Heat integrated technology assisted pressure-swing distillation for the mixture of ethylene glycol and 1,2-butanediol

Journal Pre-proofs Heat Integrated Technology Assisted Pressure-Swing Distillation for the Mixture of Ethylene Glycol and 1,2-Butanediol Weixin Mao, Yueqiang Cao, Rongchun Shen, Jinghong Zhou, Xinggui Zhou, Wei Li PII: DOI: Reference:

S1383-5866(19)34741-0 https://doi.org/10.1016/j.seppur.2020.116740 SEPPUR 116740

To appear in:

Separation and Purification Technology

Received Date: Revised Date: Accepted Date:

15 October 2019 11 February 2020 17 February 2020

Please cite this article as: W. Mao, Y. Cao, R. Shen, J. Zhou, X. Zhou, W. Li, Heat Integrated Technology Assisted Pressure-Swing Distillation for the Mixture of Ethylene Glycol and 1,2-Butanediol, Separation and Purification Technology (2020), doi: https://doi.org/10.1016/j.seppur.2020.116740

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Heat Integrated Technology Assisted Pressure-Swing Distillation for the Mixture of Ethylene Glycol and 1,2-Butanediol Weixin Mao1, #, Yueqiang Cao2, #, Rongchun Shen1, Jinghong Zhou2,*, Xinggui Zhou2, Wei Li1,* 1 School of chemical engineering, East China University of Science and Technology, Shanghai, 200237, China 2. State Key Laboratory of Chemical Engineering, East China University of Science and Technology, Shanghai, 200237, China #

*

These authors contributed equally to this work.

Corresponding authors: [email protected]; [email protected]

Abstract: Production of ethylene glycol (EG) from coal derived syngas has been an alternative to the petroleum route, where a certain amount of mixture product of EG and 1,2-butanediol (BDO) is inevitably side-produced as a low value-added product. Thus, designing a separation process with high economic performance for this azeotropic mixture is highly desirable. Herein, pressure-swing distillation (PSD) processes without and with aid of heat integration are designed and simulated by using Aspen Plus for recovering high purity EG and BDO from the mixture product. The results show that the CO2 emissions and total annual cost (TAC) of the four modified PSD process have been significantly reduced compared with the optimized basic PSD process. Among these processes, the full heat integrated pressure-swing distillation process has the best economic performance with TAC decreasing by 35%, while the bottom flash pressure-swing distillation process has the best environmental performance with CO2 emissions decreasing by 95%. These findings provide fundamentals for the rational design of industrial separation and purification of EG and BDO azeotrope in the coal-based EG production.

Keywords: Pressure-swing distillation; ethylene glycol; 1,2-butanediol; heat integration technology

Nomenclature AC AR 𝑏 B1

auxiliary condenser auxiliary reboiler exergy of stream bottom stream of normal pressure column

B2 BDO BF BF-PSD CCC [𝐶𝑂2 ]𝑓𝑢𝑒𝑙 𝐶𝑂𝑃 CostAC

bottom stream of low pressure column 1,2-butanediol bottom flash pressure-swing distillation based on bottom flash heat pump technology cold composite curve CO2 emissions from burning the heavy oil fuel coefficient of performance annual cost of compressors

CostCA Costcap CostEA Costene Costves D D1 D2 DN DL EG FHI-PSD ℎ HCC ℎ𝑝𝑟𝑜𝑐

annual cost of column capital cost of column annual cost of exchangers energy cost of column capital cost of column vessel column diameter distillate stream of normal pressure column distillate stream of low pressure column diameter of normal pressure column diameter of low pressure column ethylene glycol full heat integrated pressure-swing distillation enthalpy of stream hot composite curve the enthalpy of stream delivered to the process

IDL IDN L LC LPC LR 𝐿𝑊 𝑛 NC NFD1 NFD2 NFF

diameter of low pressure column diameter of normal pressure column column length condenser of low pressure column low pressure column reboiler of low pressure column lost work of separation molar flowrate of stream condenser of normal pressure column feeding locations of distillate stream of normal pressure column feeding locations of distillate stream of low pressure column feeding locations of feed

𝑁𝐻𝑉 NPC NR NTL

the net heating value of a fuel with a carbon content normal pressure column reboiler of normal pressure column theoretical trays of low pressure column

NTN OPTL

theoretical trays of normal pressure column outlet pressure of throttle valve of low pressure column

OPTN PHI-PSD PRL

outlet pressure of throttle valve of normal pressure column partial heat integrated pressure-swing distillation compressor pressure ratio of low pressure column

PRN PSD 𝑄𝐶 QCW 𝑄𝐶𝑜𝑚𝑝

compressor pressure ratio of normal pressure column pressure-swing distillation condenser duty

𝑄𝑐𝑜𝑛𝑠 𝑄𝑓𝑢𝑒𝑙 𝑄ℎ QFH QLC QLR QNC QNR

energy consumption the amount of fuel burnt heat delivered at high temperature hot utility requirements for reboilers condenser duty of low pressure column reboiler duty of low pressure column condenser duty of normal pressure column reboiler duty of normal pressure column

𝑄𝑅 RRL RRN 𝑠 𝑇0 TAC 𝑇𝐶 𝑇𝐹𝑇𝐵

reboiler duty reflux ratio of low pressure column reflux ratio of normal pressure column entropy of stream the environmental temperature total annual cost temperature of condenser the theoretical flame temperature

𝑇𝑅 𝑇𝑠

temperature of reboiler the temperature of heat source

𝑇𝑠𝑡𝑎𝑐𝑘 VC VF VRC VRC-PSD 𝑊 𝑊𝑐𝑜𝑚𝑝

the stack temperature vapor compression vapor fraction vapor recompression pressure-swing distillation based on vapor recompression heat pump technology work input work into or out of compressor minimum work of separation

𝑊𝑚𝑖𝑛 Greek letters

cold utility requirements for condensers compressor duty

𝛼 𝜆𝑝𝑟𝑜𝑐 𝜂

the ratio of molar masses of CO2 and C the latent heat of stream delivered to the process the thermodynamic efficiency

