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Enhanced Efficient Extractive Distillation by Combining Heat-Integrated Technology and Intermediate Heating Lumin Li, Yangqin Tu, Lanyi Sun,* Yafei Hou, Minyan Zhu, Lianjie Guo, Qingsong Li, and Yuanyu Tian State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Qingdao 266580, China ABSTRACT: In order to make full use of the sensible heat of the heavy entrainer adopted in the extractive distillation, an industrial separation case that focuses on the separation of benzene/ cyclohexane mixture using sulfolane as entrainer is studied in this work. Furthermore, several retrofitted processes are proposed using the basic case. A side reboiler that utilizes the self-heat or external heat as the heat source is introduced in the separation system to reduce the energy requirement. The economic feasibilities of these sequences are estimated by calculating the total annual cost. Moreover, the thermodynamic analysis and carbon dioxide emissions indexes are also considered to evaluate the energy efficiency and environmental impacts of the modified designs. The simulation results show that the extractive distillation with heat integration 2 (EDHI2) design, which utilizes the entrainer stream to heat the side reboiler and the bottom reboiler, is the most energy-conserving sequence. In addition, the EDHI2 sequence also shows apparent benefits on economics and environment aspects. on process intensification,9 like dividing-wall columns (DWC), heat-pump-assisted distillation, heat-integrated distillation, or multieffect distillation.10−13 An extractive dividing-wall column (EDWC) for ethyl acetate and isopropyl alcohol mixture separation was proposed by Zhang et al., and calculation results showed that the energy requirements and TAC of the advanced design can be reduced by 5% and 10%, respectively, compared with those of the conventional process.14 Sun et al.15 studied the separation of benzene and cyclohexane mixture using furfural as entrainer in an EDWC, and the results showed that more than 22% energy savings can be achieved by the EDWC process when compared with the conventional two-column process. In their work, low-pressure steam and mediumpressure steam are used in the extractive distillation column and entrainer recovery column, respectively, while only mediumpressure steam can be used in the EDWC reboiler, which explains why the steam cost and TAC savings in the EDWC case are only 1.8% and 4.8%, respectively. Luo et al.16 advocated a novel heat-pump-assisted EDWC sequence for bioethanol purification, and the results showed that approximately 24% TAC savings can be achieved for this innovative process. To separate the acetone−methanol azeotrope system, the pressureswing distillation and extractive distillation processes with and without heat integration were proposed by Luyben.17 In his work, the simulation results showed that the TAC of conventional pressure-swing distillation was higher than that of

1. INTRODUCTION Mixture separation is an important consideration in the petrochemical industry and environmental engineering, and distillation is the most common method used in separation process. However, ordinary distillation is ineffective to separate the azeotropic or close-to-boiling-point mixtures.1 Thus, it is necessary to adopt special distillation technology such as pressure-swing distillation, azeotropic distillation, and extractive distillation to achieve the separation target.2−4 Though the pressure-swing distillation can avoid the use of a third component to accomplish the separation, it is not suitable for the mixtures whose azeotropic composition is insensitive to pressure.5 Azeotropic distillation can separate the mixtures with the aid of an entrainer, although the azeotropic system also brings some challenges like multiple steady states and high energy requirement.6,7 Extractive distillation is a broadly used technology in chemical industry, which possesses the advantages of both distillation and extraction. For the extractive distillation system, the relative volatility of the feed mixture can be enhanced through introducing an entrainer, and hence, the products can be easily separated by distillation. The selection of an appropriate entrainer is significant for extractive distillation processes. One important thing is that entrainer and products should be chemically nonreactive and easily separated, and the availability, selectivity, and environmental impacts should also be considered at the entrainer screening stage. In addition, the entrainers are generally characterized by high boiling points for their nonvolatile products, which results in a large amount of energy requirements accompanied by high total annual cost (TAC) to recover the entrainer.8 Nowadays, lots of energy-saving technologies have been proposed and applied to the extractive distillation processes based © XXXX American Chemical Society

Received: March 24, 2016 Revised: July 1, 2016 Accepted: July 27, 2016

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DOI: 10.1021/acs.iecr.6b01152 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

Figure 1. Residue curve maps for the cyclohexane-benzene-sulfolane ternary system at (a) 1 atm and (b) 0.08 atm.

