. Carbon Emission. Beside economical evaluation, the emission of carbon dioxide is also an important criterion to be compared. Being different from the economical evaluation, the analysis of CO2 emission is independent to production rate. In this work, the calculation of carbon emission is based on the correlation provided by Gadalla et al, and natural gas is assumed to be used as the fuel to produce steam.30 The emission amount per GJ of different grades of steam is already collected in Table 1, while the analysis results of the two processes studied in previous sections are listed in Table 6.
a
Calculated as DMC unit price = 685 USD/MT, and BO unit price = 1600 USD/MT and BG unit price = 2200 USD/MT.
original size. Thus, the authors consider this process has the potential in the future, as long as the details in reaction section can be further studied and clarified. 5.2. IND-EC Process. In this process, the raw materials are CO2, MeOH and EO, while the products are DMC and EG. The unit prices for CO2, MeOH, and the base case unit price for DMC are set at the same values as in previous section. As referenced from ICIS.com, the base case unit prices of EO and EG are set as 1.56 (USD/kg) and 1.2 (USD/kg), respectively. Again, sensitivity analysis is performed on varying EO (95− 105%), EG (90−100%), and DMC price (95−105%) in order to investigate the uncertainties. The evaluation result of the base case is listed in Table S5, while the sensitivity analysis results are illustrated as in Figure 8. From the results, it is clear that the base case is pretty economically attractive with 26.55% ROI and 17.01% IRR. The results could be well-expected as the reaction conversion is much higher as comparing with the DIR-BO process, which leads to less system loading. Because of this, the sensitivity analysis focuses more on lowering the profit from products. From the sensitivity results, the general trend is the same as that in the DIR-BO process. However, it is found that ROI and IRR change dramatically compared with the previous case. This is resulted from the much less total grass roots cost in this
Table 6. Comparison of CO2 Emission Amounta DIR-BO emission (COL-1 Reb) emission (COL-2 Reb) emission (COL-3 Reb) emission (COL-4 Reb) emission (DWC-Reb) emission (Reactor) consumption net emission a
K
IND-EC 1.387 0.041 0.062 0.017 1.066 0.156 −0.487 2.242
emission emission emission emission
(STR Reb) (RD Reb) (EDC Reb) (SRC Reb)
consumption net emission
0.020 0.265 0.172 0.079
−0.487 0.049
Unit: (ton CO2)/(ton DMC generated). DOI: 10.1021/acs.iecr.7b02923 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX
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heat-integrated. Thus, converting it into DMC can barely reach net CO2 reduction and may not have too much upside for further improvement.
Table 6 illustrates that CO2 is in net emission in the DIR-BO process. Given the fact that CO2 is used as feed stock to produce DMC, 2.18 ton of CO2 is emitted as 1 ton of DMC is produced. The reason is that the reaction conversion is still low, with only about 10% of total methanol conversion. Although it is already improved a lot compared with the direct production process without BO addition (