1. Introduction Ethylene glycol, an important chemical raw material, is conventionally produced via ethylene oxidation originated from petroleum in industry. Recently, the production of EG from coal-based syngas increasingly attracts attention due to its potential in reducing the heavy dependence on the petroleum resources.[1] However, the purification for the EG produced from the coal-based process is more complicated than that from the petroleum-based process. Li et al.[2] have proposed a five-column separation process for refining EG, where a mixture product of EG and 1,2-butanediol (BDO) is extracted from the top of the de-butanediol column (Col 4 in Scheme 1). This mixture is difficult to be separated by conventional distillation for the boiling points of two components are extremely close and an azeotrope can be formed[3], thus normally used as low value-added products such as antifreeze agents, which definitely lowers the overall economy of the coal-based EG production process. Therefore, efficiently separating the mixture to produce highly valuable EG and BDO will greatly improve the economic benefits of the process.

Scheme 1. Schematic diagram for the five-column separation process for EG.

There are some methods for separating azeotropic mixture of EG and BDO, such as azeotropic distillation, extraction, and reactive distillation, which employ azeotropic agent, extractant, additional reactant, respectively.[3-6] However, the subsequent recovery for these mass separation agents definitely increases the complexity of separation process and thus increases the energy consumption and the separation cost. Therefore, designing a separation process with high energy efficiency and economic performance for the azeotropic mixture of EG and BDO to produce EG and BDO products of high purity is meaningful for coal-based EG production but still remains challenge. Fortunately, the azeotrope of EG and BDO are pressure-sensitive, which makes pressure-swing distillation (PSD) employing two columns operated at different pressures be possible for separating the azeotrope mixture without any mass separation agents and additional separation cost for its recovery.[7] Huizenga et al.[8] proposed a PSD process for separating EG and BDO mixture, and the product purity of 99.6 wt% and 90.0 wt% for EG and BDO streams were obtained, respectively, showing a promising potential in separating azeotrope of EG and BDO with PSD. Yet the EG product in their work still could not meet the requirements of the polyester industry for EG grade. Therefore, there is a need to search for more rigorous separation conditions, e.g., higher reflux ratio and/or more trays, for the PSD process to obtain high purity EG (≥99.9 wt%). Another concern is that even the optimized PSD process with more rigorous conditions result in high grade EG of 99.9 wt%, the correspondingly increased operation and capital cost still probably

reduce the feasibility of this process. Thus, more measures should be taken to further

reduce energy consumption and cost as well. Noting that the energy input in the reboiler at the

bottom of the column for a distillation process is removed in the condenser by the cold utility at the top of the column, which means the separation with distillation is achieved at the cost of energy loss and thus low thermodynamics efficiency.[9] On the other hand, the higher reflux ratio will lead to more energy loss and CO2 emission as the energy is obtained from burning fossil fuels. Thus, recovering the energy of the top vapor stream to improve the thermodynamic efficiency is a promising attempt to enhance the economic performance and reduce the CO2 emission of the PSD process. The typical strategy to address this issue is the heat integration technology, including the double-effect heat integration and the heat pump technology.[10] The former one is to employ the top vapor stream of the high pressure column to heat the reboiler of the low pressure column.[11] The latent heat of the top vapor stream can be recycled to reduce energy loss in this way.[9, 12-17] The latter one is to use compressors to pressurize the top vapor stream for heating the reboiler.[11, 18-20] The latent heat of the top vapor stream of both columns can be recovered to reduce energy loss in this case.[16, 21-24] Exploiting of both heat integration technologies in PSD process definitely improves the energy efficiency, but will also increase the complexity of the process and thus higher capital costs, which means energy cost compromises with capital cost. Thus, the PSD processes modified by the heat integration technology needs to be further investigated detailed for maximizing the total economic efficiency. So far, the PSD process with aid of heat integration technology for separation the EG and BDO mixture into high purity EG and BDO of 99.9 wt% has not been studied, to the best of our knowledge. Therefore, in this work, in order to recovering the high purity EG and BDO from their

mixture as low-valued side product in coal-based EG manufacture process, new heat integration assisted PSD processes are designed and evaluated for separating the mixture of EG and BDO into high grade EG and BDO (99.9 wt% for both) by using Aspen Plus. The design parameters of conventional PSD process are optimized to minimize the total annual cost (TAC). Four modified PSD processes with aid of heat integration technology have been proposed and carefully optimized, aiming to improve the economy benefit and reduce the CO2 emission. By comparing the energetic, environmental and economic performance, the optimum separation process is recommended for separating the azeotropic mixture of EG and BDO.

2. Methods 2.1 Property method The applicability and reliability of vapor-liquid equilibrium calculation depend on the thermodynamic model used. In this case, NRTL model is applied to describe the nonideality of the vapor-liquid equilibrium of the EG-BDO binary system. The binary interaction parameters used here are from corrected binary interaction parameters proposed by Li et al. [3] as shown in Table S1. The comparison of experimental and simulated data on the phase equilibrium is shown in Figure S1, in which the vapor-liquid equilibrium data calculated using the NRTL model are very close to the experimental data, suggesting the applicability of the NRTL model for this binary system.