extractive distillation. In the case of heat integration, both the energy cost and TAC of extractive distillation were lower than those of pressure-swing distillation. Self-heat recuperation technology has been extensively studied in recent years. Kansha and co-workers applied this advanced technology to pure bioethanol production, and the simulation results showed that the energy requirement for the novel sequence was only 12.3% of the original distillation process.18 It should be noted that these attractive methods mentioned above are typical applications of process intensification in distillation process, which are explored to gain further energy savings through the possible combinations, and significant achievements have been obtained.19−22 However, few works have paid attention to the utilization of sensible heat in entrainer. It is well-known that the entrainer recovery operation, especially for the entrainer with high boiling point, is an energy-intensive and inevitable process, and a large amount of sensible heat in the entrainer stream is discharged without heat recycling in most extractive distillation processes. In this work, an industrial separation case, which focuses on the separation of benzene/cyclohexane mixture with sulfolane as entrainer,23 will be retrofitted to make full use of the entrainer sensible heat. The economical evaluation and thermodynamic analysis of the new processes are carried out to estimate their feasibilities. In addition, the environmental impacts of the proposed configurations are assessed by calculating their carbon dioxide (CO2) emissions.

Figure 2. Flowsheet of the conventional extractive distillation (CED) process for the separation of benzene−cyclohexane mixture.

distillation column (EDC) and an entrainer recovery column (ERC). The flowsheet of the CED process is shown in Figure 2, in which the operating parameters are very close to a real industrial process as reported by Qin et al.23 Sulfolane is chosen as a typical organic entrainer in the extractive distillation process especially adopted for the separation of aromatics from hydrocarbon mixtures. Sulfolane has a low thermal decomposition temperature of 493.15 K, and thus, it is suggested that the operation temperature is under 493.15 K. In this study, to guarantee the highly efficient separation, operation pressure of the ERC is set at 0.08 atm, leading to the base temperature with 478.49 K, and the stream from the ERC base is available as a heat source. However, in the present CED case, the recycled entrainer is immediately cooled and then fed into EDC, and a great deal of sensible heat is wasted without consideration. Because the temperature of a distillation column determines the quality of heat source required in the system, it is possible to reduce the exergy loss of distillation system if the temperature profiles of columns are changed. An effective method is adopting a side reboiler or side condenser to the distillation system, which can redistribute the heat of the reboiler/condenser. Figure 3 presents the temperature and composition profiles of the CED process. As shown in Figure 3a, the temperature profile in ERC displays a steep slope that allows introducing an intermediate heating. Therefore, in order to improve this extractive distillation process, several modified configurations will be discussed in the following sections. 2.2. Improved Extractive Distillation Processes. It can be seen from Figure 2 that the temperatures in the bottom of the EDC and ERC are 385.06 and 478.49 K, respectively. Thus, the temperature of the recycled entrainer stream is high enough

2. PROCESS DESIGN OF EXTRACTIVE DISTILLATION 2.1. Conventional Extractive Distillation Process (CED). A residue curve map (RCM) calculated by the nonrandom two liquid (NRTL) model can be used as a convenient tool to evaluate the feasibility of extractive distillation sequences for the benzene−cyclohexane−sulfolane system. Figure 1 gives the RCM for this ternary system at 1 and 0.08 atm. As shown in Figure 1, there is no additional azeotrope that formed with sulfolane at both 1 and 0.08 atm. The sulfolane is the stable node; the benzene/cyclohexane azeotrope is the unstable node; both benzene and cyclohexane are saddles. Because there is no distillation boundary for this ternary system, extractive distillation process will be an efficient method for the separation of benzene−cyclohexane mixture. In this paper, the extractive distillation processes are conducted using the Aspen Plus simulator based on the NRTL physical properties. The specifications for the products are 99.9 mol % and the entrainer is 99.99 mol %. The conventional extractive distillation (CED) sequence includes an extractive B


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Figure 3. (a) Temperature profiles and (b) composition profiles in the CED process.

