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A combination of pressure-swing and extractive distillation for separating complex binary azeotropic system
Accepted Manuscript Title: A combination of pressure-swing and extractive distillation for separating complex binary azeotropic system Authors: Xiaoxin Gao, Biyun Zhu, Jiangquan Ma, Deming Yang PII: DOI: Reference:
S0255-2701(17)30226-X https://doi.org/10.1016/j.cep.2017.10.012 CEP 7096
To appear in:
Chemical Engineering and Processing
Received date: Revised date: Accepted date:
10-3-2017 23-6-2017 14-10-2017
Please cite this article as: Xiaoxin Gao, Biyun Zhu, Jiangquan Ma, Deming Yang, A combination of pressure-swing and extractive distillation for separating complex binary azeotropic system, Chemical Engineering and Processing https://doi.org/10.1016/j.cep.2017.10.012 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
A combination of pressure-swing and extractive distillation for separating complex binary azeotropic system Xiaoxin Gao,a,b,* Biyun Zhu,b Jiangquan Ma,a,b Deming Yanga,b a. Jiangsu Key Laboratory of Advanced Catalytic Materials and Technology, Changzhou University, Changzhou 213164, P.R. China b. College of Petrochemical Engineering, Changzhou University, Changzhou 213164, P.R. China
AUTHOR INFORMATION Corresponding Author E-mail:[email protected]. Tel: +86-519-86330255
Highlights
The double column pressure-swing with extractive distillation(DCPSED) was proposed.
The triple column pressure-swing with extractive distillation were proposed.
The heat integration and MVR heat pump distillations were also tested.
The DCPSED with MVR process is used to separate complex binary azeotropic system.
ABSTRACT: Methyl acetate/methanol/water mixture forms more than one different azeotrope, whereas its triangular diagram presents a distillation boundary at atmospheric pressure. The two different simulation processes of the double column pressure-swing with extractive distillation (DCPSED) and triple column pressure-swing with extractive distillation (TCPSED) were proposed to separate the complex ternary system. Furthermore, the heat integration and 1
mechanical vapor compression (MVR) heat pump distillations were also tested (through simulation) to separate the complex ternary system to save energy. The feasibility of the processes was confirmed using vapor-liquid equilibrium curve, while rigorous steady-state simulations were implemented using Aspen Plus (version 7.1). With the goal of achieving minimum energy consumption and minimum total annual cost (TAC), the purities of methyl acetate, methanol and water were used as the constraint variables. The results show that water was a better solvent for this separation system. Compared with the TCPSED process, the energy consumption and TAC for DCPSED process could be reduced by 44.83%, and 47.74%, respectively. Compared with the DCPSED and DCPSED with heat integration processes, DCPSED with MVR process could reduce the energy consumption by 73.34% and 59.52%, while TAC was decreased by 37.64% and 11.77%, respectively. The simulation results indicated that DCPSED with MVR process has great advantage over other processes used to separate methyl acetate/methanol/water ternary mixture. Keywords: Azeotropic compositions; Pressure-swing distillation; Extractive distillation; Heat integration; MVR technology 1. Introduction Methanol (CH3OH) and methyl acetate (C3H6O2) are extensively applied as effective organic solvents in chemical and pharmaceutical industries owing to their excellent physicochemical properties [1]. Due to great application prospects in industry, it is attractive and necessary to separate and reuse methanol and methyl acetate to comply all processing plants with modern and rigorous environmental and resource-saving regulations. However, high-purity separation of ternary mixture of methyl acetate/methanol and methyl acetate/water azeotropes could not be
2
efficiently accomplished in common practical operation. Due to this, the lower purity of methyl acetate/methanol is likely to be at the cost of higher reflux and larger number of theoretical stages by conventional extractive distillation process [2]. Distillation is the most widely used separation process in chemical industries, whereas pressure-swing distillation (PSD) and extractive distillation (ED) are the most common methods for separation of a binary homogeneous azeotrope[3]. The ternary system containing two azeotropes has attracted more and more research attention in recent years [4-9]. In comparison to some alternative methods for separation of ternary azeotropic systems, such as azeotropic distillation [10], reactive distillation [11-12], liquid-liquid extractive distillation [13], salt addition [14], membrane distillation [15] and some other new coupling separation techniques, PSD and ED are broadly applied in recently years [16-19]. Although the membrane separation technology has the advantage of energy-saving in the process of breaking the azeotropic composition, the investment for the equipment cost is very high and commercial methods need to be further developed in this regard [15]. In azeotropic distillation, a new component was introduced as an entrainment agent, which resulted in more energy-consumption to recycle the entrainer. Zhu et al. [5] studied the separation of acetonitrile/methanol/benzene mixture via triple column pressure-swing distillation and analyzed the residue-curve map to illustrate the relationship of pressure and distillation boundary. Kaam et al. [20] investigated the feasibility and conducted
experiments
about
the
heterogeneous
extractive
batch
distillation
of
chloroform/methanol/water system, especially focusing on the entrainer selection for the separation of chloroform-methanol mixture. Luyben [21] introduced a comparison of extractive distillation and pressure-swing distillation for acetone/chloroform separation, and reported that the
3
pressure-swing distillation is much more expensive than the extractive distillation for acetone/chloroform separation in terms of both capital investment and energy consumption without considering product contamination. Lladosa et al. [22] proposed that the pressure-swing distillation process is effective to some extent, provided the azeotropic composition changes significantly with pressures. For ED, the key is the selection of solvent, and therefore, the ED process is effective if a suitable solvent can be found. ED requires an appropriate solvent to improve the relative volatility of the separated components, so that the mixture can be separated more easily and effectively. Pressure-swing distillation (PSD) uses pressure difference between several columns and changes the composition of azeotrope to effectively separate the mixture, and therefore, avoids inducing a third component to achieve energy saving. Gomis et al. [23] successfully set gasoline fraction mixtures as entrainers to implement ethanol dehydration, and conducted a pilot-scale industrial test to produce bioethanol and naphta. Deorukhkar et al. [24] used dimethyl sulfoxide (DMSO) as the entrainer to purify tetrahydrofuran (THF) from its aqueous azeotrope through continuous distillation, which recycled THF with a purity greater than 99 wt.%. Wang et al. [25] studied the design of pressure-swing distillation for the mixture of n-heptane and isobutanol with steady and dynamic simulations in the commercial software package Aspen Plus (version 7.2). By comparing the conventional PSD (CPSD) with unusual PSD (UPSD), the authors [23] proposed that the two processes should be applied according to their advantages. For a complex and special azeotropic system, conventional distillation could not achieve a desired separation, due to which, a combination of pressure-swing distillation and extractive distillation was studied in this paper. The change of pressure has great impact on the azeotropic composition and azeotropic temperature to
4
determine suitable parameters in PSD process. The PSED, heat integration and MVR heat pump distillations
were
studied
to
separate
the
complex
ternary
mixture
of
methyl
acetate/methanol/water with the minimum total annual cost (TAC) and the least energy consumption in this paper. Furthermore, factors such as the choice of solvent, optimization of solvent ratio and determination of pressure were also investigated in this study. 2. Methods and evaluation 2.1 Property method. The selection of physical property model is the key to obtain reasonable simulation results. According to the ternary mixture’s vapor-liquid equilibrium data [26,27], the interaction between the two substances was calculated by using the “Data regression” function. The simulation results obtained from Aspen Plus demonstrated that NRTL model for the regressed interaction parameters with the vapor-liquid equilibrium of methyl acetate/methanol/water system agree well with those reported in literature. Additionally, the regressed interaction parameters for the SR-Polar model were calculated (see Table S1 in the Supporting Information). 2.2 Economic evaluation Total annual cost (TAC) [28], which consists of annual operating cost (OC) and capital investment (CI) is used to evaluate and compare various PSD and ED process designs. The OC contains steam cost, cooling water cost and the electricity cost of the compressor, while CI includes depreciation expense of the distillation column, plate, column internals, compressor and heat exchanger equipment cost. Many small equipment sets was usually ignored, such as valves, pipes, pumps and column skirts due to the small proportion of TAC calculation. The simulation was completed by optimizing each factor individually with a flowchart [29]. The equations used
5
for calculations are given below.