2.2 Cost evaluation TAC, i.e., the sum of energy costs and annual capital costs, has been widely used in economic

evaluation of process design.[7, 25-30] The energy costs include the cost of hot and cold utilities as well as the cost of electricity for compressors. The capital costs consist of the costs of columns, condensers, reboilers, heat exchangers and compressors. The costs of the pumps, valves and pipes are not considered in the capital costs because they are negligible compared with those of the major equipment. The payback period is assumed to be 3 years, and the annual operating time of the factory is assumed to be 8000 h. The corresponding formulas used to estimate the TAC summarized in Table 1 are from the previous report.[7] Table 1. Basis of economics and equipment sizing Vessels Column diameter (D)/m Aspen tray sizing Column length (L) /m NT trays with 2 ft spacing plus 20% extra length Capital cost /$ 17 640(D)1.066(L)0.802 Condensers Heat-transfer coefficient /kW/K·m2 0.852 Differential temperature /K reflux-drum temperature − 315 Capital cost /$ 7296(area in m2)0.65 Reboilers Heat-transfer coefficient /kW/K·m2 0.568 Differential temperature /K steam temperature − base temperature (ΔT > 20 K) Capital cost /$ 7296(area in m2)0.65 Heat exchangers Heat-transfer coefficient /kW/K·m2 0.852 Differential temperature /K LMTD of (inlet and outlet temperature differences) Capital cost /$ 7296(area in m2)0.65 Compressors* Capital cost /$ (1296)(517.3)(3.11)(hp)0.82/280 Energy cost HP steam /$/GJ 9.88(41 barg, 254 °C) MP steam /$/GJ 8.22 (10 barg, 184 °C) LP steam /$/GJ 7.78 (5 barg, 160 °C) Cooling water /$/GJ 0.354 Electricity /$/GJ 16.9 (capital cost/payback period) + energy cost TAC /$/year 3 Payback period/year *

The isentropic efficiency and mechanical efficiency are set to 80% and 90%, respectively.

2.3 Energy consumption and thermodynamic efficiency The energy consumption (𝑄𝑐𝑜𝑛𝑠 ) is an index for analyzing energy usage, which is calculated as follows:[31] 𝑄𝑐𝑜𝑛𝑠 = 𝑄𝑅 + 3𝑄𝐶𝑜𝑚𝑝

(1)

where 𝑄𝑅 and 𝑄𝐶𝑜𝑚𝑝 are reboiler heat duty and compressor duty, respectively. The thermodynamic efficiency (𝜂) is an important index for assessing the efficiency energy use, which can be calculated according to the method proposed by Seader and Henley.[32] The expression of the thermodynamic efficiency is shown in Eq. (2~5). 𝑊

𝑚𝑖𝑛 𝜂 = 𝐿𝑊+𝑊

(2)

𝑚𝑖𝑛

𝑊𝑚𝑖𝑛 = ∑𝑜𝑢𝑡 𝑛𝑏 − ∑𝑖𝑛 𝑛𝑏 (3) 𝑏 = ℎ − 𝑇0 𝑠 (4) 𝑇

𝑇

𝐿𝑊 = ∑𝑖𝑛 [𝑛𝑏 + 𝑄𝑅 (1 − 𝑇0 ) + 𝑊𝑐𝑜𝑚𝑝 ] − ∑𝑜𝑢𝑡 [𝑛𝑏 + 𝑄𝐶 (1 − 𝑇0 ) + 𝑊𝑐𝑜𝑚𝑝 ] 𝑠

𝑠

(5) where 𝑊𝑚𝑖𝑛 (kJ/h) and 𝐿𝑊 are the minimum work of separation and the lost work, respectively. 𝑏 (kJ/kmol) is the exergy, and 𝑛 (kmol/h) is molar flowrate. ℎ (kJ/kmol) and 𝑠 (kJ/kmol∙K) are the enthalpy and entropy of stream into or out of system. 𝑇0 is the environmental temperature (K), and 𝑇𝑠 is the temperature of heat source (HP stream) or heat trap (cooling water). 𝑄𝑅 and 𝑄𝐶 (kJ/h) are the reboiler duty and condenser duty, respectively. 𝑊𝑐𝑜𝑚𝑝 (kJ/h) is the work into or out of compressor.

2.4 Evaluation of CO2 emissions CO2 emissions is the key index for assessing the environmental performance of the

process.[33] Evaluating the CO2 emissions of a process is complex, because there are a variety of traditional and new energy resources that can generate steam for heating the reboiler and electricity to drive the compressor.[12] Therefore, for the convenience of calculation, in this paper, heavy fuel oil is assumed as the source of steam generation, and a given CO2 emissions value (51.1 kg CO2/GJ) for the electricity of compressor.[34] According to the model for estimating the CO2 emissions proposed by Gadalla et al[35], the emission from the boilers, [𝐶𝑂2 ]𝑓𝑢𝑒𝑙 (kg/s), is assumed to be from burning the heavy oil fuel, which can be calculated by Eq. (6) and Eq. (7). 𝑄

𝐶%

[𝐶𝑂2 ]𝑓𝑢𝑒𝑙 = ( 𝑓𝑢𝑒𝑙) × ( ) 𝛼 𝑁𝐻𝑉 100 𝑄𝑝𝑟𝑜𝑐

𝑄𝑓𝑢𝑒𝑙 = (𝜆

𝑝𝑟𝑜𝑐

) × (ℎ𝑝𝑟𝑜𝑐 − 419) × (𝑇

(6) 𝑇𝐹𝑇𝐵 −𝑇0 𝐹𝑇𝐵 −𝑇𝑠𝑡𝑎𝑐𝑘

)

(7)

where 𝑄𝑓𝑢𝑒𝑙 (kW) is the amount of fuel burnt. 𝛼 (=3.67) is the ratio of molar masses of CO2 and C, and 𝑁𝐻𝑉 (kJ/kg) is the net heating value of a fuel with a carbon content of 𝐶%. For the heavy oil fuel, 𝑁𝐻𝑉 and 𝐶% are 39771 kJ/kg and 86.5, respectively.[35] 𝜆𝑝𝑟𝑜𝑐 (kJ/kg) and ℎ𝑝𝑟𝑜𝑐 (kJ/kg) are the latent heat and enthalpy of stream delivered to the process, respectively. The 𝑇𝐹𝑇𝐵 and 𝑇𝑠𝑡𝑎𝑐𝑘 are the theoretical flame temperature and stack temperature, assumed as 1800 ℃ and 160 ℃, respectively.

2.5 Heat pump performance The coefficient of performance (𝐶𝑂𝑃) is an index for evaluating the performance of heat pump, which is proposed by Bruinsma and Spoelstra.[18] The expression of 𝐶𝑂𝑃 is the ratio of heat (𝑄ℎ ) delivered to the hot reservoir at high temperature to the input work (𝑊), is shown in Eq. (8).