composition profiles for this EDHI1 process are given in Figure 5. It can be seen that the temperature profile in EDC is almost unchanged, and the steep temperature profile in ERC is improved; simultaneously, high purity products can be obtained in the two columns. 2.2.2. Extractive Distillation with Heat Integration 2 (EDHI2). In the EDHI1 process, the temperature of the top vapor from EDC is only 353.92 K, and thus, the heat exchange in side reboiler is limited and the reboiler duty of ERC is still considerable (see Figure 4). In addition, it was found that the sensible heat of the entrainer was still substantial even though it has been used to heat the EDC reboiler. An attractive idea is to make full use of the sensible heat of the entrainer, through which the duties in both ERC reboiler and entrainer cooler can be further reduced, and this configuration, called EDHI2, is illustrated in Figure 6. As shown in Figure 6, the bottom stream of ERC is divided into two parts that are used to heat the EDC reboiler and the ERC side rebiler, respectively, and then the two streams are mixed and cooled by the entrainer cooler. Compared with the EDHI1 design, the side reboiler duty of EDHI2 design is obviously increased due to the higher temperature of heat source, and the duties in the ERC reboiler and entrainer cooler are simultaneously decreased. Figure 7 presents the corresponding temperature and composition profiles in EDHI2 process. 2.2.3. Extractive Distillation with Heat Integration 3 (EDHI3). Because the EDHI2 design utilizes part of the recycled entrainer to heat the ERC side reboiler, the total heat exchange can only reach about 438 kW (see Figure 6), and the bottom reboiler duty of ERC is 645.2 kW, which still consumes large amounts of high-pressure steam. To investigate the saving potentials of high-pressure steam required in ERC, an alternative method is proposed that uses the external heat source with low quality to heat the side reboiler; as a result, the side reboiler duty will increase, and consequently, the bottom reboiler duty will decrease.

to provide the heat required in the EDC. In addition, it is worth noting that the temperature profile of the ERC is steep due to the wide boiling points of the mixture, which brings an opportunity to further reduce the energy consumption by introducing intermediate heating. 2.2.1. Extractive Distillation with Heat Integration 1 (EDHI1). An improved extractive distillation with heat integration 1 (EDHI1) is proposed on the basis of the CED process, and Figure 4 reveals the corresponding flowsheet that utilizes

Figure 4. Flowsheet of the extractive distillation with heat integration 1 (EDHI1).

the sensible heat of the recycled entrainer. As shown in Figure 4, the low-pressure steam provided to the tower kettle of EDC in CED sequence can be saved. In the CED process, the temperature of the top vapor stream from EDC is 353.92 K, which is higher than the temperature in stage 6 of ERC (332.04 K), as shown in Figure 3a. Thus, the top vapor stream from EDC can exchange heat energy with the stream withdrawn from stage 6 of ERC as shown in Figure 4. The temperature and C


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Figure 5. (a) Temperature profiles and (b) composition profiles in EDHI1 process.

the corresponding extractive distillation with heat integration 4 (EDHI4), and the reboiler duty of ERC is only 556.71 kW. Figure 12 displays the temperature profiles and composition profiles of the EDHI4 process, and it also can be seen that the temperature profile of the ERC is apparently improved compared with that in CED design. 2.2.5. Extractive Distillation with Thermal Coupling (EDTC). In the conventional extractive distillation column, there is a sharp temperature increase from stage 28 to 29, which is mainly caused by the decrease of benzene composition on stage 29. This remixing phenomenon can be observed in Figure 3b, which will cause unwanted energy consumption and decreased energy efficiency of the ERC. Studies show that thermally coupled distillation configurations are effective alternatives to reduce or eliminate the remixing,24,25 and thus, in this section, the thermally coupled distillation sequence is applied to the benzene/cyclohexane separation system by combining the EDC and the ERC. The proposed extractive distillation with thermal coupling (EDTC) configuration is given in Figure 13. In the EDTC process, the reboiler of the EDC is removed, and a vapor stream withdrawn from the ERC is compressed and fed to the EDC bottom. The liquid stream from the EDC bottom flows into the ERC as the feed stream. In addition, the recycled entrainer with high temperature is used as the heat source of the side reboiler. Figure 14 presents the temperature and composition profiles of the EDTC design. It can be seen that the remixing phenomenon is apparently alleviated in this new sequence. To provide a clear understanding about the heat integration of the proposed configurations, detailed information on the heat exchangers in the EDC and ERC is given in Table 1.

Figure 6. Flowsheet of the extractive distillation with heat integration 2 (EDHI2).