TAC OC CI /
(1)
OC
(2)
(3)
CW (7500 3600)(QC / 4.18 6) /1000
(4)
7500 CM WM
(5)
CI CC (1.35 H 2 / 4) C A AT CY
(6)
The diameter of the column is calculated by “pack sizing” function in Aspen Plus. The operating time of designs was set 7500 h/year. Among the formulae, the payback period is denoted by θ, whereas the limit of recycle has been set to 5. All other symbols used in equations have been explained in Table S2 of Supporting Information. 3. Processes and discussion The ternary mixture is fed at 3500 kg/h with the mass fraction of 45 wt.% methyl acetate, 40 wt.% methanol, and 15 wt.% water. The purity of three products was up to 99.8 wt.%, with the lowest recovery of solvent set to be 99.7 wt.%. The boiling points of methyl acetate, methanol and water are 330.2 K, 337.68 K and 373.15 K, respectively. 3.1 Composition analysis using triangular diagram Triangular diagram is used to describe the properties of residue curves, distillation boundary and azeotropes (methyl acetate/methanol and acetate/water azeotropic systems) at 101.325 kPa. The diagram illustrates the boiling point of pure components, whereas the minimum-boiling azeotropes under normal temperature and atmospheric pressure have compositions of 82 wt.% methyl acetate (in methyl acetate and methanol mixture, while rest being the methanol) with the 6
boiling point of 326.74 K, and 97 wt.% methyl acetate at 329.81 K (in methyl acetate and water mixture, while rest being the water). Distillation boundary divides the map into two parts (Part 1 and Part 2) as illustrated in Fig. 1. 3.2 Selection of operating pressure The principle of PSD is that the binary mixture deviates from the Raoult's law. Azeotropic composition can be broken or changed by adjusting the pressure, while at the same time, the high purity product can be obtained from the top and bottom of the column by distillation. The main purpose of pressure-selection is to destroy the methyl acetate/water azeotrope for methyl acetate recovery column. The pressure of the column was varied between 20 - 101.325 kPa. Fig. 2 illustrates the trend of the mass fraction change of methyl acetate with pressure. The results show that when the pressure decreased from 101.325 kPa to 20 kPa, the composition of azeotrope changed rapidly. Even the azeotrope disappeared as the pressure was varied, which saved a great quantity of energy-consumption by avoiding the addition of extraction agent. The data in Table 1 shows that methyl acetate/water azeotropic composition is pressure-sensitive. The pressure of the first column (P1) is determined to be 20 kPa. The pressure of the second column (P2) was also determined by the change in azeotropic composition. The pressure of this column was also varied between 20 - 101.325 kPa. Fig. 3 describes the trend of the mass fraction change of methyl acetate with pressure. When the lower-pressure column was operated at 20 kPa, the concentration of methyl acetate increased from 82 wt.% to 90 wt.%, which reached the requirement of PSD, as is shown in Table 2. Furthermore, the operating pressure had close relationship with the steam temperature. Considering the condensation cost of overhead vapor, the operating pressure cannot be too low. In
7
this way, a small amount of solvent can meet the requirement of separation. Additionally, the choice of pressure makes great contribution to the energy utilization in subsequent heat integration. Finally, the pressures of DCPSED were set to be 20 kPa and 54 kPa, while the pressures of TCPSED were set to be 20 kPa, 20 kPa and 75 kPa. 3.3 Optimization of solvent The choice of solvent is critical to the extraction operation, and is the first step in modeling the extraction process. In the separation process of methyl acetate/methanol, the extraction agent (also known as the solvent) is still needed to break the azeotropic composition to increase the relative volatility of the component. This way, the goal of separating the azeotropic system or the near boiling system was achieved. The solvent ratio (extraction ratio) was set as 1:1, while several candidate extraction agents, such as water, propylene glycol, N,N-dimethylformamide (DMF), ethylene glycol and glycerol were selected for comparison. It is noteworthy that for higher relative volatility of the components, the solvent works better, and vice versa [24]. After using the two-phase flash stream model (Flash2 model in Aspen Plus software), the calculation results in Table 3 show that water is the optimal choice. Recycling of solvent is also an evaluation index of the process. Being economical, accessible, harmless, thermally and chemically stable, and non-corrosive, the most important and primary attribute of water is the large selectivity among methyl acetate/methanol mixture, which determines the difficulty of separation due to its impact on the number of stages and reflux ratio (affect the diameter) [20]. Furthermore, water was detected in the raw waste water, which further increases the amount of solvent in circulation.
8
3.4 Simulation processes 3.4.1 DCPSED process The effects of solvent ratio (S/ABC, the mass ratio of solvent to feed) on the purity of methyl acetate and the heat load of column bottoms were investigated to determine the optimum solvent ratio. Considering the first column (C1) as an example, the results in Fig. 4 demonstrated that the purity of methyl acetate changed rapidly with the increase of solvent ratio. Once the solvent ratio reached 0.2, the trend tended to be gentle and basically remained unchanged. Eventually, the solvent ratio was set to be 0.2. Fig. 5 shows the PSED process of the ternary azeotrope. The DCPSED process is mainly composed of two columns with different operating pressures and includes a lower-pressure column (C1, 20kPa) and a higher-pressure column (C2, 54kPa). Fresh feed (ABC) consists of the mixture of 45 wt.% methyl acetate, 40 wt.% methanol, and 15 wt.% water. The mixture of additional extraction agent (C) and the stream from solvent recovery column were added to break the methyl acetate/methanol azeotropic composition. Fig. 5 presents the specific process flowsheet of the process. Streams ABC and C flowed into C1 (20kPa) from different feed stages, while methyl acetate (99.8 wt.%) was obtained as product from the top of C1. Stream BC was fed to the next column at a pressure of 54 kPa. The separation of methanol and water is based on the difference in their boiling temperatures. High purity of methanol can be achieved, whereas the water can be recycled in the bottom. The mass balance and heat balance of DCPSED is presented in Table S3 (see Supporting Information). 3.4.2 TCPSED process Fig. 6 shows the TCPSED process, which is mainly composed of three different-pressure
9
columns, namely the low-pressure column C1 (20kPa), the low-pressure column C2 (20kPa), the high-pressure column C3 (75kPa). Stream ABC flowed into C1 and a mixture near the azeotropic composition was obtained from the top of C1. The azeotropic composition and water (solvent) flowed together into C2, whereas methyl acetate was separated from the top of C2. Streams BC from the bottom of C1 and C2 were transported to C3. High purity methanol was obtained from the top of C3. The water separated from the bottom of C3 can be recycled as the solvent was added to C2. A comparison of results for DCPSED and TCPSED processes is presented in Table 4. All the parameters were optimized. Obviously, the DCPSED process is superior to TCPSED process, no matter what the energy consumption or TAC is. The mass and heat balance of TCPSED is shown in Table S4 (see Supporting Information). 3.4.3 PSED with heat integration The high temperature and quality heat source provide the needed energy to column bottom, while the external cold source removes the vapor’s latent heat at the top of column in a conventional distillation process, thus irreversibly bringing about the loss of energy and reduction in thermodynamic efficiency. Due to this, the efficient use of energy and improvement in thermodynamic efficiency have become a hot research topic. Based on the conventional process, PSED with heat integration was proposed to achieve the energy-saving effect after meeting the temperature difference in heat exchanger, which means using the vaporization heat from the high-pressure column to heat the low-pressure column’s bottom [17]. In DCPSED with heat integration process, the heat of the top of C2 served as the heat source for the bottom of C1 to indirectly reach energy-saving. The flowsheet is described in Fig. 7.
10
In TCPSED with heat integration process, the heat of the top of C3 is matched to that of the bottom of the C2. The detailed process is shown in Fig. 8. On the basis of heat-matching, the hot stream was fully used to heat the cold stream to reduce the heating and cooling duties, thus improving the system's heat recovery capacity and reducing the investment costs. Unfortunately, even under the most ideal conditions, the latent heat of column’s top vapors cannot be used and the bottom of the column must be heated using external heat. Therefore, the best energy-saving effect cannot be reached. Table 5 presents all the optimized simulation parameters. 3.4.4 PSED with MVR As discussed above, the temperature approach between the bottom and top of the column caused
a
lot
of
energy
wastage,
which
urged
the
authors
to
explore
new
energy conservation techniques. MVR heat pump technology possesses energy-efficient characteristics due to the utilization of top vapor’s latent heat [30,31]. On one hand, the application of MVR has to meet a certain temperature difference for the transfer of heat, while on the other, the heat match between the condensation and reboiler must also be accomplished. The compression ratio of compressor is determined by the heat transfer temperature difference, which affects heat transfer area in the bottom of column. C1 in DCPSED process was used as an example to analyze the results. The compression ratio of compressor increases with the enlargement of heat transfer temperature difference, as shown in Fig. 9. Considering the industry’s requirement and thermodynamic efficiency, the heat transfer temperature difference is determined to be 10 °C. Finally, the compression ratio was defined to be 3.6, when the required temperature difference is achieved.
11
After designing and contrasting the energy-saving effect, the self-sufficient MVR with single compression was confirmed as the optimum. In other words, the selection of a proper compression ratio helps in increasing the temperature and pressure of top vapors in individual columns and heats the bottom stream. Fig. 10 and 11 explain the specific process. Table 6 presents the optimized parameters as obtained from the simulation. The simulation results imply that the overhead heat is smaller than the reboiler heat in all columns, while the auxiliary reboilers are needed to make up for the extra energy. The data in Table 6 shows that the DCPSED with MVR is better than TCPSED with MVR process with 16.39% reduction in energy consumption and 40.70% decrease in TAC. 4. Conclusions The simulated results manifested that water is the best solvent for separation of ternary mixture containing methyl acetate, methanol and water which originated from chemical waste. The feasibility of PSED was analyzed at different pressures by Aspen plus and the purity of methyl acetate/methanol/water in distillation reached 99.8 wt.%. The low-pressure column and solvent are beneficial to increase the relative volatility of methyl acetate-methanol mixture and reduce the energy consumption. The results demonstrated that DCPSED is superior to TCPSED process with reducing reduction in energy consumption and TAC of 44.83% and 47.74%, respectively. DCPSED and TCPSED processes combined with the heat integration energy-saving process reduced the energy consumption by 40.72% and 17.36%, and decreased the TAC by 37.72% and 22.09% respectively, thereby attained a remarkable energy-saving effect. Surprisingly, once the MVR technology was applied to DCPSED and TCPSED processes, a huge change was brought about in the processes’ outcomes. Compared to DCPSED and TCPSED processes, the energy
12
consumption was reduced by about 76.00% and 77.96%, and TAC was decreased by about 45.05% and 51.41%, respectively. Through contrastive analysis, the DCPSED with MVR heat-pump is confirmed as the most economical process. Furthermore, among the four processes, DCPSED coupled with MVR heat-pump is the optimum process to separate methyl acetate/methanol/water mixture, and provides the best energy-saving advantage. These studies reveal that DCPSED with MVR process has significant potential as an economic and feasible technological route for separating methyl acetate/methanol/water azeotropic system.
ASSOCIATED CONTENT Supporting Information Additional information as noted in text. This material is available free charge via the Internet at http://pubs.acs.org.
ACKOWLEDGEMENT This research was funded by Advanced Catalysis and Green Manufacturing Collaborative Innovation Center，Changzhou University，213164.
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Figure captions Fig. 1. Ternary map for methyl acetate/methanol/water system Fig. 2. Effect of pressure on methyl acetate/water system’s vapor-liquid equilibrium Fig. 3. Effect of pressure on methyl acetate/methanol system’s vapor-liquid equilibrium Fig. 4. Effect of solvent ratio on methyl acetate’s purity and heat duty Fig. 5. Process flow diagram of DCPSED process Fig. 6. Process flow diagram of TCPSED process Fig. 7. Process flow diagram of DCPSED with heat integration process Fig. 8. Process flow diagram of TCPSED with heat integration process Fig. 9. Effect of compression ratio on temperature difference Fig. 10. Process flow diagram of DCPSED process with MVR Fig. 11. Process flow diagram of TCPSED process with MVR Table caption Table 1. Effect of pressure on methyl acetate-water azeotrope’s composition Table 2. Effect of pressure on methyl acetate/methanol azeotrope’s composition Table 3.Screening results for the extraction agent Table 4.Simulation results for conventional PSED process Table 5. Simulation results for PSED with heat integration 17
Table 6. Simulation results for PSED process with MVR
Fig. 1. Ternary map for methyl acetate/methanol/water system
Vapor massfraction of methyl acetate
1.0
0.8
a
b
c
d
0.6
0.4
a b c d
0.2
0.0 0.0
0.2
0.4
101.325 kPa 50 kPa 40 kPa 20 kPa
0.6
Liquid massfraction of methyl acetate
18
0.8
1.0
Fig. 2. Effect of pressure on methyl acetate/water system’s vapor-liquid equilibrium
1.0
d
Vapor massfraction of methyl acetate
c b
0.8
a 0.6
0.4
a b c d
0.2
0.0 0.0
0.2
0.4
101.325 kPa 50 kPa 30 kPa 20 kPa
0.6
0.8
1.0
Liquid massfraction of methyl acetate
Fig. 3. Effect of pressure on methyl acetate/methanol system’s vapor-liquid equilibrium
0.9985
1020
Mass fraction of methyl acetate Heat duty 1010
0.9975 1000 0.9970
990 0.9965
0.9960 0.0
0.1
0.2
0.3
0.4
980 0.5
Solvent ratio
Fig. 4. Effect of solvent ratio on methyl acetate’s purity and heat duty
19
Heat duty (kW)
Mass fraction of methyl acetate
0.9980
Fig. 5. Process flow diagram of DCPSED process
Fig. 6. Process flow diagram of TCPSED process
Fig. 7. Process flow diagram of DCPSED with heat integration process
Fig. 8. Process flow diagram of TCPSED with heat integration process
20
30
temperature difference
Temperature difference (℃)
20
10
0
-10
-20 1
2
3
4
5
6
Compression ratio
Fig. 9. Effect of compression ratio on temperature difference
Fig. 10. Process flow diagram of DCPSED process with MVR
Fig. 11. Process flow diagram of TCPSED process with MVR
21
Table 1. Effect of pressure on methyl acetate-water azeotrope’s composition Parameter
101.325 kPa
50 kPa
40 kPa
20 kPa
Mass fraction of methyl acetate
0.97
0.99
0.996
--
Mass fraction of methanol
0.03
0.01
0.004
--
Azeotropic temperature (°C)
56.66
38.28
32.83
--
Table 2. Effect of pressure on methyl acetate/methanol azeotrope’s composition Parameter
101.325 kPa
50 kPa
30 kPa
20 kPa
Mass fraction of methyl acetate
0.82
0.86
0.88
0.90
Mass fraction of methanol
0.18
0.14
0.12
0.10
Azeotropic temperature (°C)
53.59
36.18
24.58
16.05
Table 3. Screening results for the extraction agent Extractant
Water
Boiling point (°C)
100
Relative volatility
11.1
Propylene glycol
DMF
124.4 152 Methyl acetate - Methanol system 1.0 1.4
22
Ethylene glycol
Propanetriol
197.3
287.9
1.4
1.2
Table 4. Simulation results for conventional PSED process DCPSED
TCPSED
Parameter C1
C2
C1
C2
C3
Number of stages
75
25
40
43
24
Operating pressure (kPa)
20
54
20
20
75
Reflux ratio
8
2
6
6
1.35
Feed stage
70
19
35
40
18
Extractant feed stage
25
--
--
19
--
Solvent ratio
0.20
--
--
0.29
--
Diameter (m)
1.27
0.97
1.74
1.43
0.81
Top temperature (°C)
17.18
49.23
16.06
17.17
56.94
Bottom temperature (°C)
38.10
83.00
34.93
47.26
91.43
Condenser duty (kW)
973.05
1360.53
1852.04
1362.26
1053.25
Reboiler duty (kW)
991.27
1443.25
1852.78
1392.68
1167.41
Annual coal consumption 2242.83
4065.41
423.12
1166.59
Annual steam cost ($)
1.00×106
1.81×106
Annual water cost ($)
2.03×105
5.07×105
Annual equipment cost ($)
2.55×104
4.65×104
TAC ($)
1.23×106
2.35×106
(×103 kg) Total heat exchange area (m2)
23
Table 5. Simulation results for PSED with heat integration DCPSED with heat integration
TCPSED with heat integration
Parameter C1
C2
C1
C2
C3
Number of stages
75
25
40
43
24
Operating pressure (kPa)
20
54
20
20
75
Reflux ratio
8
2
6
6
1.35
Feed stage
70
19
35
40
18
Extractant feed stage
25
--
--
19
--
Solvent ratio
0.20
--
--
0.29
--
Diameter (m)
1.27
0.97
1.74
1.43
0.81
Top temperature (°C)
17.18
49.23
16.06
17.17
56.94
Bottom temperature (°C)
38.10
83.00
34.93
47.26
91.43
Condenser duty (kW)
991.27
361.30
1852.04
1362.26
--
--
1443.25
1852.78
339.43
1167.41
Reboiler duty (kW) Annual coal consumption (×103 kg)
1329.61
3359.62
Total heat exchange area (m2)
499.16
996.34
Annual steam cost ($)
5.94×105
1.41×106
Annual water cost ($)
1.46×105
4.17×105
Annual equipment cost ($)
2.55×104
3.03×104
TAC ($)
7.66×105
1.83×106
24
Table 6. Simulation results for PSED process with MVR DCPSED with MVR
TCPSED with MVR
Parameter C1
C2
C1
C2
C3
Number of stages
75
25
40
43
24
Operating pressure (kPa)
20
54
20
20
75
Reflux ratio
8
2
6
6
1.35
3.75
5.19
3.6
5
4.8
163.15
298.88
290.53
281.48
225.51
Feed stage
70
19
35
40
18
Extractant feed stage
25
--
--
19
--
Solvent ratio
0.20
--
--
0.29
--
Diameter (m)
1.27
0.97
1.74
1.43
0.81
Top temperature (°C)
17.18
49.23
16.06
17.17
56.94
Bottom temperature (°C)
38.10
83.00
34.93
47.26
91.43
--
--
--
--
--
26.18
95.32
0.74
30.42
114.16
Compression ratio Compressor power (kW)
Condenser duty (kW) Reboiler duty (kW) Annual coal consumption (×103 kg)
538.26
895.84
253.43
456.09
Annual steam cost ($)
4.95×104
5.93×104
Annual water cost ($)
5.71×105
9.83×105
Total heat exchange area (m2)
25
Annual equipment cost ($)
--
--
Annual equipment cost ($)
5.62×104
1.01×105
TAC ($)
6.76×105
1.14×106
26
S0255-2701(17)30226-X https://doi.org/10.1016/j.cep.2017.10.012 CEP 7096
To appear in:
Chemical Engineering and Processing
Received date: Revised date: Accepted date:
10-3-2017 23-6-2017 14-10-2017
Please cite this article as: Xiaoxin Gao, Biyun Zhu, Jiangquan Ma, Deming Yang, A combination of pressure-swing and extractive distillation for separating complex binary azeotropic system, Chemical Engineering and Processing https://doi.org/10.1016/j.cep.2017.10.012 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
A combination of pressure-swing and extractive distillation for separating complex binary azeotropic system Xiaoxin Gao,a,b,* Biyun Zhu,b Jiangquan Ma,a,b Deming Yanga,b a. Jiangsu Key Laboratory of Advanced Catalytic Materials and Technology, Changzhou University, Changzhou 213164, P.R. China b. College of Petrochemical Engineering, Changzhou University, Changzhou 213164, P.R. China
AUTHOR INFORMATION Corresponding Author E-mail:[email protected]. Tel: +86-519-86330255
Highlights
The double column pressure-swing with extractive distillation(DCPSED) was proposed.
The triple column pressure-swing with extractive distillation were proposed.
The heat integration and MVR heat pump distillations were also tested.
The DCPSED with MVR process is used to separate complex binary azeotropic system.
ABSTRACT: Methyl acetate/methanol/water mixture forms more than one different azeotrope, whereas its triangular diagram presents a distillation boundary at atmospheric pressure. The two different simulation processes of the double column pressure-swing with extractive distillation (DCPSED) and triple column pressure-swing with extractive distillation (TCPSED) were proposed to separate the complex ternary system. Furthermore, the heat integration and 1
mechanical vapor compression (MVR) heat pump distillations were also tested (through simulation) to separate the complex ternary system to save energy. The feasibility of the processes was confirmed using vapor-liquid equilibrium curve, while rigorous steady-state simulations were implemented using Aspen Plus (version 7.1). With the goal of achieving minimum energy consumption and minimum total annual cost (TAC), the purities of methyl acetate, methanol and water were used as the constraint variables. The results show that water was a better solvent for this separation system. Compared with the TCPSED process, the energy consumption and TAC for DCPSED process could be reduced by 44.83%, and 47.74%, respectively. Compared with the DCPSED and DCPSED with heat integration processes, DCPSED with MVR process could reduce the energy consumption by 73.34% and 59.52%, while TAC was decreased by 37.64% and 11.77%, respectively. The simulation results indicated that DCPSED with MVR process has great advantage over other processes used to separate methyl acetate/methanol/water ternary mixture. Keywords: Azeotropic compositions; Pressure-swing distillation; Extractive distillation; Heat integration; MVR technology 1. Introduction Methanol (CH3OH) and methyl acetate (C3H6O2) are extensively applied as effective organic solvents in chemical and pharmaceutical industries owing to their excellent physicochemical properties [1]. Due to great application prospects in industry, it is attractive and necessary to separate and reuse methanol and methyl acetate to comply all processing plants with modern and rigorous environmental and resource-saving regulations. However, high-purity separation of ternary mixture of methyl acetate/methanol and methyl acetate/water azeotropes could not be
2
efficiently accomplished in common practical operation. Due to this, the lower purity of methyl acetate/methanol is likely to be at the cost of higher reflux and larger number of theoretical stages by conventional extractive distillation process [2]. Distillation is the most widely used separation process in chemical industries, whereas pressure-swing distillation (PSD) and extractive distillation (ED) are the most common methods for separation of a binary homogeneous azeotrope[3]. The ternary system containing two azeotropes has attracted more and more research attention in recent years [4-9]. In comparison to some alternative methods for separation of ternary azeotropic systems, such as azeotropic distillation [10], reactive distillation [11-12], liquid-liquid extractive distillation [13], salt addition [14], membrane distillation [15] and some other new coupling separation techniques, PSD and ED are broadly applied in recently years [16-19]. Although the membrane separation technology has the advantage of energy-saving in the process of breaking the azeotropic composition, the investment for the equipment cost is very high and commercial methods need to be further developed in this regard [15]. In azeotropic distillation, a new component was introduced as an entrainment agent, which resulted in more energy-consumption to recycle the entrainer. Zhu et al. [5] studied the separation of acetonitrile/methanol/benzene mixture via triple column pressure-swing distillation and analyzed the residue-curve map to illustrate the relationship of pressure and distillation boundary. Kaam et al. [20] investigated the feasibility and conducted
experiments
about
the
heterogeneous
extractive
batch
distillation
of
chloroform/methanol/water system, especially focusing on the entrainer selection for the separation of chloroform-methanol mixture. Luyben [21] introduced a comparison of extractive distillation and pressure-swing distillation for acetone/chloroform separation, and reported that the
3
pressure-swing distillation is much more expensive than the extractive distillation for acetone/chloroform separation in terms of both capital investment and energy consumption without considering product contamination. Lladosa et al. [22] proposed that the pressure-swing distillation process is effective to some extent, provided the azeotropic composition changes significantly with pressures. For ED, the key is the selection of solvent, and therefore, the ED process is effective if a suitable solvent can be found. ED requires an appropriate solvent to improve the relative volatility of the separated components, so that the mixture can be separated more easily and effectively. Pressure-swing distillation (PSD) uses pressure difference between several columns and changes the composition of azeotrope to effectively separate the mixture, and therefore, avoids inducing a third component to achieve energy saving. Gomis et al. [23] successfully set gasoline fraction mixtures as entrainers to implement ethanol dehydration, and conducted a pilot-scale industrial test to produce bioethanol and naphta. Deorukhkar et al. [24] used dimethyl sulfoxide (DMSO) as the entrainer to purify tetrahydrofuran (THF) from its aqueous azeotrope through continuous distillation, which recycled THF with a purity greater than 99 wt.%. Wang et al. [25] studied the design of pressure-swing distillation for the mixture of n-heptane and isobutanol with steady and dynamic simulations in the commercial software package Aspen Plus (version 7.2). By comparing the conventional PSD (CPSD) with unusual PSD (UPSD), the authors [23] proposed that the two processes should be applied according to their advantages. For a complex and special azeotropic system, conventional distillation could not achieve a desired separation, due to which, a combination of pressure-swing distillation and extractive distillation was studied in this paper. The change of pressure has great impact on the azeotropic composition and azeotropic temperature to
4
determine suitable parameters in PSD process. The PSED, heat integration and MVR heat pump distillations
were
studied
to
separate
the
complex
ternary
mixture
of
methyl
acetate/methanol/water with the minimum total annual cost (TAC) and the least energy consumption in this paper. Furthermore, factors such as the choice of solvent, optimization of solvent ratio and determination of pressure were also investigated in this study. 2. Methods and evaluation 2.1 Property method. The selection of physical property model is the key to obtain reasonable simulation results. According to the ternary mixture’s vapor-liquid equilibrium data [26,27], the interaction between the two substances was calculated by using the “Data regression” function. The simulation results obtained from Aspen Plus demonstrated that NRTL model for the regressed interaction parameters with the vapor-liquid equilibrium of methyl acetate/methanol/water system agree well with those reported in literature. Additionally, the regressed interaction parameters for the SR-Polar model were calculated (see Table S1 in the Supporting Information). 2.2 Economic evaluation Total annual cost (TAC) [28], which consists of annual operating cost (OC) and capital investment (CI) is used to evaluate and compare various PSD and ED process designs. The OC contains steam cost, cooling water cost and the electricity cost of the compressor, while CI includes depreciation expense of the distillation column, plate, column internals, compressor and heat exchanger equipment cost. Many small equipment sets was usually ignored, such as valves, pipes, pumps and column skirts due to the small proportion of TAC calculation. The simulation was completed by optimizing each factor individually with a flowchart [29]. The equations used
5
for calculations are given below.
TAC OC CI /
(1)
OC
(2)
(3)
CW (7500 3600)(QC / 4.18 6) /1000
(4)
7500 CM WM
(5)
CI CC (1.35 H 2 / 4) C A AT CY
(6)
The diameter of the column is calculated by “pack sizing” function in Aspen Plus. The operating time of designs was set 7500 h/year. Among the formulae, the payback period is denoted by θ, whereas the limit of recycle has been set to 5. All other symbols used in equations have been explained in Table S2 of Supporting Information. 3. Processes and discussion The ternary mixture is fed at 3500 kg/h with the mass fraction of 45 wt.% methyl acetate, 40 wt.% methanol, and 15 wt.% water. The purity of three products was up to 99.8 wt.%, with the lowest recovery of solvent set to be 99.7 wt.%. The boiling points of methyl acetate, methanol and water are 330.2 K, 337.68 K and 373.15 K, respectively. 3.1 Composition analysis using triangular diagram Triangular diagram is used to describe the properties of residue curves, distillation boundary and azeotropes (methyl acetate/methanol and acetate/water azeotropic systems) at 101.325 kPa. The diagram illustrates the boiling point of pure components, whereas the minimum-boiling azeotropes under normal temperature and atmospheric pressure have compositions of 82 wt.% methyl acetate (in methyl acetate and methanol mixture, while rest being the methanol) with the 6
boiling point of 326.74 K, and 97 wt.% methyl acetate at 329.81 K (in methyl acetate and water mixture, while rest being the water). Distillation boundary divides the map into two parts (Part 1 and Part 2) as illustrated in Fig. 1. 3.2 Selection of operating pressure The principle of PSD is that the binary mixture deviates from the Raoult's law. Azeotropic composition can be broken or changed by adjusting the pressure, while at the same time, the high purity product can be obtained from the top and bottom of the column by distillation. The main purpose of pressure-selection is to destroy the methyl acetate/water azeotrope for methyl acetate recovery column. The pressure of the column was varied between 20 - 101.325 kPa. Fig. 2 illustrates the trend of the mass fraction change of methyl acetate with pressure. The results show that when the pressure decreased from 101.325 kPa to 20 kPa, the composition of azeotrope changed rapidly. Even the azeotrope disappeared as the pressure was varied, which saved a great quantity of energy-consumption by avoiding the addition of extraction agent. The data in Table 1 shows that methyl acetate/water azeotropic composition is pressure-sensitive. The pressure of the first column (P1) is determined to be 20 kPa. The pressure of the second column (P2) was also determined by the change in azeotropic composition. The pressure of this column was also varied between 20 - 101.325 kPa. Fig. 3 describes the trend of the mass fraction change of methyl acetate with pressure. When the lower-pressure column was operated at 20 kPa, the concentration of methyl acetate increased from 82 wt.% to 90 wt.%, which reached the requirement of PSD, as is shown in Table 2. Furthermore, the operating pressure had close relationship with the steam temperature. Considering the condensation cost of overhead vapor, the operating pressure cannot be too low. In
7
this way, a small amount of solvent can meet the requirement of separation. Additionally, the choice of pressure makes great contribution to the energy utilization in subsequent heat integration. Finally, the pressures of DCPSED were set to be 20 kPa and 54 kPa, while the pressures of TCPSED were set to be 20 kPa, 20 kPa and 75 kPa. 3.3 Optimization of solvent The choice of solvent is critical to the extraction operation, and is the first step in modeling the extraction process. In the separation process of methyl acetate/methanol, the extraction agent (also known as the solvent) is still needed to break the azeotropic composition to increase the relative volatility of the component. This way, the goal of separating the azeotropic system or the near boiling system was achieved. The solvent ratio (extraction ratio) was set as 1:1, while several candidate extraction agents, such as water, propylene glycol, N,N-dimethylformamide (DMF), ethylene glycol and glycerol were selected for comparison. It is noteworthy that for higher relative volatility of the components, the solvent works better, and vice versa [24]. After using the two-phase flash stream model (Flash2 model in Aspen Plus software), the calculation results in Table 3 show that water is the optimal choice. Recycling of solvent is also an evaluation index of the process. Being economical, accessible, harmless, thermally and chemically stable, and non-corrosive, the most important and primary attribute of water is the large selectivity among methyl acetate/methanol mixture, which determines the difficulty of separation due to its impact on the number of stages and reflux ratio (affect the diameter) [20]. Furthermore, water was detected in the raw waste water, which further increases the amount of solvent in circulation.
8
3.4 Simulation processes 3.4.1 DCPSED process The effects of solvent ratio (S/ABC, the mass ratio of solvent to feed) on the purity of methyl acetate and the heat load of column bottoms were investigated to determine the optimum solvent ratio. Considering the first column (C1) as an example, the results in Fig. 4 demonstrated that the purity of methyl acetate changed rapidly with the increase of solvent ratio. Once the solvent ratio reached 0.2, the trend tended to be gentle and basically remained unchanged. Eventually, the solvent ratio was set to be 0.2. Fig. 5 shows the PSED process of the ternary azeotrope. The DCPSED process is mainly composed of two columns with different operating pressures and includes a lower-pressure column (C1, 20kPa) and a higher-pressure column (C2, 54kPa). Fresh feed (ABC) consists of the mixture of 45 wt.% methyl acetate, 40 wt.% methanol, and 15 wt.% water. The mixture of additional extraction agent (C) and the stream from solvent recovery column were added to break the methyl acetate/methanol azeotropic composition. Fig. 5 presents the specific process flowsheet of the process. Streams ABC and C flowed into C1 (20kPa) from different feed stages, while methyl acetate (99.8 wt.%) was obtained as product from the top of C1. Stream BC was fed to the next column at a pressure of 54 kPa. The separation of methanol and water is based on the difference in their boiling temperatures. High purity of methanol can be achieved, whereas the water can be recycled in the bottom. The mass balance and heat balance of DCPSED is presented in Table S3 (see Supporting Information). 3.4.2 TCPSED process Fig. 6 shows the TCPSED process, which is mainly composed of three different-pressure
9
columns, namely the low-pressure column C1 (20kPa), the low-pressure column C2 (20kPa), the high-pressure column C3 (75kPa). Stream ABC flowed into C1 and a mixture near the azeotropic composition was obtained from the top of C1. The azeotropic composition and water (solvent) flowed together into C2, whereas methyl acetate was separated from the top of C2. Streams BC from the bottom of C1 and C2 were transported to C3. High purity methanol was obtained from the top of C3. The water separated from the bottom of C3 can be recycled as the solvent was added to C2. A comparison of results for DCPSED and TCPSED processes is presented in Table 4. All the parameters were optimized. Obviously, the DCPSED process is superior to TCPSED process, no matter what the energy consumption or TAC is. The mass and heat balance of TCPSED is shown in Table S4 (see Supporting Information). 3.4.3 PSED with heat integration The high temperature and quality heat source provide the needed energy to column bottom, while the external cold source removes the vapor’s latent heat at the top of column in a conventional distillation process, thus irreversibly bringing about the loss of energy and reduction in thermodynamic efficiency. Due to this, the efficient use of energy and improvement in thermodynamic efficiency have become a hot research topic. Based on the conventional process, PSED with heat integration was proposed to achieve the energy-saving effect after meeting the temperature difference in heat exchanger, which means using the vaporization heat from the high-pressure column to heat the low-pressure column’s bottom [17]. In DCPSED with heat integration process, the heat of the top of C2 served as the heat source for the bottom of C1 to indirectly reach energy-saving. The flowsheet is described in Fig. 7.
10
In TCPSED with heat integration process, the heat of the top of C3 is matched to that of the bottom of the C2. The detailed process is shown in Fig. 8. On the basis of heat-matching, the hot stream was fully used to heat the cold stream to reduce the heating and cooling duties, thus improving the system's heat recovery capacity and reducing the investment costs. Unfortunately, even under the most ideal conditions, the latent heat of column’s top vapors cannot be used and the bottom of the column must be heated using external heat. Therefore, the best energy-saving effect cannot be reached. Table 5 presents all the optimized simulation parameters. 3.4.4 PSED with MVR As discussed above, the temperature approach between the bottom and top of the column caused
a
lot
of
energy
wastage,
which
urged
the
authors
to
explore
new
energy conservation techniques. MVR heat pump technology possesses energy-efficient characteristics due to the utilization of top vapor’s latent heat [30,31]. On one hand, the application of MVR has to meet a certain temperature difference for the transfer of heat, while on the other, the heat match between the condensation and reboiler must also be accomplished. The compression ratio of compressor is determined by the heat transfer temperature difference, which affects heat transfer area in the bottom of column. C1 in DCPSED process was used as an example to analyze the results. The compression ratio of compressor increases with the enlargement of heat transfer temperature difference, as shown in Fig. 9. Considering the industry’s requirement and thermodynamic efficiency, the heat transfer temperature difference is determined to be 10 °C. Finally, the compression ratio was defined to be 3.6, when the required temperature difference is achieved.
11
After designing and contrasting the energy-saving effect, the self-sufficient MVR with single compression was confirmed as the optimum. In other words, the selection of a proper compression ratio helps in increasing the temperature and pressure of top vapors in individual columns and heats the bottom stream. Fig. 10 and 11 explain the specific process. Table 6 presents the optimized parameters as obtained from the simulation. The simulation results imply that the overhead heat is smaller than the reboiler heat in all columns, while the auxiliary reboilers are needed to make up for the extra energy. The data in Table 6 shows that the DCPSED with MVR is better than TCPSED with MVR process with 16.39% reduction in energy consumption and 40.70% decrease in TAC. 4. Conclusions The simulated results manifested that water is the best solvent for separation of ternary mixture containing methyl acetate, methanol and water which originated from chemical waste. The feasibility of PSED was analyzed at different pressures by Aspen plus and the purity of methyl acetate/methanol/water in distillation reached 99.8 wt.%. The low-pressure column and solvent are beneficial to increase the relative volatility of methyl acetate-methanol mixture and reduce the energy consumption. The results demonstrated that DCPSED is superior to TCPSED process with reducing reduction in energy consumption and TAC of 44.83% and 47.74%, respectively. DCPSED and TCPSED processes combined with the heat integration energy-saving process reduced the energy consumption by 40.72% and 17.36%, and decreased the TAC by 37.72% and 22.09% respectively, thereby attained a remarkable energy-saving effect. Surprisingly, once the MVR technology was applied to DCPSED and TCPSED processes, a huge change was brought about in the processes’ outcomes. Compared to DCPSED and TCPSED processes, the energy
12
consumption was reduced by about 76.00% and 77.96%, and TAC was decreased by about 45.05% and 51.41%, respectively. Through contrastive analysis, the DCPSED with MVR heat-pump is confirmed as the most economical process. Furthermore, among the four processes, DCPSED coupled with MVR heat-pump is the optimum process to separate methyl acetate/methanol/water mixture, and provides the best energy-saving advantage. These studies reveal that DCPSED with MVR process has significant potential as an economic and feasible technological route for separating methyl acetate/methanol/water azeotropic system.
ASSOCIATED CONTENT Supporting Information Additional information as noted in text. This material is available free charge via the Internet at http://pubs.acs.org.
ACKOWLEDGEMENT This research was funded by Advanced Catalysis and Green Manufacturing Collaborative Innovation Center，Changzhou University，213164.
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Figure captions Fig. 1. Ternary map for methyl acetate/methanol/water system Fig. 2. Effect of pressure on methyl acetate/water system’s vapor-liquid equilibrium Fig. 3. Effect of pressure on methyl acetate/methanol system’s vapor-liquid equilibrium Fig. 4. Effect of solvent ratio on methyl acetate’s purity and heat duty Fig. 5. Process flow diagram of DCPSED process Fig. 6. Process flow diagram of TCPSED process Fig. 7. Process flow diagram of DCPSED with heat integration process Fig. 8. Process flow diagram of TCPSED with heat integration process Fig. 9. Effect of compression ratio on temperature difference Fig. 10. Process flow diagram of DCPSED process with MVR Fig. 11. Process flow diagram of TCPSED process with MVR Table caption Table 1. Effect of pressure on methyl acetate-water azeotrope’s composition Table 2. Effect of pressure on methyl acetate/methanol azeotrope’s composition Table 3.Screening results for the extraction agent Table 4.Simulation results for conventional PSED process Table 5. Simulation results for PSED with heat integration 17
Table 6. Simulation results for PSED process with MVR
Fig. 1. Ternary map for methyl acetate/methanol/water system
Vapor massfraction of methyl acetate
1.0
0.8
a
b
c
d
0.6
0.4
a b c d
0.2
0.0 0.0
0.2
0.4
101.325 kPa 50 kPa 40 kPa 20 kPa
0.6
Liquid massfraction of methyl acetate
18
0.8
1.0
Fig. 2. Effect of pressure on methyl acetate/water system’s vapor-liquid equilibrium
1.0
d
Vapor massfraction of methyl acetate
c b
0.8
a 0.6
0.4
a b c d
0.2
0.0 0.0
0.2
0.4
101.325 kPa 50 kPa 30 kPa 20 kPa
0.6
0.8
1.0
Liquid massfraction of methyl acetate
Fig. 3. Effect of pressure on methyl acetate/methanol system’s vapor-liquid equilibrium
0.9985
1020
Mass fraction of methyl acetate Heat duty 1010
0.9975 1000 0.9970
990 0.9965
0.9960 0.0
0.1
0.2
0.3
0.4
980 0.5
Solvent ratio
Fig. 4. Effect of solvent ratio on methyl acetate’s purity and heat duty
19
Heat duty (kW)
Mass fraction of methyl acetate
0.9980
Fig. 5. Process flow diagram of DCPSED process
Fig. 6. Process flow diagram of TCPSED process
Fig. 7. Process flow diagram of DCPSED with heat integration process
Fig. 8. Process flow diagram of TCPSED with heat integration process
20
30
temperature difference
Temperature difference (℃)
20
10
0
-10
-20 1
2
3
4
5
6
Compression ratio
Fig. 9. Effect of compression ratio on temperature difference
Fig. 10. Process flow diagram of DCPSED process with MVR
Fig. 11. Process flow diagram of TCPSED process with MVR
21
Table 1. Effect of pressure on methyl acetate-water azeotrope’s composition Parameter
101.325 kPa
50 kPa
40 kPa
20 kPa
Mass fraction of methyl acetate
0.97
0.99
0.996
--
Mass fraction of methanol
0.03
0.01
0.004
--
Azeotropic temperature (°C)
56.66
38.28
32.83
--
Table 2. Effect of pressure on methyl acetate/methanol azeotrope’s composition Parameter
101.325 kPa
50 kPa
30 kPa
20 kPa
Mass fraction of methyl acetate
0.82
0.86
0.88
0.90
Mass fraction of methanol
0.18
0.14
0.12
0.10
Azeotropic temperature (°C)
53.59
36.18
24.58
16.05
Table 3. Screening results for the extraction agent Extractant
Water
Boiling point (°C)
100
Relative volatility
11.1
Propylene glycol
DMF
124.4 152 Methyl acetate - Methanol system 1.0 1.4
22
Ethylene glycol
Propanetriol
197.3
287.9
1.4
1.2
Table 4. Simulation results for conventional PSED process DCPSED
TCPSED
Parameter C1
C2
C1
C2
C3
Number of stages
75
25
40
43
24
Operating pressure (kPa)
20
54
20
20
75
Reflux ratio
8
2
6
6
1.35
Feed stage
70
19
35
40
18
Extractant feed stage
25
--
--
19
--
Solvent ratio
0.20
--
--
0.29
--
Diameter (m)
1.27
0.97
1.74
1.43
0.81
Top temperature (°C)
17.18
49.23
16.06
17.17
56.94
Bottom temperature (°C)
38.10
83.00
34.93
47.26
91.43
Condenser duty (kW)
973.05
1360.53
1852.04
1362.26
1053.25
Reboiler duty (kW)
991.27
1443.25
1852.78
1392.68
1167.41
Annual coal consumption 2242.83
4065.41
423.12
1166.59
Annual steam cost ($)
1.00×106
1.81×106
Annual water cost ($)
2.03×105
5.07×105
Annual equipment cost ($)
2.55×104
4.65×104
TAC ($)
1.23×106
2.35×106
(×103 kg) Total heat exchange area (m2)
23
Table 5. Simulation results for PSED with heat integration DCPSED with heat integration
TCPSED with heat integration
Parameter C1
C2
C1
C2
C3
Number of stages
75
25
40
43
24
Operating pressure (kPa)
20
54
20
20
75
Reflux ratio
8
2
6
6
1.35
Feed stage
70
19
35
40
18
Extractant feed stage
25
--
--
19
--
Solvent ratio
0.20
--
--
0.29
--
Diameter (m)
1.27
0.97
1.74
1.43
0.81
Top temperature (°C)
17.18
49.23
16.06
17.17
56.94
Bottom temperature (°C)
38.10
83.00
34.93
47.26
91.43
Condenser duty (kW)
991.27
361.30
1852.04
1362.26
--
--
1443.25
1852.78
339.43
1167.41
Reboiler duty (kW) Annual coal consumption (×103 kg)
1329.61
3359.62
Total heat exchange area (m2)
499.16
996.34
Annual steam cost ($)
5.94×105
1.41×106
Annual water cost ($)
1.46×105
4.17×105
Annual equipment cost ($)
2.55×104
3.03×104
TAC ($)
7.66×105
1.83×106
24
Table 6. Simulation results for PSED process with MVR DCPSED with MVR
TCPSED with MVR
Parameter C1
C2
C1
C2
C3
Number of stages
75
25
40
43
24
Operating pressure (kPa)
20
54
20
20
75
Reflux ratio
8
2
6
6
1.35
3.75
5.19
3.6
5
4.8
163.15
298.88
290.53
281.48
225.51
Feed stage
70
19
35
40
18
Extractant feed stage
25
--
--
19
--
Solvent ratio
0.20
--
--
0.29
--
Diameter (m)
1.27
0.97
1.74
1.43
0.81
Top temperature (°C)
17.18
49.23
16.06
17.17
56.94
Bottom temperature (°C)
38.10
83.00
34.93
47.26
91.43
--
--
--
--
--
26.18
95.32
0.74
30.42
114.16
Compression ratio Compressor power (kW)
Condenser duty (kW) Reboiler duty (kW) Annual coal consumption (×103 kg)
538.26
895.84
253.43
456.09
Annual steam cost ($)
4.95×104
5.93×104
Annual water cost ($)
5.71×105
9.83×105
Total heat exchange area (m2)
25
Annual equipment cost ($)
--
--
Annual equipment cost ($)
5.62×104
1.01×105
TAC ($)
6.76×105
1.14×106
26