𝐶𝑂𝑃 =

𝑄ℎ

(8)

𝑊

The simple coefficient of performance 𝐶𝑂𝑃𝑆 is an index for evaluating whether the application of heat pump technology can improve the benefits of the process, which is proposed by Pleşu et al.[36] The expression of 𝐶𝑂𝑃𝑆 is shown in Eq. (9). 𝐶𝑂𝑃𝑆 =

𝑄𝑅 𝑊

=𝑇

𝑇𝐶 𝑅 −𝑇𝐶

(9)

Where, 𝑄𝑅 is the reboiler duty of column, 𝑊 is the work input, 𝑇𝑅 and 𝑇𝐶 is temperature (K) of reboiler and condenser. When the 𝐶𝑂𝑃𝑆 value exceeds 10, the use of a heat pump will bring benefits. Between 5 and 10, it should be evaluated more deeply. If it is lower than 5, then the heat pump should not be used.

3. Enhanced PSD process for EG and BDO mixture 3.1 Process design, description and simulation The PSD process was previously proposed to separate the mixture of EG and BDO according to Aspen simulation, but the purities of the obtained EG could not meet the requirement of polyester industry for EG grade[37]. It is necessary to design a new process to enhance the performance of the PSD technology. In principle, the prerequisite for a PSD process is that the difference in the azeotropic composition under two different operating pressures is larger than 5%.[38] The T-xy diagram of EG-BDO binary mixture under 100.0 and 10.0 kPa estimated by Aspen Plus is shown in Figure 1. The molar fraction of EG increases from 54.20 mol% at 100.0 kPa to 62.30 mol% at 10.0 kPa (Table S2), which can meet the requirement of a PSD process. Moreover, the azeotropic temperature decreases from 466.8 K at 100.0 kPa to 403.7 K at 10.0 kPa.

Within this temperature range, cooling water and high-pressure steam can be used as the cold and hot utilities rather than special coolant and heat agent. This can reduce the cost of utilities as exemplified by the comparison of the cooling water and low temperature refrigerated water as the coolants (Table S3). Thus, these two pressures are chosen as the operating pressures for the PSD process, i.e., 100.0 kPa for the normal pressure column (NPC) and 10.0 kPa for the low pressure column (LPC).

Figure 1. T-xy diagrams for EG/BDO at 10.0 and 100.0 kPa. The flowsheet of the PSD process is schematically shown in Figure 2. After completing the pressure selection, the flowrate and composition of the recycle stream (i.e., D2 stream) need to be further determined. This is because the recycle stream dominates the workload of the subsequent multi-variable optimization, and the corresponding optimization efficiency would be greatly improved with an appropriate initial values for the recycle stream. The flowrate, composition,

temperature and pressure of the feed for the PSD process are set at 50.00 kmol/h, 90.00 mol% of EG, 421.2 K and 20.0 kPa, respectively, according to the industrial operating parameters of the five-column separation process.[2] The design specification for the purities of EG and BDO products are both 99.90 wt%. Based on the above specification, the flowrates of the bottom streams of the NPC and LPC can be given by the component balance as Eq (10) and (11): 𝑛𝐹 = 𝑛𝐵1 + 𝑛𝐵2

(10)

𝑛𝐹 𝑥𝐹 = 𝑛𝐵1 𝑥𝐵1 + 𝑛𝐵2 𝑥𝐵2

(11)

where 𝑛𝐹 , 𝑛𝐵1 , and 𝑛𝐵2 are the flowrate of feed, B1 and B2 streams, respectively. 𝑥𝐹 , 𝑥𝐵1 , and 𝑥𝐵2 are the composition of feed, B1 and B2 streams, respectively. The compositions of the distillate streams of the NPC and LPC are set to be close to the azeotropic compositions at the two operating pressures, i.e., 55.00 and 61.50 mol% of EG, respectively. Based on the above specification, the flowrates of D1 and D2 streams can be also given by the component balance as below: 𝑛𝐷1 = 𝑛𝐷2 + 𝑛𝐵2

(12)

𝑛𝐷1 𝑥𝐷1 = 𝑛𝐷2 𝑥𝐷2 + 𝑛𝐵2 𝑥𝐵2

(13)

where 𝑛𝐷1 , 𝑛𝐷2 , and 𝑛𝐵2 are the flowrate of D1, D2 and B2 streams, respectively. 𝑥𝐷1 , 𝑥𝐷2 , and 𝑥𝐵2 are the composition of D1, D2 and B2 streams, respectively.

Figure 2. Schematic flowsheet of a basic PSD process.

It is noted that the flowrate and the composition of the LPC distillate stream (i.e., D2) are correlated with the heat duty, which influences the TAC of the total process. Therefore, the flowrate and composition of D2 stream were optimized starting with the initial parameters estimated above. The dependence of TAC on the composition and the flowrate is shown in Figure 3, in which the optimum composition and flowrate of D2 stream is 60.87 mol% of EG and 46.00 kmol/h, respectively.

Figure 3. Optimization results for the determination of the LPC distillate stream variables. 3.2 Process optimization and analysis According to the flowsheet of the PSD process with the given feed conditions and designed specifications shown in Figure 2, there are overall 9 variables, i.e., pressures of NPC and LPC, flowrate and composition of the recycle stream (i.e., D2 stream), theoretical trays of NPC and LPC, feeding locations of feed stream, D1 stream and D2 stream, to be optimized to minimize the TAC of the PSD process. The operating pressures of the two columns were first chosen according to the thermodynamic equilibrium data. As discussed above, the flowrate and composition of the recycle stream have been optimized towards a minimized TAC of the process based on the sensitivity analysis using Aspen plus. The left five degrees of freedom, i.e., the theoretical trays of NPC (NTN) and LPC (NTL), locations of feed stream (NFF), D1 stream (NFD1) and D2 stream (NFD2), were further optimized to minimize the TAC of the process by the sequential iterative optimization

method schematically shown in Figure 4. NFD1 was first optimized, and the results are shown in Figure 5, in which the optimum of NFD1 is 32. During the next optimization for NFF, NFD1 was set at 32, and the minimum of the TAC is obtained at the NFF of 83. Along this line, the optimum of NFD2, NTL and NTN can be obtained, as seen in the Figure 5.