Therefore, it is possible to further reduce the total steam cost of this system because the amount of high-pressure steam is decreased. Figure 8 presents this alternative extractive distillation with heat integration 3 (EDHI3). As shown in Figure 8, the heat load of side reboiler is 691 kW, which is higher than that in the EDHI2 design, and the bottom reboiler duty of ERC is significantly reduced. The corresponding temperature profiles and composition profiles are given in Figure 9. 2.2.4. Extractive Distillation with Heat Integration 4 (EDHI4). The influence of the temperature of feed to ERC (TF2) on the ERC reboiler duty and the benzene product purity is presented in Figure 10. As shown in Figure 10, the benzene product purity almost maintains at 99.9 mol % when the TF2 increases from 70 to 128 °C, and at the same time, the reboiler duty shows a sustained downward trend. Hence, using the hot recycled entrainer stream to heat the feed to ERC is a feasible method to reduce the energy consumption of the CED process, which is also much simpler than implementing a side reboiler or exchanging heat with the EDC reboiler. Figure 11 presents

3. PERFORMANCE EVALUATION METHODS In this work, three commonly used criteria including total annual cost, thermodynamic efficiency, and CO2 emissions are used to evaluate the proposed configurations. The TAC is selected to D


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Figure 7. (a) Temperature profiles and (b) composition profiles in the EDHI2 process.

where the Wmin [kJ/h] and LW [kJ/h] are the minimum work of separation and the lost work, respectively, which can be described as eqs 3 and 4:

∑

Wmin =

out of system

nb −

∑

nb (3)

into system

b is the availability function and is defined as b = h − T0s, where n [kmol/s] is mole flow rate; h [kJ/kmol] is enthalpy; s [kJ/kmol K] is molar entropy; and T0 [K] is the surrounding temperature. Ex = Wmin + LW = Figure 8. Flowsheet of the extractive distillation with heat integration 3 (EDHI3).

⎛ T ⎞ Q R ⎜1 − 0 ⎟ − TR ⎠ ⎝ into system

∑

⎛ T ⎞ Q C⎜1 − 0 ⎟ + Ws TC ⎠ ⎝ out of system

∑

(4)

where Ex [kW] presents exergy consumption of the system; QR [kW] and TR [K] are heat duty and temperature of reboilers, respectively; QC [kW] and TC [K] are heat duty and temperature of condensers, respectively; and Ws [kW] is shaft work. The thermodynamic efficiency of the different configurations can be calculated on the basis of the simulation results using the above equations. Distillations are energy-intensive processes that contribute greatly to CO2 emissions. Therefore, reducing the energy requirements of distillation processes will not only cut down the TAC but also decrease greenhouse gases emissions. The total CO2 emissions can be used to evaluate the environmental benefits from the energy-conserving configurations. The CO2 emissions generated from fuel combustion can be described as the stoichiometric eq 5:

evaluate the economic feasibility of the five energy-saving designs, and it can be described as eq 126 TAC ($/year) = operating costs + capital costs/payback period (1)

where the operating costs include the steam cost, cooling water cost, chilled cost, and electricity cost, and their corresponding prices are given in Table 2.5,13,27 The capital costs contain the cost for the trays, column shell, heat exchangers and compressor, and a mean payback period of 3 years is assumed to calculate the annual capital cost. The overall heat transfer coefficients and Marshall & Swift index (M&S) are adopted from the work of Li et al.28 The thermodynamic efficiency (η) can be used to estimate the energy economy of the distillation systems. The thermodynamic efficiency reported by Seader et al.29 is shown in eq 2: Wmin η= LW + Wmin (2)

⎛ y⎞ y Cx Hy + ⎜x + ⎟O2 → xCO2 + H 2O ⎝ ⎠ 4 2

(5)

where x and y are defined as the number of carbon C and hydrogen H atoms existing in the fuel compositions, respectively, E


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and excess air is assumed in the reaction to avoid forming carbon monoxide. The total CO2 emissions, [CO2]Emiss [kg/s], can be described as eq 6:30 ⎛Q ⎞⎛ C % ⎞ ⎟α [CO2 ]Emiss = ⎜ Fuel ⎟⎜ ⎝ NHV ⎠⎝ 100 ⎠

(6)

where QFuel [kW] is the amount of fuel burnt; NHV [kJ/kg] denotes the net heating value of a fuel with a carbon content of C% (dimensionless); α (= 3.67) is the ratio of molar masses between CO2 and C.

4. SIMULATION AND OPTIMIZATION To match the industrial case as mentioned above, the parameters of the conventional extractive distillation process remain unchanged,23 and the product purities can be satisfied through adjusting the reflux ratio and distillate rate of each column. There are some variables in the proposed configurations needed to be optimized, which includes the number of stages in the EDC (NT1), the number of stages in the ERC (NT2), the feed stage locations of entrainer and the fresh feed streams in the EDC (NFE, NFF), the flow rate of entrainer (FE), the feed location in the ERC (NFM), the flow rate (FW), and position (NFW) of the withdrawal in the ERC. The minimum TAC of the modified configurations can be calculated using the following optimization procedure. (1) Fixed the top pressures of the EDC and ERC at 1 and 0.08 atm, respectively. (2) Set the entrainer flow rate (FE). (3) Set the total stages of the EDC (NT1). (4) Set the feed stages of the EDC (NFE, NFF). (5) Adjust the reflux ratio and distillate rate of the EDC until the product specifications are met. (6) Go back to step 4 and alter NFE and NFF until the reboiler duty of the EDC is minimized.

Figure 10. Influence of the feed to ERC temperature (TF2) on the reboiler duty of ERC and the benzene product purity.

Figure 11. Flowsheet of the extractive distillation with heat integration 4 (EDHI4). F


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It should be noted that once the total stages of the ERC is determined, the position of the withdrawal can be confirmed due to the limited available minimum approach temperature. Figure 15 presents the optimum withdrawal flow rates in the four modified configurations. Note that QR means the reboiler duty of the ERC in the EDHI1, EDHI2, and EDTC processes, and QT is the sum of the side reboiler duty and bottom reboiler duty in the ERC of the EDHI3 process. The optimized parameters of the CED, EDHI1, EDHI2, EDHI3, EDHI4, and EDTC processes are summarized in Table 3.

5. RESULTS AND DISCUSSION Regarding the energy consumption, the EDHI2 design achieves the largest energy savings because of the synergetic advantage of the two types of self-heat recuperation. Though the energy consumptions of the modified sequences are all smaller than those of the CED process, the crucial index, TAC, should also be considered to evaluate the economic feasibility of the proposed designs. Therefore, the TACs of all configurations are calculated on the basis of the simulation results, and the corresponding economic results are presented in Table 4. It can be seen that the most economic EDHI2 design can achieve more than 18.75% TAC savings when compared with CED design, and the suboptimal EDHI4 design also has a 17.59% TAC savings. In addition, the EDTC sequence is not recommended because of the high capital cost and TAC, although it can relieve the remixing phenomenon of EDC. To get a clear comparison, the detailed cost items for the CED and modified sequences are illustrated in Figure 16. As shown in Figure 16, the operating costs in the modified processes are all lower than that in the CED process because of the decreased energy consumption. In terms of capital costs there is little difference among all sequences, except the EDTC design for the large compressor cost.

Figure 13. Flowsheet of the extractive distillation with thermal coupling (EDTC).

(7) Go back to step 3 and alter NT1 until the TAC is minimized. (8) Go back to step 2 and alter FE until the TAC is minimized. (9) Set the total stages of the ERC (NT2). (10) Set the feed stage of the ERC (NFM). (11) Set the flow rate of the withdrawal in the ERC (FW). (12) Adjust the reflux ratio and distillate rate of the ERC until the product specifications are met. (13) Go back to step 11 and alter FW until the reboiler duty of the ERC is minimized. (14) Go back to step 10 and alter NFM until the reboiler duty of the ERC is minimized. (15) Go back to step 9 and alter NT2 until the TAC is minimized. G


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Figure 14. (a) Temperature profiles and (b) composition profiles in the EDTC process.

Table 1. Information of the Heat Exchangers in the Proposed Extractive Distillation Processes EDHI1 hot stream flow, kmol/h Tin, K Tout, K pressure, atm mole fraction cyclohexane benzene sulfolane cold stream flow, kmol/h Tin, K Tout, K pressure, atm mole fraction cyclohexane benzene sulfolane QHX, kW

EDHI2

EDHI3

EDHI4

EDTC

EDC

ERC

EDC

ERC

EDC

ERC

ERC

ERC

109.81 478.49 434.68 0.124

18.94 354.15 353.92 1.007

50.47 478.49 378.29 0.124

59.25 478.49 356.56 0.124

110.02 474.27 430.28 0.124

433.15 433.15 5.922

110.00 478.49 400.16 0.124

110.00 474.43 370.10 0.110

trace trace 1

0.999 0.001 trace

trace trace 1

trace trace 1

trace trace 1

trace trace 1

trace trace 1

188.54 376.93 385.02 1.189

110.00 333.33 352.14 0.105

189.07 376.82 384.85 1.189

122.00 333.79 386.22 0.105

188.74 376.83 384.81 1.189

100.00 332.47 433.15 0.105

160.00 343.95 399.15 0.099

90.99 368.95 462.42 0.095

0.001 0.416 0.583 304

trace 0.099 0.901 157

0.001 0.418 0.581 310

trace 0.098 0.902 438

0.002 0.415 0.583 304

trace 0.103 0.897 691

0.001 0.312 0.687 534

trace 0.033 0.967 700

EDHI2 design as an example, the vapor flow profiles of the ERC in the CED and EDHI2 designs are displayed in Figure 17. As shown in Figure 17, the vapor flow rate above the side reboiler location of the EDHI2 design is similar to that of the CED design, and the vapor flow rate below the side reboiler is much lower than that of the basic case. This phenomenon can be explained as follows. On one hand, the bottom reboiler duty of the ERC in EDHI2 design (645.2 kW) is much lower than that in the CED process (1091.03 kW), and thus, the corresponding vapor flow from the column bottom in EDHI2 design is lower than that in the CED case. On the other hand, the introduction of the side reboiler alters the column

Table 2. Utility prices utility

price

high-pressure steam (41 barg, 254 °C) low-pressure steam (5 barg, 160 °C) chilled water cooling water electricity

$9.88/GJ $7.78/GJ $4.43/GJ $0.354/GJ $16.9/GJ

In the EDHI1, EDHI2, EDHI3, and EDTC designs, the introduction of the side reboiler may change the original vapor flow rate of the ERC, which will have an adverse effect of increasing the column diameter. Taking the most economic H


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Figure 15. Effects of the withdrawal flow rate on the reboiler duty of the (a) EDHI1, (b) EDHI2, (c) EDHI3, and (d) EDTC designs.

temperature distribution and increases the vapor flow rate above the side reboiler location. Therefore, in this case, the addition of a side reboiler leads to an overall reduction on the vapor flow rate near the bottom of the ERC. Note that the column diameter is usually set by the maximum vapor rate, and hence, the diameter of the ERC in the EDHI2 design is smaller than that in the CED design (see Table 3). This trend also can be found in the cases of the EDHI1, EDHI3, EDHI4 designs. Because the reductions of the diameters of the ERC for the modified designs are very small, the corresponding costs for column shell remains about the same. And the slight increase of capital cost for the EDHI1, EDHI2, and EDHI3 designs mainly comes from the additional cost of side reboiler. Because the EDHI4 design has no side reboiler in the ERC, its total capital cost is smaller than that of the CED design. Table 5 lists the calculation results of the thermodynamic efficiency and carbon dioxide emissions of the proposed configurations for the benzene/cyclohexane separation system. As shown in Table 5, the minimum work of separation for the CED system is essentially equivalent to that of other configurations. The thermodynamic efficiency of EDHI1, EDHI2, EDHI3, and EDHI4 designs is higher than that of the CED design, and for the EDHI2 design, it is greatly improved from 30.59% to 44.64%. It also demonstrates that the EDTC sequence is not an effective separation method for the low thermodynamic efficiency. The CO2 emissions for different sequences show a similar trend to their energy consumptions. As a result, the EDHI2 design presents the lowest CO2 emission with 329.25 kg/h among the modified configurations. Therefore, from the analysis results, it can be concluded that the EDHI2 design reveals not only a high thermodynamic efficiency but also a greater advantage than other designs in the economic and environmental aspects.

Table 3. Comparison of the Optimum CED, EDHI1, EDHI2, EDHI3, EDHI4, and EDTC Designs designs EDC total stages (NT1) reflux ratio diameter, m entrainer flow rate, kmol/h entrainer position (NFE) fresh feed position (NFF) entrainer cooler duty, kW condenser duty, kW reboiler duty, kW ERC total stages (NT2) reflux ratio diameter, m mixture position (NFM) withdrawal position (NFW) withdrawal flow rate, kmol/h condenser duty, kW side reboiler duty, kW reboiler duty, kW compressor duty, kW energy savings

CED

EDHI1 EDHI2 EDHI3 EDHI4 EDTC

29 1.419 1.032 110 4 20 931 1005 304

29 1.420 1.033 110 4 20 625 834 304

29 1.436 1.036 110 4 19 180 1012 310

29 1.420 1.033 110 4 19 619 1006 304

29 1.419 1.032 110 4 20 393 1005 304

28 2.179 1.106 110 4 20 200 1321 0

10 0.425 1.461 5

10 0.411 1.361 5 6 122 676 438 645

10 0.450 1.380 5 6 105 694 728 368

10 0.425 1.367 5

1091

10 0.390 1.350 5 6 120 666 172 901

0

35.4%

53.7%

21.4%

38.3%

7 0.270 1.754 4 4 91 608 700 769 166 9.2%

682

682 557

6. CONCLUSIONS The heavy entrainer with high boiling point is usually added to the extractive distillation system to alter the relative volatility of the mixture and realize easy separation, which consequently requires large amounts of energy to recover the entrainer. In the present work, five modified configurations, which can make full use of the sensible heat of the entrainer stream, are proposed for the separation of I


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Industrial & Engineering Chemistry Research Table 4. Economic Results for the CED, EDHI1, EDHI2, EDHI3, EDHI4, and EDTC Designs capital costs, 103$ operating costs, 103$/y TAC, 103$/y TAC savings, %

CED

EDHI1

EDHI2

EDHI3

EDHI4

EDTC

1618.475 485.449 1024.941 0

1647.381 356.220 905.346 11.67

1652.590 281.930 832.793 18.75

1613.651 372.958 910.842 11.13

1562.414 323.865 844.669 17.59

2874.458 392.709 1350.861 −31.80

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel.: +86 13854208340. Fax: +86 0532 86981787. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This work was supported by the National Natural Science Foundation of China (Grant: 21276279 and Grant: 21476261) and the Key Research and Development Plan Project of Shandong Province (Grant: 2015GGX107004) and the Fundamental Research Funds for the Central Universities (Grant: 15CX06042A). Lastly, the authors are grateful to the editor and the anonymous reviewers.

Figure 16. Comparison of the operating cost, capital cost, and TAC for the CED and modified processes.

■

Figure 17. Vapor flow profiles for the ERC of the CED and EDHI2 designs.

benzene/cyclohexane mixture. Energy savings can be obtained in these modified processes, and the corresponding TACs are calculated to evaluate their economic benefits. The results indicate that the TACs of the retrofitted sequences are all reduced to various degrees except the EDTC design, and the EDHI2 design can achieve the most TAC savings compared with CED design. Lastly, the thermodynamic analysis and CO2 emissions are compared for the proposed processes. The calculation results show that the proposed designs except the EDTC one also show advantages on the energy efficiency and environmental benefits aspects. It can be concluded that the heat-integrated extractive distillation, which makes good use of the sensible heat in the heavy entrainer, is an effective way to gain the economic and environmental benefits.

NOMENCLATURE CED = conventional extractive distillation CO2 = carbon dioxide DWC = dividing-wall column EDC = extractive distillation column EDHIi = extractive distillation with heat integration i EDTC = extractive distillation with thermal coupling EDWC = extractive dividing-wall column ERC = entrainer recovery column Ex = exergy consumption of the system FE = flow rate of entrainer FW = flow rate of the withdrawal in the ERC LW = lost work M&S = Marshall & Swift index NFE = feed stage location of entrainer NFF = fresh stream location in the EDC NFM = feed location in the ERC NFW = withdrawal position of the ERC NHV = net heating value of a fuel NRTL = nonrandom two liquid NT1 = the number of stages in the EDC NT2 = the number of stages in the ERC QC = heat duty of the condenser QFuel = the amount of fuel burnt QHX = heat transfer amount of the heat exchanger QR = heat duty of the reboiler RCM = residue curve map TAC = total annual cost

Table 5. Calculation Results of the Thermodynamic Efficiency and Carbon Dioxide Emissions for the Conventional and Modified Extractive Distillation Configurations minimum work of separation, kW lost work, kW thermodynamic efficiency total CO2 emissions, kg/h CO2 emissions savings

CED

EDHI1

EDHI2

EDHI3

EDHI4

EDTC

90.20 204.67 30.59% 712.10 0

90.20 154.55 36.85% 459.93 35.4%

90.20 111.88 44.64% 329.25 53.8%

90.20 94.08 48.95% 559.47 21.4%

90.20 111.60 44.70% 439.51 38.28%

90.20 275.80 24.64% 634.65 10.9%

J


Article


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TC = condenser temperature TF2 = temperature of the feed to ERC Tin = inlet temperature Tout = outlet temperature TR = reboiler temperature Wmin = minimum work of separation Ws = shaft work α = the ratio of molar masses of CO2 and C η = thermodynamic efficiency

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