Figure 4. Sequential iterative optimization procedure for PSD process.

Figure 5. Optimization results for the determination of PSD process design variables.

The optimized PSD process with the detailed parameters simulated by Aspen Plus is illustrated in Figure 6. The pressurized fresh feed and recycle streams are co-fed at the 83th and 53th tray of the NPC operated at 100.0 kPa, respectively, where the mixture is separated into the bottom stream (i.e., B1) with the 99.90 wt% of EG and the distillate stream (i.e., D1) with the composition close to the azeotropic mixture at 100.0 kPa. The D1 stream with 54.94 mol% of EG is depressurized to 25.0 kPa by the throttle valve and then fed at the 32th tray of LPC operated at 10.0 kPa, where the mixture is separated into the bottom stream (i.e., B2) with 99.90 wt% of BDO and the distillate stream (i.e., D2) recycled to NPC with the composition close to the azeotropic

mixture at 10.0 kPa. The detailed results of the optimized PSD process are summarized in Tables 2 and 3. The minimized TAC of the PSD process is 3.706 × 106 $/year, with energy cost of 2.727 × 106 $/year and capital cost of 2.936 × 106 $. It can be clearly seen that the energy cost of the PSD is much high, up to 73% of the TAC. Meanwhile, the estimated CO2 emission is as high as 4046.9 kg/h, probably due to the high energy consumption. These results indicate a great potential for enhancing the economic and environmental performance of the PSD process.

Figure 6. Flow sheet of the optimized PSD process. The NC, NR, RR and ID represent the condenser and reboiler of normal pressure column, reflux ratio and column diameter, respectively.

Table 2. Design parameters of the PSD process. Case PSD Parameters NPC NT 100 P/kPa 100.0

LPC 94 10.0

RR B/kmol/h NFD1 NFF NFD2 F/kmol/h

4.78 45.02

5.24 4.98 32

83 53 50.00

Table 3. Performance evaluation indexes for different designed process Case

PSD

PHI-PSD

FHI-PSD

VRC-PSD

BF-PSD

Capital cost/106 $

2.936

2.977

2.958

6.720

6.810

Energy cost/106 $/year

2.727

1.441

1.414

0.676

0.546

TAC/106 $/year

3.706

2.433

2.400

2.916

2.816

TAC saving/%

0

34

35

21

24

𝑄𝑐𝑜𝑛𝑠 /kW

9259.2

4893.1

4802.8

3642.1

3312.2

CO2 emissions/kg/h

4046.9

2138.6

2099.2

441.9

203.1

CO2 emissions saving/%

0

47

48

89

95

1.66%

3.60%

3.66%

5.46%

6.35%

0

117

120

229

283

Thermodynamic efficiency Thermodynamic efficiency increasing/%

4. PSD process with heat integration Heat integration is commonly employed to decrease the energy cost and thus increase the energy efficiency of separation processes. the heat integration was further employed for the PSD process. The temperature-enthalpy diagram (T-H)[39] is first used to see if it is possible to recover energy for the optimized PSD process. The hot composite curve (HCC) and the cold composite curve (CCC) shown in Figure 7 indicate the overall heat availability and the heat requirement for the process, respectively. The displacement of the HCC curve on the X-axis represents the cold

utility requirement (i.e., QCW), and that of the CCC curve represents the hot utility requirement (i.e., QFH). No overlap between the HCC and CCC curves is observed from the T-H diagram, suggesting no heat recovery in the PSD process. However, the temperature of HCC curve is higher than that of CCC curve at a certain region in Figure 7, i.e., some streams that need to be cooled in the process have higher temperatures than those that need to be heated. Thus, the energy consumption of PSD process will be reduced effectively upon using the streams with higher temperature to heat ones with lower temperature via the heat integration technology.

Figure 7. Temperature-enthalpy(T-H) diagram of PSD process

The temperature difference between the top stream of NPC (466.8 K) and the bottom stream of LPC (442.6 K) is larger than 20 K, which meets the requirements for the application of double-effect integration technology.[40] In addition, in the abovementioned PSD process, the

𝐶𝑂𝑃𝑆 values of NPC (21.8) and LPC (10.4) are higher than 10, meeting the requirements for the application of heat pump technology. Therefore, the double-effect integration technology and heat pump technology were further employed to improve the performance of the PSD process.

4.1 PSD process with double-effect heat integration The double-effect heat integrated pressure-swing distillation process is first designed, including the double-effect partial heat integrated pressure-swing distillation (PHI-PSD) and full heat integrated pressure-swing distillation (FHI-PSD) processes. As shown in Figure 8, in the PHI-PSD process based on the optimized parameters of the PSD process shown above, a heat transfer is employed to recover the latent heat of NPC top vapor steam heat the LPC reboiler. The pressurized fresh feed and recycle streams are co-fed into the NPC, where co-feed stream is separated into the bottom stream (i.e., B1) with the 99.90 wt% of EG and the top vapor steam with the composition close to the azeotropic mixture at 100.7 kPa. The top vapor steam is condensed into saturated liquid stream in heat exchanger (i.e., E1) with releasing the latent heat. Then, the condensed steam of E1 is split into two streams. One is fed into NPC as the reflux on the top stage, and the other one (i.e., D1) is fed into LPC through the throttle valve, where the NPC distillate is separated into the bottom stream (i.e., B2) with the 99.90 wt% of BDO and the LPC distillate stream (i.e., D2) recycled to NPC with the composition close to azeotropic mixture at 10.0 kPa.

Figure 8. Flow sheet of PHI-PSD process. The VF and AR represent the vapor fraction and auxiliary reboiler, respectively. Figure 9 shows the T-H diagram of the PHI-PSD process. QCW and QFH are 4728.8 and 4893.1 kW, respectively. Compared with the PSD process, the PHI-PSD process saves 47% of QFH and 48% of QCW, respectively. The overlap between HCC and CCC curves represents the amount of heat recovery in the process, which is calculated to be 4366.1 kW. The detailed performances of PHI-PSD process are summarized in Tables 3. The TAC of the PHI-PSD process is 2.433 × 106 $/year, with the energy cost of 1.441 × 106 $/year and the capital cost of 2.977 × 106 $. Compared to those of the PSD process, the TAC and energy cost of the PHI-PSD process decrease by 34% and 47%, respectively, and the capital cost increases by 1%, which is mainly due to decrease om the energy cost by recovering the latent heat of top vapor stream of NPC. The slight increase in capital cost of TAC is resulted from the addition of the new heat exchangers. In addition, the energy consumption and CO2 emissions both decrease by 47%, and thermodynamic

efficiency increases by 117%, suggesting that the PHI-PSD process can effectively reduce greenhouse gas emissions and the degree of irreversibility of the separation process in comparison to the PSD process.

Figure 9. Temperature-enthalpy(T-H) diagram of PHI-PSD process. Notably, in the PHI-PSD process, an auxiliary reboiler (i.e., AR1) is added since that the heat duty of LPC reboiler is not equal to that of NPC condenser, which increases the total capital cost and hot utilities. The FHI-PSD process without auxiliary reboilers is thus designed towards reduced capital cost and hot utilities. As shown in Figure 10, the pressurized fresh feed and recycled streams are co-fed into the NPC column, where co-feed stream is separated into the bottom stream (i.e., B1) with the 99.90 wt% of EG and the top vapor steam with the composition close to the azeotropic mixture at 100.7 kPa. The top vapor steam is condensed into saturated liquid stream in heat exchanger (i.e., E1) with releasing the latent heat. Then, the condensed steam

of E1 is split into two streams. One is fed into NPC as the reflux on the top stage, and the other one (i.e., D1) is fed into LPC through the throttle valve, where the NPC distillate is separated into the bottom stream (i.e., B2) with the 99.90 wt% of BDO and the LPC distillate stream (i.e., D2) with the composition close to azeotropic mixture at 10.0 kPa. The LPC distillate stream is recycled to the NPC. In this process, in order to achieve full heat integration, the reboiler duty of LPC (i.e., QLR) should be equal to the condenser duty of NPC (i.e., QNC). Therefore, the process is modeled in Aspen Plus by setting QLR - QNC = 0 as the design specification and the flowrate of LPC distillate stream as a variable.

Figure 10. Flow sheet of FHI-PSD process Figure 11 shows the T-H diagram of the optimized FHI-PSD process. QCW and QFH are 4638.5 and 4802.8 kW, respectively. Moreover, compared with the PHI-PSD process, the FHI-PSD process saves 48% of QFH and 49% of QCW, respectively. The amount of heat recovery in

this process is calculated to be 4456.6 kW. The detailed results of FHI-PSD process are summarized in Tables 3. The TAC of the FHI-PSD process is 2.400 × 106 $/year, with the energy cost of 1.414 × 106 $/year and the capital cost of 2.958 × 106 $. The TAC and energy cost of FHI-PSD process decrease by 35% and 48% respectively, and the capital cost increases by 1%, compared with PSD process. The energy consumption and CO2 emissions both decrease by 48%, and thermodynamic efficiency increases by 120% compared with PSD process. Moreover, compared with PHI-PSD process, TAC and energy cost of FHI-PSD process decrease by 1 and 2% respectively, and capital cost decreases by 1%, due to the absence of the auxiliary reboilers.

Figure 11. Temperature-enthalpy(T-H) diagram of FHI-PSD process 4.2 PSD process with heat pump technology Despite that 4456.6 kW of heat is recovered, the hot and cold utility requirements of the FHI-PSD process are still large, which means that only recovering the latent heat of NPC top

vapor steam is insufficient. Thus, heat pump assisted distillation technologies, including the vapor recompression pressure-swing distillation (VRC-PSD) and the bottom flash pressure-swing distillation (BF-PSD) processes, were employed to recover latent heat of both NPC and LPC to further improve the performance of the process.[41] It should be noted that the maximum temperature of the outlet vapor of compressors should be less than 150 ℃ to avoid the potential issues such as worn rings, liquid strike and oil breakdown.[21] The focus of this work is to compare the heat pump technology assisted PSD process to the basic one for gaining a sense that if and how much the performance of PSD process can be improved by such assisted technology. Therefore, the temperature limitation for the outlet vapor of the compressors are not considered in the VRC-PSD and BF-PSD processes. However, the potential problems of the compressors associated with the temperature of outlet vapor should be seriously considered in the practical application. The VRC-PSD process is first designed based on the PSD process with optimized parameters shown in section of 3.2, and the corresponding detailed information with simulated results are shown in Figure 12. In this process, the column top vapor stream is pressurized by the added compressor (i.e., Comp1) to increase its energy quality, which ensures that the high-quality latent heat can be used as the heat source for the reboiler. The pressurized fresh feed and recycled streams are co-fed into the NPC column, where the mixture is separated into the bottom stream (i.e., B1) with the 99.90 wt% of EG and the top vapor steam with the composition close to the azeotropic mixture at 100.7 kPa. The top vapor steam is compressed to increase the temperature and pressure. In heat transfer of E1, the NPC top vapor steam is condensed into high-pressure

saturated liquid phase stream, which is subsequently decompressed and partially vaporized in the throttle valve, and then condensed into saturated liquid phase through the auxiliary condenser (i.e., AC1). The liquid stream is further split into two streams. One is fed into the NPC column as the reflux on the top stage, and the other one is fed into the LPC column through the throttle valve. The LPC process is similar to the NPC process. The latent heat of LPC top vapor steam can meet the heat duty of the reboiler in the LPC process, and the auxiliary reboiler is thus not needed. The LPC bottom stream (i.e., B2) with 99.90 wt% of BDO is extracted as the product, while the distillate stream (i.e., D2) with the composition close to the azeotropic mixture at 10.0 kPa is recycled to NPC column. In NPC process, an auxiliary reboiler (i.e., AR1) is added since that the latent heat of top vapor steam is smaller than heat duty of the NPC reboiler.

Figure 12. Flow sheet of VRC-PSD process with optimized compressor pressure ratios for LPC and NPC columns. The AC represent the auxiliary condenser.

For VRC-PSD process, the amount of the latent heat released from the top vapor steam in the heat exchanger depends on the differential temperature driving force, which is determined by the compressor pressure ratio (i.e., PR, the ratio of outlet pressure to the inlet one).[12] Larger value of the PR endows higher temperature and pressure of the steam at the compressor outlet, which increases the differential temperature driving force, and thus reduces the capital cost of the heat exchanger. However, more driving energy of the compressor is necessary for the larger value of the PR, meaning the higher capital cost and energy cost of the compressor. The PR value is thus optimized by employing the TAC for the VRC-PSD process as the objective function, and the optimal PR of NPC (i.e., PRN) and LPC ratio (i.e., PRL) are 2.15 and 5.52, respectively. The VRC-PSD process with the optimal PRN and PRL are schematically shown in Figure 12. Based on the optimal PR values, the T-H diagram of the VRC-PSD process is shown Figure 13, where the QCW and QFH are 1335.5 and 581.6 kW, respectively, and the amount of heat recovery in the process is 8677.6 kW. Compared with the FHI-PSD process, the VRC-PSD process saves 88% of QFH and 71% of QCW, respectively. The TAC of the VRC-PSD process based on the optimized PR values is 2.916 × 106 $/year, and the energy cost and capital cost are 0.676 × 106 $/year and 6.720 × 106 $, respectively, as listed Table 3. Compared to those of the PSD process, the TAC and energy cost of VRC-PSD process decrease by 21 and 75%, respectively, and the capital cost increases by 129%. Meanwhile, the energy consumption and CO2 emissions decrease by 61 and 89%, respectively, and thermodynamic efficiency increases by 229%. Moreover, compared to those of the FHI-PSD process, the energy cost was reduced by 52%, suggesting that the energy cost can be effectively reduced by applying the heat pump technology.

Figure 13. Temperature-enthalpy (T-H) diagram of the VRC-PSD process with optimized compressor pressure ratios for LPC and NPC columns. In addition to the VRC-PSD process, the BF-PSD process is also considered here to improve the performance of the PSD process. The corresponding detailed information with simulated results are shown in Figure 14. In this process, the liquid stream from column bottom is depressurized in the throttle valve in order to decrease its bubble point and dew point temperature, which ensures the stream with low temperature and pressure can be used as the cold source for the condenser. The pressurized fresh feed and recycled streams are co-fed into the NPC column, where co-feed stream is separated into the liquid stream from the 99th tray of the NPC and the top vapor steam with the composition close to the azeotropic mixture at 100.7 kPa. The liquid stream is partially vaporized in the heat exchanger of E1, and then the partially vaporized stream is separated into vapor phase stream and liquid phase stream through the vapor-liquid separator (i.e., SEP1).

Then, the vapor phase stream and the liquid stream are compressed by the compressor and the pump, respectively. These two streams are mixed and heated to the specified vapor fraction by the superheated vapor-phase stream. The mixed stream with the specified vapor fraction is separated in another vapor-liquid separator (i.e., SEP2), where the vapor-phase stream is refluxed to the NPC column, and the liquid-phase stream with 99.90 wt% of EG (i.e., B1) is extracted as the product at the bottom. The top vapor steam of the NPC column is condensed with releasing the latent heat, which is subsequently decompressed and partially vaporized in the throttle valve, and then condensed into saturated liquid phase through the auxiliary condenser (i.e., AC1). The liquid stream is further split into two streams, which are fed into the NPC column as the reflux on the top stage and the LPC column through the throttle valve, respectively. The LPC process is similar to the NPC process, while the bottom stream (i.e., B2) with 99.90 wt% of BDO is extracted as the product, and the distillate stream (i.e., D2) with the composition close to the azeotropic mixture at 10.0 kPa is recycled to the NPC column.

Figure 14. Flow sheet of BF-PSD process with optimized outlet pressure of the throttle valve for LPC and NPC columns. For the BF-PSD process, the amount of the heat absorbed by the bottom stream in the heat exchanger depends on the differential temperature driving force, which is affected by the outlet pressure of the throttle valve (i.e., OPT). Lower OPT gives lower bubble point and dew point of the throttle valve outlet stream, which increases the differential temperature driving force, and thus reduce the capital cost. However, more driving energy of the compressor is necessary for the slower OPT, meaning the higher capital cost and energy cost of the compressor. Thus, the OPT values were optimized towards the minimized the TAC for the BF-PSD process. The optimal OPT for the NPC (i.e., OPTN) and LPC (i.e., OPTL) are 76.0 and 7.0 kPa, respectively. The BF-PSD process with the optimal OPTN and OPTL are schematically shown in Figure 14. The T-H diagram for the BF-PSD process based on the optimized OPT values (Figure 15) suggests that the QFH is close to zero. The calculated QCW is only 829.5 kW. Compared to the

VRC-PSD process, the BF-PSD process saves 100% of QFH and 38% of QCW, respectively. The amount of heat recovered in the optimized BF-PSD process is 8265.4 kW. The detailed performances of the BF-PSD process with optimized OPT values are summarized in Table 3. The TAC of the process is 2.816 × 106 $/year, and the energy cost and capital cost are 0.546 × 106 $/year and 6.810 × 106 $, respectively. Compared to those of the PSD process, the TAC and energy cost of BF-PSD process decrease by 24 and 80% respectively, and the capital cost increases by 132%; the energy consumption and CO2 emissions decrease by 64 and 95%, respectively, and thermodynamic efficiency increases by 283%. Moreover, compared to those of the VRC-PSD process, the TAC and energy cost of BF-PSD process decrease by 3 and 19%, respectively. These indicate that BF-PSD process is more attractive in economy than the VRC-PSD process.

Figure 15. Temperature-enthalpy(T-H) diagram of BF-PSD process

4.3 Discussion The economic and environmental efficiency, i.e. TAC and CO2 emissions, of the 5 PSD processes designed in this work are illustrated with histogram for easy comparison in Figure 16. In general, the modified PSD processes, utilizing whether double-effect heat integration technology or heat pump technology, have considerably reduced energy consumption and CO2 emissions (Figure 16b), thus enhancing the economic and environmental benefits. However, the energy cost reductions in these cases are all achieved at the cost of gain in capital cost. Among the modified PSD processes, the FHI-PSD process has the lowest of TAC of 2.400 × 106 $/year with the capital cost of 2.958 × 106 $, which is only 1% increase compared to the conventional PSD process, and achieves significant energy cost decreases by 48% (Figure 16a). This one could be the best choice for the industry to attempt to recovering high purity EG and BDO from their azeotrope according to the computation results in this work.

Figure 16. (a) Comparisons of different processes in terms of key indicators: capital cost, energy cost and TAC. (b) Energy consumption and CO2 emissions.

A comparison between the double-effect heat integration technology and the heat pump technology shows that the PSD processes assisted by double-effect heat integration technology presented remarkable energy cost (about 50%) decrease with slight capital cost increase (about 1%), while the PSD processes assisted by heat pump technology exhibited strikingly significant energy cost decrease (about 80%) with more significant capital cost increase (about 130%). This is because that double-effect heat integration technology needs to add only a heat exchanger, while the heat pump technology needs to add two compressors and two heat exchangers, which results in the capital cost increase and thus insufficient economic competitiveness. It can be known from the TAC formula in Table 1 that the capital cost accounts for a large ratio in the TAC, which is unfavorable for the heat pump technology assisted process with high capital cost. This explains why the TAC of heat pump technology assisted PSD processes are higher than those of the double-effect heat integration technology in this work. But TAC is related only to the annual capital cost rather than total capital cost, if a longer payback period such as 5 years or more is allowed, then the TAC of the process using heat pump technology will be more competitive under another circumstance. In addition to the economic performance, the climate and environmental performances are becoming more and more important nowadays for the new process and technology and for the sustainability of chemical industry. CO2 emission, as defined in Methods section, is employed to evaluate the climate and environmental performances. The process with the less CO2 emissions is undoubtedly more environmental-friendly. As seen in Figure 16b, the CO2 emissions for the 4 modified PSD processes are remarkably reduced compared to the conventional PSD process,

among which the BF-PSD process has the least CO2 emission. Compared to that of the basic PSD process, the CO2 emissions of the PHI-PSD, FHI-PSD, VRC-PSD, and BF-PSD processes decrease by 47, 48, 89, and 95%, respectively. Similarly, the energy consumption of the PHI-PSD, FHI-PSD, VRC-PSD, and BF-PSD processes decrease by 47, 48, 61 and 64%, respectively, compared to that of the basic PSD process. These suggest that the application of heat integration technology can reduce energy consumption and CO2 emissions more effectively, especially for the CO2 emissions, mainly due to the maximized recovery of hot utilities and thus the reduced the amount of fossil fuels required to burn for providing high pressure steam. The operating flexibility of the heat pump technology enhanced PSD processes were also addressed by comparing the basic PSD process with the VRC-PSD and BF-PSD processes under two different operating cases. In case 1, the feed flowrate is increased by 60% with the EG ratio in the feed mixture up to 95%, while in case 2, the feed flowrate is decreased by 60% with the EG ratio in the feed mixture down to 80%. Both the basic PSD process and the heat pump technology assisted PSD processes, either VRC-PSD or BF-PSD, are simulated under new cases, the detailed calculation results, including TAC, energy cost, CO2 emission and thermodynamic efficiency, are listed in Table S6. Apparently, as in the regular case we have discussed in the above, the TACs, energy costs and CO2 emissions of VRC-PSD and BF-PSD processes in both new cases, are remarkably smaller than those of the basic PSD process. Likewise, the thermodynamic efficiencies of the VRC-PSD and BF-PSD processes are also significantly increased, about 3 times of the corresponding values of the basic PSD process. These results exemplify the good operating

flexibility of the heat pump technology enhanced PSD process for separating the EG-BDO mixture.

5. Conclusions In this work, pressure-swing distillation (PSD) processes aided with heat integration technology were designed for recovering high purity EG and BDO from their mixture. The parameters of the basic PSD process were first optimized targeting minimum TAC by a sequential iterative optimization method. Then, the double-effect heat integration technology and heat pump technology were employed to reduce both energy consumption and CO2 emission and thus improve the performance and feasibility of the PSD process. Four modified PSD processes with aid of heat integration method, i.e., PHI-PSD, FHI-PSD, VRC-PSD and BF-PSD processes, have been proposed and carefully optimized. The analysis of T-H diagram and thermodynamic efficiency, as well as the TAC has been used to evaluate the modified PSD processes comparing with the conventional one. The results show that the CO2 emission and TAC of the four modified PSD process have been significantly reduced compared with the non-heat integrated PSD process. Among these processes, FHI-PSD process exhibits the best economic performance with TAC reduced by 35% over the conventional PSD process, while BF-PSD process shows the best environmental performance with CO2 emission reduced by 95%. Acknowledgements

This work was supported by Ministry of Science and Technology of the People’s Republic of China, under the Research Fund for National Key R&D Program of China (2018YFB0604700).

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Weixin Mao: Investigation, Visualization; Yueqiang Cao: Writing- Original draft preparation, Writing - Review & Editing; Rongchun Shen: Formal analysis, Methodology; Jinghong Zhou: Resources, Writing - Review & Editing, Project administration; Xinggui Zhou: Conceptualization Wei Li: Supervision, Validation

1. New pressure-swing distillation process for the mixture of EG and 1,2-butanediol. 2. Pressure-swing distillation process improved by heat integrated technology. 3. Full heat integrated PSD process shows the best economic performance. 4. Bottom flash PSD process exhibits the best environmental performance.

Declaration of interests ☒ The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper.

☐The authors declare the following financial interests/personal relationships which may be considered as potential competing interests: