Improving the performance of heterogeneous azeotropic distillation via self-heat recuperation technology

Accepted Manuscript Title: Improving the performance of heterogeneous azeotropic distillation via self-heat recuperation technology Authors: Jingxing Chen, Qing Ye, Tong Liu, Hui Xia, Shenyao Feng PII: DOI: Reference:

S0263-8762(18)30602-6 https://doi.org/10.1016/j.cherd.2018.11.022 CHERD 3433

To appear in: Received date: Revised date: Accepted date:

3 May 2018 25 August 2018 16 November 2018

Please cite this article as: Chen, Jingxing, Ye, Qing, Liu, Tong, Xia, Hui, Feng, Shenyao, Improving the performance of heterogeneous azeotropic distillation via self-heat recuperation technology.Chemical Engineering Research and Design https://doi.org/10.1016/j.cherd.2018.11.022 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.

Improving

the

performance

of

heterogeneous

azeotropic distillation via self-heat recuperation technology

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Jingxing Chen, Qing Ye*, Tong Liu, Hui Xia, and Shenyao Feng

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Jiangsu Key Laboratory of Advanced Catalytic Materials and Technology, School of Petrochemical Engineering, Changzhou University, Changzhou, Jiangsu 213164,

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China

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AUTHOR INFORMATION

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Highlights:

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*Tel.: +86 519 86330355. Fax: +86 519 86330355. E-mail: [email protected].

The self-heat reception technology was applied to heterogeneous distillation.

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The vapor compression system was optimized to reduce energy consumption.

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The heat exchanger network was used to design the heat integration.

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The HAD-SHRT process can reduce energy consumption and cost effectively.

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Abstract Azeotropic distillation is one of the most energy intensive processes and widely used to separate thermos-sensitive azeotropes. In this work, several heterogeneous

azeotropic distillation processes based on the vapor recompression heat pump (VRHP) and self-heat recuperation technology (SHRT) are proposed to study the energy-saving performance. The results show that the heterogeneous azeotropic distillation process based on self-heat recuperation technology (HAD-SHRT) has better performance in economic and environment. In HAD-SHRT process, the sensible heat and latent heat of the vapor streams are recuperated and supplied to the reboilers. The remaining

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sensible heat is supplied to the preheaters through the heat exchanger network (HEN).

Adding a preheater can reduce the duty of compressor by increasing the initial

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temperature of the stream. Compared with the conventional heterogeneous distillation

process, the HAD-SHRT process can save 38.50% of TAC, 93.30% of CO2 emissions, and 80.92% of energy consumption. The thermodynamic efficiency can increase to

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13.51%.

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Keywords: azeotropic distillation, energy-saving, self-heat recuperation technology.

Nomenclature

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Abbreviations

A-DWC = azeotropic dividing-wall column

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CCC = cold composition curve EX = the exergy of the process

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GCC = the grand composition curve HAD = heterogeneous azeotropic distillation HADC = heterogeneous azeotropic distillation column

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HAD-FP = HAD with feed preheating HAD-SHRT = HAD process based on SHRT HAD-VRHP1 = HAD with vapor recompression heat pump HAD-VRHP2= HAD-VRHP with compressor inlet preheating HCC = hot composition curve HEN = heat exchanger network

NF2= the feed stage of the HADC NFAQ = the feed stage of the aqueous phase NFF = the feed stage of the fresh feed NT1 = the total stages of the PDC NT2 = the total stages of the HADC PDC = preconcentration distillation column

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Pi = the inlet pressure of the compressor PSD = pressure swing distillation

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Pt = the outlet pressure of the compressor QC = the heat duty of the condenser Qcomp = the compression duty

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Qcons = energy consumption

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QCW = the cold utility requirements

QR = the heat duty of the reboiler

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Qpre = the heat duty of the preheater

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QFH = the hot utility requirements

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Qrecovery = the amount of heat recovery S = entropy

SHRT = self-heat recuperation technology

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T0 = the temperature of the environment TAC = total annual cost

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T-butanol = Tert butanol

Tc = the temperature of condenser TCC = total capital cost

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TOC = total operating cost TR = the temperature of reboiler U = Overall heat-transfer coefficients VRC = vapor recompression column VRHP = vapor recompression heat pump Wmin = the ideal work of the process

1 Introduction Since mixture separation by distillation columns consumes large amounts of fossil fuels as an energy source, energy-saving has attracted growing attention from many chemical industries and some energy-saving technologies for distillation have been developed [1]. Azeotropic distillation, one of the most energy-consuming processes,

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has a lower operating temperature and is more suitable for the separation of the thermos-sensitive system compared with ordinary distillation. The heterogeneous

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azeotropic distillation process takes advantage of a natural liquid−liquid separation in

a decanter [2–4]. With the azeotropic distillation process and by selecting toluene as the entrainer to form a binary heterogeneous azeotrope with water, Wu et al. [5] easily

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separated the pyridine and water mixture into an organic phase and an aqueous phase

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by a natural liquid−liquid separation technology. Yu et al. [6] studied the separation of

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ethylene glycol and neopentyl glycol via azeotropic distillation by using the entrainer of para-xylene, which interacted with ethylene glycol to form a heterogeneous

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azeotrope at atmospheric pressure. The results indicate that the process can achieve

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99.9 mol% of ethylene glycol and neopentyl glycol. In conventional azeotropic distillation process, the low-boiling azeotropes are distilled from the top of the column, with a large amount of heat supplied to the bottom reboiler, while the overhead vapor

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stream is directly condensed by cold utility. However, the heat supply to the reboiler is mostly discarded at the condenser. To reduce energy consumption and greenhouse gas

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emissions, some researchers designed an azeotropic dividing wall column (A-DWC) to improve thermodynamic efficiency. Dejanović et al. [7] designed a packed dividingwall column for an aromatic processing plant, and it saved 43.3% of energy

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consumption compared with a two-columns-in-series configuration. Yu et al. [8] proposed an A-DWC for the t-butanol dehydration system, and they saved 23.8% of energy consumption and 19.93% of total annual costs compared with the conventional process. Recently, energy consumption reduction in distillation has been widely studied by using a vapor recompression heat pump (VRHP), in which the overhead vapor stream

is compressed to raise temperature and the compressed stream is utilized as a heating medium in the reboiler. In other words, VRHP can upgrade low-grade heat to highquality energy and then provides the heat for the reboiler. The bubble point temperature of the compressed stream should surpass the temperature of the bottom stream for the sake of successful heat exchange. Hossein et al. [9] proposed a vapor recompression column (VRC) in methanol-water separation, and it reduced energy consumption by

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49.6% compared with the conventional distillation process. In the process, the reboiler

was heated by the heat of the overhead vapor without any external heat input, indicating

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the introduction of VRC can significantly decrease energy consumption. However, the heat of the overhead vapor stream can be only partially recovered and supplied to the reboiler, while the resting sensible heat is directly condensed by cold utility. Some

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researchers also designed profitable A-DWC equipped with VRHP. Liu et al. [10]

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found the heat pump-assisted distillation reactive dividing wall columns for n-propyl

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propionate synthesis saved 9.6% of total annual cost with a payback period of 8 years. Li et al. [11] designed A-DWCs with different VRHP models to study the energy-saving

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effect and found the VRHP-based A-DWC saved 32.22% of total annual cost.

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To further save energy, Kansha et al. [12] developed a self-heat recuperation technology (SHRT) which can greatly reduce the energy consumption in distillation process. The heat of the SHRT process is divided into two parts, including (1) the latent

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and sensible heat of the compressed stream, which is recuperated to heat the reboiler, and (2) all the remaining sensible heat, which is recuperated and supplied to the

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preheater. Thus, all the latent and sensible heat in the process can be recuperated without any external heat, and the electric power for the compressor becomes the only source of energy consumption. Nguyen et al. [13] proposed an SHRT-based natural gas

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liquid recovery process, in which the heat of the compressed vapor was used to heat the bottom streams and feed stream. The feed stream was also heated by the bottom product streams. The advanced process saved 67.19% of annual operating cost compared with the conventional process. Christopher et al. [14] proposed a SHRT- based energysaving process for propylene-propane separation, in which the heat of the compressed stream was supplied to the reboiler, and the remaining sensible heat was used to heat

the feed and the compressor inlet stream. The advanced process saved 45% of energy consumption compared with the process with simple vapor recompression. Xia et al. [15] designed a novel SHRT-based energy-saving pressure swing distillation (PSD) process, which saved 72.39% of energy consumption compared with conventional PSD. Kansha et al. [16] proposed a new SHRT-based azeotropic distillation process for ethanol-water system, in which the heat of the vapor stream was recovered by

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compressors and exchanged with the heat of the corresponding preheater and reboiler. Simulations showed the new process reduced 87.7% of energy consumption compared

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with the conventional process. Although many studies indicate the good performance

of SHRT in reducing energy consumption, few studies show its ability to further decrease total annual cost by reducing the duty of vapor compressor. Thus, optimizing

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the vapor compressor system of SHRT-based azeotropic distillation is a potential

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choice.

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The heat exchanger network (HEN) is an effective way to design the heat integration in SHRT-based distillation process. HEN integrates networks of all hot

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streams and cold streams to optimize the energy utilization. Song et al. [17] optimized

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the hydrogen production process by HEN, and the results showed the energy consumption was reduced by 61.5% compared with the conventional process. Poddar et al. [18] designed a reactive distillation process for biodiesel production by setting up

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HEN to design the heat integration, and found the energy consumption was reduced obviously compared with the original process. The above cases imply that HEN can

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make full use of all the heat of the process with the help of an optimal heat integration. However, few researchers have proposed the energy- saving processes by combining

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HEN and SHRT into azeotropic distillation. Although the energy-saving advantages of these methods have been discussed, the

effect of SHRT on the azeotropic distillation system has rarely been studied. In this study, the conventional azeotropic distillation is optimized firstly by preheating the fresh feed and reflux streams. Then different processes are proposed based on the concepts of VRHP and SHRT. The vapor compression system was optimized in terms

of energy consumption and total annual cost（TAC）, and all the heat of the SHRTbased process can be fully utilized by HEN. All processes were optimized to obtain the optimal operating conditions and compared in terms of energy consumption, total annual cost, and CO2 emission. 2 Design

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Tert butanol (T-butanol), as an organic synthesis intermediates, is widely used to make high purity isobutylene through a dehydration process. However, t-butanol

usually forms a homogeneous azeotrope with water at atmospheric pressure. In

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previous studies, Ka-Man et al. [19] investigated an extraction distillation process for tert-butanol dehydration. The savings of total annual cost and energy consumption of the proposed design are very significant. However, the pressure of the entrainer

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recovery column is set as 0.02atm, which is very difficult to operate. Yu designed a

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heterogeneous azeotropic distillation process for the t-butanol dehydration system [8].

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Cyclohexane is selected as the entrainer. The t-butanol/water/ cyclohexane mixture

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has a ternary heterogeneous azeotrope at 65.2℃. The NRTL model is used to predict the thermodynamic behaviors, and the model parameters can be seen in Table1. In this

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process, the reboilers must provide large amounts of energy for the vaporization of the azeotropic mixture, which are distilled from the top of the columns. To improve

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energy efficiency, some energy-saving processes are proposed, and the conventional heterogeneous azeotropic distillation process is selected as a basic process.

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2.1 Conventional Heterogeneous Azeotropic Distillation Process (HAD) Figure 1 shows the optimized conventional heterogeneous azeotropic distillation

process for t-butanol dehydration. The feed flow rate is 100 kmol/h, with 50 mol% t-

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butanol and 50 mol% water. Both t-butanol and water product specifications are set at 99.8 mol%. Cyclohexane is selected as the entrainer to form a ternary heterogeneous azeotrope with t-butanol and water. Other details and key parameters have been calculated in Yu et al.[8] The fresh feed is fed into a preconcentration distillation column (PDC), where most of the water is removed. The specified purity water is obtained at the bottom of PDC, and the composition of the distillate of the PDC is near

to that of the binary azeotrope, which flows to the heterogeneous azeotropic distillation column (HADC). In the HADC, 99.8 mol% t-butanol is produced at the bottom. The ternary azeotrope mixture distilled from the top is condensed and separates into two liquid phases in a decanter. The aqueous phase containing high concentration of water is recycled back to the PDC. The organic phase, which is rich in t-butanol and cyclohexane, is brought into the HADC.

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2.2 HAD with Feed Preheating（HAD-FP）

In the conventional azeotropic distillation process, the temperature of the fresh

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feed is 25 ℃, and the temperature of the decanter is set at 30℃. The temperatures of

the feed, aqueous phase and organic phase are all much lower than the tray temperatures, respectively. Thus, large amount of heat is required to heat and vaporize these liquid

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streams. Meanwhile, vapor streams are directly cooled at overhead condensers, which

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leads to more cold utility requirement. To improve the energy efficiency, the conditions

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of the feed, aqueous phase and organic phase must be improved. As shown in Figure 2, these streams are all preheated to the bubble point by adding three preheaters. The fresh

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feed is preheated by E1 before entering PDC, the aqueous phase and organic phase are

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preheated by E2, E3, and then recycled back to the columns, respectively. The preheating streams give the column some heat to reduce the reboiler duty. It is noted that the new process（HAD-FP）need to reoptimize the operating parameters because

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of the changes of the feed conditions. The total annual cost (TAC) is selected as the economic indicator in optimization

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procedure. The TAC is divided into annualized capital cost and the total operating cost (TOC). The annualized capital cost equals the total capital cost (TCC) divided by a payback period. The calculation data and equipment sizing refer to Luyben’s paper

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[20,21], which are listed in Table 2.The formula for TAC calculation is as follow: TAC = TOC +

TCC Payback period

There are five main design variables to be determined in HAD-FP process, including the total stages of the PDC (NT1), the feed stage of the fresh feed (NFF), the feed stage of the aqueous phase (NFAQ), the total stages of the HADC (NT2), and feed

stage of the HADC (NF2). PDC and HADC are optimized through a sequential iterative strategy as follows: Optimization for PDC (1) Select the total number of stages of PDC (NT1) (2) Select the feed stage of the fresh feed (NFF) (3) Select the feed stage of the aqueous phase (NFAQ)

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(4) Vary reflux ratio and reboiler duty to meet product specifications

(5) Go back to step 2,3 and vary NFF ,NFAQ until the TAC of PDC is minimized

Optimization for HADC (1) Select the total number of stages of HADC (NT2)

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(2) Select the feed stage of the HADC (NF2)

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(6) Go back to step 1 and vary NT1 until the TAC of PDC is minimized

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(3) Vary reflux ratio and reboiler duty to meet product specifications

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(4) Go back to step 2 and vary NF2 until the TAC of HADC is minimized (5) Go back to step 1 and vary NT2 until the TAC of HADC is minimized

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Figure 3 (a) shows the influence of NT1 on the TAC, Figure 3 (b) shows the

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influence of NT2 on the TAC. It is clear that the TAC is minimal when NFF is 7, NFAQ is 11, NT1 is 12, NF2 is 1, and NT2 is 32. More favorable conditions and detailed information can be seen in Figure 2. Compared with HAD process, the total reboiler

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duty is decreased from 6000 kW to 3717.4 kW, the total condenser duty is decreased from 5695 kW to 4207.6 kW, the total preheater duty is 808.2 kW. Thus, adding

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preheaters is a feasible way to reduce the energy consumption. Figure 4 shows the temperature-heat flow (T-H) diagram of the HAD-FP. QCW

represents the cold utility requirements, while QFH represents the hot utility

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requirements. The hot utility requirements and cold utility requirements of the HADFP are 4525.6 kW and 4207.6 kW, respectively. As shown in Figure 4, there is no overlap between hot composition curve (HCC) and cold composition curve (CCC), which means the heat cannot be exchanged between hot and cold streams in the system. The reboilers require a large amount of heat from the hot utility, and the condensers also require cold utility. Therefore, HAD-FP process has huge development space in

reducing the energy consumption. 2.3 HAD with Vapor Recompression Heat Pump（HAD-VRHP1） The HAD-FP process uses the preheaters to reduce the energy consumption, but its effect is not obvious. In the process, the vapor streams from the top columns are condensed to saturated liquid by cooling water, so massive latent heat is wasted. In order to make full use of latent heat, the vapor recompression heat pump (VRHP) will

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be adopted. Two compressors are added to compress the top vapor streams, then the compressed streams supply heat to the reboilers, respectively. Because of the internal

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heat transfer, both hot utility requirements and the cold utility requirements can be significantly reduced.

Figure 5 shows the flowsheet of the HAD-VRHP1. The preheated feed is fed to

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the PDC, where the fresh feed is separated into high purity water and the vapor stream

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with a near azeotropic mixture. The distillate from the PDC is compressed in a

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compressor (C1) to increase its pressure and temperature, and then the compressed stream (S1) exchange the latent and sensible heat in the reboiler (E4), where the stream

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is changed from vapor to liquid. The liquid stream (S3) is condensed in cooler 1 and

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split into two streams. One is recycled back to PDC, the other is brought into HADC. In HADC, the overhead vapor is compressed in the compressor (C2) to increase energy grade; the discharged stream (S2) can supply heat to the reboiler (E5); the

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outlet stream (S4) is condensed in cooler 2 and then flows to the decanter. In the decanter, mixture splits into an organic and an aqueous phase. The organic phase is

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preheated in E3 before entering HADC. The aqueous phase, containing mostly water, is preheated in E2 and recycled back to PDC. The parameters and operating conditions of the columns are maintained, so only

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the vapor compression system which contains two compressors, two exchangers and two coolers needs to be optimized. Figure 6 shows the influence of compression ratio on temperature difference and compressor duty. The compression ratio (RC) represents the ratio of the outlet pressure of the compressor with the inlet pressure of the compressor. The temperature difference, which refers to the temperature difference of

the heat exchanger, is usually over 5 ℃[22,23]. However, as the compression ratio increases, the temperature difference and the duty of compressor are all obviously increased. As a result, the temperature difference is chosen as 5 ℃; the compression ratio of C1 is set as 2.86; and the compression ratio of C2 is set as 3.65. In PDC, the temperature of the overhead vapor is compressed to 127.6 ℃ , and then the compressed stream S1 is decreased to 103.3 ℃ after heating the bottom stream in E4.

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The discharged stream S3 is cooled to 80 ℃ in cooler 1 and divided into two liquid

streams. In HADC, the temperature of the vapor stream is compressed to 117.6℃, and

stream S4 is further decreased to 30 ℃ by the cooler 2.

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the compressed stream S2 decreased to 89.5℃ through E5. The temperature of outlet

Figure 7 shows the T-H diagram of the HAD-VRHP1. Qrecovery is the total latent

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and sensible heat of the outlet streams S1, S2 which is used to heat bottom streams of

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PDC and HADC, respectively. QFH comprises the hot utility requirement of the E1,

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E2, and E3, while QCW comprises the cold utility requirement of the Cooler 1. Cooler 2. The amount of heat recovery is about 3702.4 kW. The hot utility requirements and

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the cold utility requirements are about 804.9 kW and 993.0 kW, respectively. Through

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the comparison between HAD-FP and HAD-VRHP1, HAD-VRHP1 can save 82.2% of QFH and 76.4% of QCW. It can be included that the VRHP is an effective way to reduce the hot and cold requirements in the process. However, the large amount of

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electric power for the compressors is acquired. To achieve the greatest benefit of the VRHP, the duty of the compressor must be decreased.

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2.4 HAD-VRHP with Compressor Inlet Preheating（HAD-VRHP2） In HAD-VRHP1, large amount of latent heat and sensible heat of the discharged

streams from the compressors is used to heating the reboilers. However, the huge

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electricity demand of the compressors cannot be neglected. To achieve further energy saving, the overhead vapor is brought into a preheater to increase the initial temperature of inlet stream before its compression. Figure 8 shows the improved process with compressor inlet preheating (HADVRHP2). In PDC, the overhead vapor is preheated in E3, then the preheated stream is compressed in C1 to upgrade the energy level. The energy of the discharged stream is

used to heat the E4 and the minimum temperature difference of the reboiler is still maintained at 5℃. The vapor stream from HADC is also preheated in a preheater before being compressed in C2. To achieve the optimal operating conditions, the temperature of the preheated stream (Tp) and compression ratio need to be determined. Figure 9 shows the effects of Tp on the compression ratio and TAC in PDC and HADC, respectively. In Figure 9 (a), as the temperature of the preheated stream increases, the

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compression ratio decreases quickly and then keep steady. When the temperature of

the preheated stream increased to 88.4℃, TAC reaches the lowest and compression

preheated to 83.7℃, and compression ratio is set as 2.49.

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ratio is set as 2.36. In Figure 9 (b) the minimum TAC exists when the top streams is

Figure 10 shows the T-H diagram of the HAD-VRHP2. The hot utility

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requirements and cold utility requirements of the HAD-VRHP2 are 929.6 kW and

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992.8 kW, respectively. The amount of heat recovery is about 3702.4 kW. With the

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addition of two preheaters, the hot utility of HAD-VRHP2 is slightly increased compared with that of HAD- VRHP1. However, the duty of C1 is decreased from

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148.81 kW to 123.46 kW, and the duty of C2 is decreased from 357.56 kW to 258.21

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kW. The energy consumption equals hot utility requirements plus three times the duty of compressor [24,25], therefore the decrease of the compressor duty can reduce the energy consumption. On the other hand, the preheaters of the process requires hot

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utility for heating, while the heat of the discharged streams (S3, S4) is wasted. Thus, the hot utility requirements and the cold utility requirements of HAD-VRHP2 can be

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further reduced.

2.5 HAD process based on SHRT（HAD-SHRT） In HAD-VRHP1 and HAD-VRHP2, the overhead vapor streams are compressed

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to increases the energy level and then the compressed streams provide the most of heat for the reboilers. However the remaining sensible heat of the streams is discarded. In addition, the heat of the bottom product streams is neglected in previous processes. To achieve further energy saving, the self-heat recuperation technology (SHRT) is applied to the process, which can utilizes not only the sensible and latent heat of the overhead vapor streams but also the heat of other hot streams in the process. First, the

heat of the vapor stream is recuperated and provided to the reboiler. Then the sensible heat of all hot streams is supplied to the preheaters by heat integration. All the heat of the process is recuperated without any external heat. HEN is an effective way to make full use of all the heat of the process by designing the heat integration. Thus, the HEN is used to integrate networks of all hot streams and cold streams to optimize the energy utilization in HAD-SHRT process. The HEN is optimized by reducing the external

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heat and heat exchangers. Table 3 summarizes all of the hot and cold streams in HADSHRT. Figure 11 shows the optimal HEN design based on minimal TAC of the heat

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exchangers. It comprises 8 self-heat exchangers, 3 condensers and 2 compressors. The self-heat exchangers use the latent and sensible heat of the hot stream, and the condensers use cooling water. The whole HEN demand 332.6 kW cold utility and 0

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kW hot utility

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Figure 12 shows the final HAD-SHRT process with optimal parameters. The

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fresh feed is preheated via E1, E2 and then fed to the PDC. The vapor stream of PDC is heated by E3 (25.2 kW) and compressed in C1 (123.46kW), the latent and sensible

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heat of the C1 discharged stream is used to heat the bottom stream in E4 (1324.3kW),

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the rest sensible heat is used to provide heat for E3 (25.2 kW) and E6 (99.5 kW), the outlet stream of E6 is cooled in cooler1 (33.15 kW). Last, the stream is separated into two liquid streams and recycled back to the top of PDC and HADC, respectively. After

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providing heat for the preheater E5 (58.81 kW), the product stream of PDC is cooled to 30℃ in cooler 2 (12.65 kW). In HADC, the top stream is preheated by E6 (99.5

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kW) and compressed in C2 (258.21 kW), the thermal energy of the C2 discharged stream is used to heat the bottom stream in E7 (2378.1kW), the rest sensible heat is used to heat the fresh feed and the aqueous phase of the decanter in E2 and E8,

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respectively, the discharged stream is cooled to 30℃ in cooler 3 (286.8 kW) and flows to the decanter. In decanter, the mixture splits into an organic and an aqueous phase. Organic phase is preheated in E8 (447 kW) and fed at the top stage of HADC, while aqueous phase is preheated in E5 (58.81 kW) and fed to the tenth tray of PDC. The product stream of HADC provides heat for the preheater E1 (197.7 kW). In vapor recompression system, the discharged stream from compressor is

superheated vapor, so the design of the exchanger consists of two distinct regions of heat transfer: desuperheating and condensing sections [26]. It is noteworthy that the overall heat-transfer coefficient (U) depends on the phases on the two sides of the heat exchanger. When the hot side is vapor phase, a value of 0.28 kW m−2 K−1 is assumed with boiling liquid on the cold side of the reboiler. When the vapor reaches dew-point temperature, a value of 1.2 kW m−2 K−1 is assumed with condensing vapor and boiling

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liquid on the two sides of the reboiler. In desuperheating section (E7A), the

superheated vapor discharged from compressor is cooled to dew-point temperature

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(95.5 °C) at the pressure of 2.52 bar. The heat-transfer duty is 171.2 kW. The

temperature of boiling liquid on the cold side is 84.5 °C. With a U = 0.28 kW m−2 K−1, the heat-transfer area is 30.5 m2. The temperature difference at the hot end is 33.1 °C

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(117.6 - 84.5 = 33.1 °C). At the cold end, the temperature difference is 11.0 °C (95.5-

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84.5 = 11.0 °C). The average logarithmic temperature difference is 20.10°C. In condensing section (E7B), the saturated vapor is cooled into saturated liquid. The heat-

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transfer duty is 2206.9 kW. With a U = 1.2 kW m−2 K−1, the heat-transfer area is 240.8

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m2. The temperature difference at the hot end is 11.0 °C (95.5- 84.5 = 11.0 °C). At the

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cold end, the temperature difference is 5.0 °C (89.5- 84.5 = 5.0 °C). The average logarithmic temperature difference is 7.66°C. The heat-transfer area of the reboiler is equal to the total area in two heat exchangers of desuperheating and condensing

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sections[27],so the total area of the reboiler is 271.3 m2 (30.5 + 240.8 = 271.3 m2). Figure 13 shows the conditions in the reboilers for the HAD-SHRT process. Table 4

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presents the detailed data for the HAD-SHRT process. Figure 14 shows the T-H diagram of the HAD-SHRT. The shading between HCC

and CCC represents the amount of heat recovery (Qrecovery), and QCW represents the

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cold utility requirements. As shown in Figure 14, Qrecovery of HAD-SHRT consists of two parts: latent heat and sensible heat. The whole process only requires 332.6 kW cold utility, and the hot utility is no longer required. Thus, the energy consumption of HAD-SHRT is much lower than HAD-VRHP2. Figure 15 shows the grand composition curve (GCC) diagram of the HAD-SHRT. The shading (I, II, III) in the Figure 15 represents the heat recovery pockets where the heat can be exchanged

between hot and cold streams. It is obvious that HAD-SHRT only requires 332.6 kW cold utility. 3 Performance evaluation methods All processes will be compared based on economic performance, energy analysis and environmental analysis. The economic performance is evaluated by calculating TAC, and the calculation formula and basic data have been explained in the previous

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chapter. In addition, although adding the compressors and preheaters can achieve expected effects in energy consumption, it is difficult to achieve the maximum

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economic benefits in short time because of the rising TCC. Thus, the pay-back period is set as 5 and 8 years in this paper.

Energy consumption (Qcons) is a main parameter to calculate the amount of

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energy, while the thermodynamic efficiency (η) is a key parameter to evaluate the

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quality of energy in exergy analysis. The calculation formulas of Qcons and η are as

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follows

Qcons = QR + Qpre + 3Qcomp

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EX = H − T0 S

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Wmin = ∑ (EX ) − ∑ (EX ) product

EXQ = ∑ QR (1 −

T0 T0 ) − ∑ QC (1 − ) + Wcomp TR TC outof system

η=

Wmin EXQ

CC E

PT

into system

feeds

The QR, QC represents the heat duty of the reboiler and condenser, respectively,

the Qpre (kW) is the heat duty of the preheater, and the Qcomp (kW) is the duty of

A

compressor. The factor of 3 for the duty of compressor is assumed to convert the work of the compressor into thermal energy with same effort of the electrical work [10]. T0 (K) is the temperature of the environment and H, S represents the enthalpy (J/mol) and entropy (J/mol·K), respectively. EX (J/mol) is the exergy of the process, Wmin is the ideal work of the process, TR and Tc (K) represent the temperature of reboiler and condenser, respectively.

Additionally, the carbon emission is a key factor that must be considered in environmental performance. The carbon emissions are usually related to energy inputs, and the electrical energy mainly comes from burning fossil fuels. In this study, the emission index of steam is considered as 224 kg CO2/t, and the electrical energy is taken as 51.1 kg CO2/GJ [28,29]. 4 Comparison

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Table 5 summarizes the detailed comparisons of all processes. To have fair

comparison, the result of azeotropic dividing-wall column (A-DWC) investigated in

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recent paper is also listed in Table 5 [11].

Compared with the HAD process, the energy consumption and carbon emission of HAD-FP is reduced by 24.57% and 20.21%, respectively. However, the 13.72%

U

savings in TOC has not bring the expected economic benefits in short time due to the

N

addition of preheaters. The TAC of HAD-FP only decreases by 7.34% with the payback

A

period of 5 years, thus simply preheating the feed streams cannot achieve obvious improvement in economic performances. Due to the applying of VRHP, large amounts

M

of heat from the compressed vapor streams could be used to heat the bottom streams.

ED

So the HAD- VRHP1 process can save 61.22% of energy consumption and 54.79% of TOC. Although the investment cost of the HAD- VRHP1 increases greatly due to the use of the compressors, the 5 years of TAC still decreases by 29.01%. HAD- VRHP1

PT

process also reduce 76.92% of carbon emission. The HAD- VRHP2 process shows better performance in TAC and energy consumption. Compared with the HAD process,

CC E

the HAD- VRHP2 process saves 34.94% of TAC, 59.92% of TOC, and 65.43% of energy consumption. Compared with the HAD process, the HAD-SHRT process saves 38.50% of TAC, 93.30% of CO2 emissions, and 80.92% of energy consumption.

A

Compared with A-DWC, the HAD-SHRT process saves 22.67% of TAC, 58.13% of TOC, and 75.37% of energy consumption. All the heat of the HAD-SHRT process is circulated without any external heat, thus the energy consumption can be decreased greatly to compensate for the cost of compressors. The thermodynamic efficiency of four processes is 9.70%, 8.67%, 10.60%, and 13.51%, respectively. The HAD-SHRT process is found more attractive. It is worth noting that the HAD-SHRT can acquire

more economic benefits in longer payback period. Compared with the HAD process, the HAD-SHRT process can save 48.84% of TAC with the payback period of 8 years. Based on the comparison and analysis, the HAD-SHRT process shows better performance in energy saving and environment. 5 Conclusions In this work, four energy-saving processes are proposed for the heterogeneous

IP T

azeotropic distillation system. All processes are compared in energy saving, economic,

and environmental performances. After the addition of VRHP, the sensible and latent

SC R

heat of the overhead vapor streams are recuperated and supplied to the reboilers. HEN can be used to design the optimal heat integration of the process, which leads to the

sensible heat of the hot streams being supplied to the preheaters. All the heat of HAD-

U

SHRT process can be utilized. Thus, the HAD-SHRT process is most attractive process

N

in energy consumption and thermodynamic efficiency. Compared with the

A

conventional HAD process, the HAD-SHRT process saves 38.50% of TAC, 93.30% of CO2 emissions, and 80.92% of energy consumption. With payback period growth, the

M

economic advantages become more obvious. HAD-SHRT process also shows better

ED

performances than A-DWC proposed in recent study [11]. Although applying the SHRT to the conventional HAD process is an effective way to reduce the energy consumptions and costs, dynamics and control of the HAD-SHRT process is still an

PT

important issue to be addressed in the future.

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Acknowledgements

We are thankful for the assistance from the staff at the Jiangsu Key Laboratory of Advanced Catalytic Materials and Technology from the School of Petrochemical

A

Engineering (Changzhou University).

Reference

A

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A

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IP T

[1] Liang S, Cao Y, Liu X, Li X, Zhao Y, Wang Y, et al. Insight into pressure-swing distillation from azeotropic phenomenon to dynamic control. Chemical Engineering Research and Design 2017;117:318–35. doi:10.1016/j.cherd.2016.10.040. [2] Huang X, Li Z, Tian Y. Process optimization of an industrial acetic acid dehydration progress via heterogeneous azeotropic distillation. Chinese Journal of Chemical Engineering 2017. doi:10.1016/j.cjche.2017.10.030. [3] Mortaheb HR, Kosuge H. Simulation and optimization of heterogeneous azeotropic distillation process with a rate-based model. Chemical Engineering and Processing: Process Intensification 2004;43:317–26. doi:10.1016/S02552701(03)00131-4. [4] Luyben WL. Control of the Heterogeneous Azeotropic n -Butanol/Water Distillation System. Energy & Fuels 2008;22:4249–58. doi:10.1021/ef8004064. [5] Wu Y-C, Chien I-L. Design and Control of Heterogeneous Azeotropic Column System for the Separation of Pyridine and Water. Industrial & Engineering Chemistry Research 2009;48:10564–76. doi:10.1021/ie901231s. [6] Yu H, Ye Q, Xu H, Dai X, Suo X, Li R. Investigation on Thermodynamics in Separation for Ethylene Glycol + Neopentyl Glycol System by Azeotropic Distillation. Journal of Chemical & Engineering Data 2016;61:2330–4. doi:10.1021/acs.jced.5b01044. [7] Dejanović I, Matijašević L, Jansen H, Olujić Ž. Designing a Packed Dividing Wall Column for an Aromatics Processing Plant. Industrial & Engineering Chemistry Research 2011;50:5680–92. doi:10.1021/ie1020206. [8] Yu H, Ye Q, Xu H, Zhang H, Dai X. Design and Control of Dividing-Wall Column for tert -Butanol Dehydration System via Heterogeneous Azeotropic Distillation. Industrial & Engineering Chemistry Research 2015;54:3384–97. doi:10.1021/ie504325g. [9] Shahandeh H, Jafari M, Kasiri N, Ivakpour J. Economic optimization of heat pump-assisted distillation columns in methanol-water separation. Energy 2015;80:496–508. doi:10.1016/j.energy.2014.12.006. [10] Liu Y, Zhai J, Li L, Sun L, Zhai C. Heat pump assisted reactive and azeotropic distillations in dividing wall columns. Chemical Engineering and Processing: Process Intensification 2015;95:289–301. doi:10.1016/j.cep.2015.07.001. [11] Li R, Ye Q, Suo X, Dai X, Yu H, Feng S, et al. Improving the Performance of Heat Pump-Assisted Azeotropic Dividing Wall Distillation. Industrial & Engineering Chemistry Research 2016;55:6454–64. doi:10.1021/acs.iecr.6b00937. [12] Kansha Y, Tsuru N, Sato K, Fushimi C, Tsutsumi A. Self-Heat Recuperation Technology for Energy Saving in Chemical Processes. Industrial & Engineering Chemistry Research 2009;48:7682–6. doi:10.1021/ie9007419. [13] Van Duc Long N, Lee M. A novel NGL (natural gas liquid) recovery process based on self-heat recuperation. Energy 2013;57:663–70. doi:10.1016/j.energy.2013.04.078.

A

CC E

PT

ED

M

A

N

U

SC R

IP T

[14] Christopher CCE, Dutta A, Farooq S, Karimi IA. Process Synthesis and Optimization of Propylene/Propane Separation Using Vapor Recompression and Self-Heat Recuperation. Industrial & Engineering Chemistry Research 2017;56:14557–64. doi:10.1021/acs.iecr.7b03432. [15] Xia H, Ye Q, Feng S, Li R, Suo X. A novel energy-saving pressure swing distillation process based on self-heat recuperation technology. Energy 2017;141:770–81. doi:10.1016/j.energy.2017.09.108. [16] Kansha Y, Tsuru N, Fushimi C, Tsutsumi A. New Design Methodology Based on Self-Heat Recuperation for Production by Azeotropic Distillation. Energy & Fuels 2010;24:6099–102. doi:10.1021/ef100671w. [17] Song C, Liu Q, Ji N, Kansha Y, Tsutsumi A. Optimization of steam methane reforming coupled with pressure swing adsorption hydrogen production process by heat integration. Applied Energy 2015;154:392–401. doi:10.1016/j.apenergy.2015.05.038. [18] Poddar T, Jagannath A, Almansoori A. Use of reactive distillation in biodiesel production: A simulation-based comparison of energy requirements and profitability indicators. Applied Energy 2017;185:985–97. doi:10.1016/j.apenergy.2015.12.054. [19] Lo K-M, Chien I-L. Efficient separation method for tert -butanol dehydration via extractive distillation. Journal of the Taiwan Institute of Chemical Engineers 2017;73:27–36. doi:10.1016/j.jtice.2016.07.040. [20] Luyben WL. Distillation Design and Control Using AspenTM Simulation: Luyben/Distillation Design and Control Using Aspen Simulation. Hoboken, NJ, USA: John Wiley & Sons, Inc.; 2006. doi:10.1002/0471785253. [21] Luyben WL. Design and Control of the Butyl Acetate Process †. Industrial & Engineering Chemistry Research 2011;50:1247–63. doi:10.1021/ie100103r. [22] Waheed MA, Oni AO, Adejuyigbe SB, Adewumi BA, Fadare DA. Performance enhancement of vapor recompression heat pump. Applied Energy 2014;114:69– 79. doi:10.1016/j.apenergy.2013.09.024. [23] Fonyo Z, Benkö N. Comparison of Various Heat Pump Assisted Distillation Configurations. Chemical Engineering Research and Design 1998;76:348–60. doi:10.1205/026387698524776. [24] Kumar V, Kiran B, Jana AK, Samanta AN. A novel multistage vapor recompression reactive distillation system with intermediate reboilers. AIChE Journal 2013;59:761–71. doi:10.1002/aic.13862. [25] Gadalla MA, Olujic Z, Jansens PJ, Jobson M, Smith R. Reducing CO2 emissions and energy consumption of heat-integrated distillation systems. Environ Sci Technol 2005;39:6860–70. [26] Luyben WL. Series versus parallel reboilers in distillation columns. Chemical Engineering Research and Design 2018;133:294–302. doi:10.1016/j.cherd.2018.03.025. [27] Luyben WL. Heat exchanger simulations involving phase changes. Computers & Chemical Engineering 2014;67:133–6. doi:10.1016/j.compchemeng.2014.04.002. [28] You X, Rodriguez-Donis I, Gerbaud V. Reducing process cost and CO2 emissions

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for extractive distillation by double-effect heat integration and mechanical heat pump. Applied Energy 2016;166:128–40. doi:10.1016/j.apenergy.2016.01.028. [29] Waheed MA, Oni AO, Adejuyigbe SB, Adewumi BA, Fadare DA. Performance enhancement of vapor recompression heat pump. Applied Energy 2014;114:69– 79. doi:10.1016/j.apenergy.2013.09.024.

Temperature 70.4℃ 1

Organic phase 312.89 kmol/h 33.1mol% T-butanol 3.5mol%Water 63.5mol% Cyclohexane

RR=0.54

6

U

3

D=0.98m

HADC

N

PDC

1

Decanter

6

A

Aqueous phase 43.56 kmol/h 9.4mol% T-butanol 90.6mol% Water 3.30e-2mol% Cyclohexane

8

Product 50kmol/h 99.8mol% Water 0.2mol% T-butanol

ED

Reboiler duty: 1942 kW Temperature 98.3℃

M

Feed: 25 ℃ 100kmol/h 50mol% T-butanol 50mol% Water

IP T

Cyclohexane makeup Condenser duty: -4081 kW 0.004kmol/h

SC R

Condenser duty: -1614 kW Temperature 80.0℃

D1 93.56 kmol/h 57.7mol% T-butanol 42.3mol% Water 1.54e-2mol% Cyclohexane

D=1.74m

32

Reboiler duty: 4058 kW Temperature 84.5℃

A

CC E

PT

Figure 1. Process flow sheet of conventional HAD with details

Product 50kmol/h 99.8mol% T-butanol 0.2mol% Water

D1 92.3103 kmol/h 58.4mol% T-butanol 41.6mol% Water 1.57e-2mol% Cyclohexane

84.4℃ E1 299.1 kW

E3 450.0 kW

Decanter

Aqueous phase 42.32 kmol/h 9.5mol% T-butanol 90.5mol% Water PDC 3.43e-2mol% Cyclohexane 7

11

Cooler2 -2874kW

70.5℃ 1

67.0℃ Organic phase 197.78 kmol/h 42.6mol% T-butanol 7.2mol% Water 50.2mol% Cyclohexane HADC

79.6℃ E2 59.1 kW

12

32

E5 2393.2 kW

A

CC E

PT

ED

M

N

A

Product 98.3℃ 50kmol/h 99.8mol% Water 0.2mol% T-butanol

U

E4 1324.2kW

Figure 2. Flow sheet of HAD-FP

IP T

80.0℃ Feed: 25 ℃ 1 100kmol/h 50mol% T-butanol 50mol% Water

Cyclohexane makeup 1×10-10kmol/h

SC R

Cooler1 -1333.6kW

Product 84.5℃ 50kmol/h 99.8mol% T-butanol 0.2mol% Water

1.6720 1.6715

NFF=7 NFAQ=13

1.6705

IP T

NFF=6 NFAQ=9

1.6700 1.6695

NFF=7 NFAQ=10

1.6690

NFF=7 NFAQ=12

1.6685 1.6680

NFF=7 NFAQ=11 9

10

11

12

13

14

15

NT1

(a)

U

1.60 1.55

N

NF2=1 NF2=2 NF2=3

A

1.50

1.40

M

TAC/10^6$

1.45

1.35

ED

1.30 1.25

29

PT

1.20 1.15

SC R

TAC/10^6$

1.6710

30

31

32

33

34

35

NT2

A

CC E

(b) Figure 3. The influence of NT1 (a) and NT2 (b) on the TAC

110 100

QFH

QCW 80

CCC HCC

70 60

IP T

Temperature (℃ )

90

50 40 30

10

0

2000

4000

6000

8000

SC R

20 10000

Heat flow (kW)

C1 80.0℃

PT

CC E A

E2 58.8 kW

C2

Makeup V2 Cyclohexane

RC=3.65 357.56 kW

69.9℃

1 E3 447 kW

V1

PDC 7

E1 299.1 kW

Cooler2 -835kW

Decanter 1

ED

Feed: 25 ℃ 1.20 bar 100kmol/h 50mol% T-butanol 50mol% Water

RC=2.86 148.81 kW

M

A

N

U

Figure 4. Temperature-Heat flow (T-H) diagram of the HAD-FP

11

12

Cooler1 -158kW

HADC

S2 117.6℃

S1 127.6℃

E4 1324.3kW

S3 103.3℃ Product 98.3℃ 50kmol/h 99.8mol% Water 0.2mol% T-butanol

Figure 5. Flow sheet of HAD-VRHP 1

S4 89.5℃

32

E5 2378.1 kW

Product 84.5℃ 50kmol/h 99.8mol% T-butanol 0.2mol% Water

175

8.5

Compressor duty（ kW） Temperature difference (℃)

8.0

6.5

160

6.0 155

5.5 5.0

150

4.5 145 4.0 140 2.7

2.8

2.9

3.0

3.1

3.2

3.3

3.5 3.4

Compression ratio

U

(a)

7.5

N

400

Compressor duty（ kW） Temperature difference (℃)

7.0

A M

380

370

PT

350

340 3.4

3.5

3.6

3.7

3.8

6.5 6.0 5.5

ED

Compressor duty（ kW）

390

360

IP T

7.0

SC R

165

5.0 4.5

Temperature difference(℃ )

Compressor duty（ kW）

7.5

Temperature difference(℃ )

170

4.0

3.9

4.0

4.1

3.5 4.2

CC E

Compression ratio

A

(b) Figure 6. The influence of compression ratio of C1 (a) and C2 (b) on compressor duty and temperature difference;

140

QCW

Heat recovery zone （ Qrecovery）

120

HCC

100

80

CCC

IP T

Temperature (℃ )

QFH

60

SC R

40

20 0

1000

2000

3000

4000

5000

Heat flow (kW)

88.4℃ 25.2 kW C1

80℃

127.6℃ 2.39 bar

CC E

E1 299.1 kW

PT

PDC 7

A

E2 58.8 kW

11

12

E4 1324.3kW

Cooler1 -157.8kW

S3 103.3℃ Product 98.3℃ 50kmol/h 99.8mol% Water 0.2mol% T-butanol

Makeup Cyclohexane

99.5 kW C2

123.46 kW

1

Cooler2 -835kW 83.7℃

69.9 ℃

ED

Feed: 25 ℃ 100kmol/h 50mol% T-butanol 50mol% Water

M

A

N

U

Figure 7. Temperature- Heat flow (T-H) diagram of the HAD-VRHP 1

1

Decanter

258.21 kW 67℃

E3 447 kW

HADC

32

117.6℃ 2.52 bar E5 2378.1 kW

S4 89.5℃ Product 84.5℃ 50kmol/h 99.8mol% T-butanol 0.2mol% Water

Figure 8. The improved process with compressor inlet preheating

9.52

2.45

9.46

2.40

9.44

2.35

9.42

84

85

86

87

88

89

90

91

92

93

SC R

9.48

94

Tp（ ℃）

U

(a) 2.80

9.05

N

Compression ratio TAC

A

2.70

2.60 2.55

PT

2.50

CC E

79

80

81

82

83

9.00 8.95 8.90

M

2.65

8.85 8.80

ED

compression ratio

2.75

2.45

9.40 95

TAC/10^5$

compression ratio

2.50

2.30

(b)

9.50

IP T

Compression ratio TAC

2.55

TAC/10^5$

2.60

8.75 8.70

84

85

86

87

88

89

8.65 90

Tp（ ℃）

A

Figure 9. The effects of TP on the compression ratio and TAC in PDC (a) and HADC (b)

140

Heat recovery zone

QCW QFH

HCC

100 80

CCC 60

IP T

Temperature (℃ )

120

40

0

0

1000

2000

3000

SC R

20

4000

5000

Heat flow (kW)

A

N

U

Figure 10. Temperature- Heat flow (T-H) diagram of the HAD-VRHP2

PDC-vapor

M

.

88.4 ℃

ED

E3

1324.3 kW 103.3 ℃

99.5 kW

99.6 ℃

83.7 ℃

E4

84.9 ℃

58.81 kW E5

80 ℃

42.8 ℃

30 ℃ 117.6 ℃

C2

PT

HADC-vapor

E6

98.3 ℃

Cooling water

127.6 ℃

C1 25.2 kW

PDC-bottom

Cooler 1 Cooler 2 Cooler 3 33.15 kW 12.65 kW 286.8 kW 25 ℃

101.4 kW

2378.1 kW 89.5 ℃

E2

E7

447 kW

82.7 ℃

51.4 ℃ 197.7 kW

E8

CC E

HADC-bottom 84.5 ℃

98.3 ℃

E7

A

E2

66.5 ℃

E1

25 ℃

69.9 ℃

E6

79.6 ℃

67 ℃

84.4 ℃ 80 ℃

E3

84.4 ℃

83.7 ℃

81.4 ℃

E4

84.5 ℃ 88.4 ℃

30 ℃ 30 ℃

E1

E5 E8

Figure 11. The optimal heat exchanger networks design

30 ℃

30 ℃

PDC-reboiler HDC-reboiler PDC-vapor preheat

PDC- feed HADC-vapor preheat PDC-aqueous feed HADC-organic feed

A ED

PT

CC E

IP T

SC R

U

N

A

M

82.7 ℃ 89.5 ℃

99.6℃ Cooler3 -286.8 kW

84.9 ℃ Cooler1 -33.15 kW

88.4℃

E3 25.2 kW

80℃

E6 99.5 kW

69.9 ℃

80 ℃

C1

30.0℃ 51.4℃ C2

Decanter

1

1 127.6℃ 2.39 bar

PDC

258.21 kW 67℃

30.0℃ E8 447 kW

84.4 ℃ 7

HADC

E2 101.4 kW

117.6℃ 2.52 bar 11

103.3 ℃

12

32

E4 1324.3 kW

Product 30℃ 50kmol/h 99.8mol% T-butanol 0.2mol% Water

98.3 ℃

42.8℃

Cooler2 -12.65 kW

84.5 ℃

Product 30℃ 50kmol/h 99.8mol% Water 0.2mol% T-butanol

79.6 ℃

A

CC E

PT

ED

M

A

Figure 12. Flow sheet of HAD-SHRT

30.0℃

E7 2378.1 kW

N

E5 58.81 kW

89.5 ℃

SC R

66.5℃

U

E1 197.7 kW

IP T

123.46 kW

Feed: 25 ℃ 100kmol/h 50mol% T-butanol 50mol% Water

Makeup Cyclohexane

83.7℃

Reboiler of PDC: total Q =1324.3 kW 103.3 ℃

119.18 kmol/h 1.01 bar 80.0 ℃

Cold side boiling liquid at 98.3℃

123.46 kW

Superheater E3

88.4 ℃

C1

127.6 ℃ 2.39 bar

25.2 kW 7.4 m2 U = 0.2

Desuperheating E4A 77.44 kW 20 m2 U = 0.28

Condensing E4B 103.33 ℃ V/F = 1

1246.86 kW 207 m2 U = 1.2

Reboiler of HADC: total Q = 2378.1 kW 99.6 ℃

238.86 kmol/h 1.01 bar 69.9 ℃

Superheater E6

83.7 ℃

C2

117.6 ℃ 2.52 bar

99.5 kW 32.2 m2 U = 0.2

Cold side boiling liquid at 84.5℃

258.21 kW

Desuperheating E7A

171.2 kW 30.5 m2 U = 0.28

SC R

84.9 ℃

95.5 ℃ V/F = 1

CC E

PT

ED

M

A

N

U

Figure 13. Desuperheating and condensing sections

A

103.3 ℃ V/F = 0

IP T

99.6 ℃

Condensing E7B

2206.9 kW 240.8 m2 U = 1.2

89.5 ℃ V/F = 0

140

Heat recovery zone

QCW

100

HCC

80

CCC

60

IP T

Temperature (℃ )

120

40 20

0

1000

2000

3000

4000

5000

SC R

0

Heat flow (kW)

A

CC E

PT

ED

M

A

N

U

Figure 14. Temperature- Heat flow (T-H) diagram of the HAD-SHRT

140

heat recovery pockets

100

80

60

IP T

Temperature (℃ )

120

40

0

500

1000

1500

2000

2500

Net heat flow (kW)

SC R

QCW 20

A

CC E

PT

ED

M

A

N

U

Figure 15 Grand composition curve (GCC) diagram of the HAD-SHRT

Table1. NRTL Parameters of the System Component i

t-butanol

t-butanol

H2O

H2O

cyclohexane

cyclohexane

Bij (J/mol)

203.4185

559.5207

−1066.9764

Bji (J/mol)

-1372.3835

245.1068

4954.897

Cij

0.3

0.47

0.2

A

CC E

PT

ED

M

A

N

U

SC R

IP T

Component j

Table2. Basis of Economics Condensers costs Heat transfer coefficient (U): U = 0.852 kW/K · m2 Q

Heat transfer coefficient (A): A =

U∆T

0.65

Capital cost = 7296A Reboilers costs Heat transfer coefficient (U): U = 0.568 kW/K · m2 Q U∆T

Capital cost = 7296A0.65 Self-heat exchangers Heat transfer coefficient (U): Liquid -liquid Liquid - vapor Vapor-vaporizing liquid Vapor - vapor Condensing vapor-vaporizing liquid

SC R

U

Q U∆T

N

Heat transfer coefficient (A): A =

U = 0.57 kW/K · m2 U = 0.20 kW/K · m2 U = 0.28 kW/K · m2 U = 0.17 kW/K · m2 U = 1.2 kW/K · m2

1293 280

ED

M

A

Capital cost = 7296A0.65 Column vessel costs Capital cost = 17640D1.066 L0.802 Column length (L): L = (NT − 1) × 0.61 + 6 Column diameter (D): Aspen tray sizing Compressor costs Capital cost =

× 1264.75 × (hp)0.82

CC E

PT

hp: the horse power of compressor Annual steam costs Steam cost = CS× Q × 8000 × 3600 Low steam: CS = $ 7.78 / GJ Annual cooling water cost Q

C Cooling water cost = 0.03 × ∆T×4.183×1000 × 8000 × 3600

Electricity costs

hp

3.6

A

Electricity cost = 16.8 × 0.72 × 8000 × 1000 TAC =

total captical cost paybackperiod

+ annuanl energy cost

Payback period: 5 years

IP T

Heat transfer coefficient (A): A =

Table3. Stream Energy Loads of HAD-SHRT Stream

Stream type

Target Tem(℃)

Supply

Duty（kW）

Tem(℃)

88.4

80

PDC-bottom

Hot

98.3

30

HADC-vapor

Hot

83.7

30

HADC-bottom

Hot

84.5

30

PDC-reboiler

cold

81.4

HDC-reboiler

cold

84.4

PDC-vapor preheat

cold

80

PDC- feed

cold

25

cold

HADC-organic feed

cold

PT CC E A

30

345.61

2959.65 197.7

98.3

1324.3

84.5

2378.1

88.4

25.2

84.4

299.1

83.7

99.5

79.6

58.81

67

447

U N

30

ED

PDC-aqueous feed

69.9

A

cold

M

HADC-vapor preheat

1461.65

IP T

Hot

SC R

PDC-vapor

Item

HAD-SHRT process

E1

Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2) Heat duty (kW) Heat-transfer areas (m2) Overall heat-transfer coefficients (kW/K·m2)

E2

E3

SC R

E4A

U

E4B

A

N

E5

M

E6

ED

E7A

CC E

PT

E7B

E8

197.7 34.30 0.57 101.4 18.59 0.57 25.2 7.4 0.2 77.44 20 0.28 1246.86 207 1.2 58.81 6.73 0.57 99.5 32.2 0.2 171.2 30.5 0.28 2206.9 240.8 1.2 447 42.76 0.57 33.15 0.65 0.852 12.65 1.1 0.852 286.8 19.93 0.852

IP T

Table 4. Basic parameters of the HAD-SHRT process

Cooler 1

A

Cooler 2

Cooler 3

compressor

Compressor duty (kW)

123.46

Comp 2

Compressor duty (kW)

258.21

A

CC E

PT

ED

M

A

N

U

SC R

IP T

Comp 1

Table 5. Comparison of All the Processes Parameter

HAD

A-DWC

HAD-FP

HAD- VRHP1

HAD- VRHP2

HAD-

SHRT 0.98/1.

Diameter (m)

1.65

0.82/1.28

0.82/1.27

0.82/1.27

0.82/1.27

11/32

8/32

2/6/31

12/32

11/32

11/32

Condenser duty (kW)

5694

4296

4476.8

1263.3

992.8

Reboiler duty (kW)

6000

4650

3717.4

/

/

Compressor pressure ratio

/

/

/

2.86/3.65

Compressor duty (kW)

/

/

/

148.8/357.6

Energy consumption (kW)

6000

4650

4526

2327

Thermodynamic

7.88

9.56

9.70

8.67

CO2 emissions (kg/h)

1044

716

833

TCC (105 US$)

6.88

7.52

TOC (105 US$)

12.10

9.22

2.36/2.49

123.5/258.2

123.5/258.2

2074.0

1145.1

U

2.36/2.49

10.60

13.51

241

240

70

10.25

20.50

19.57

22.14

10.44

5.47

4.85

3.86

10.72

12.49

9.57

8.77

8.29

10.16

11.72

8.03

7.30

6.63

13.48

PT

5 years

CC E

8 years

M

ED

TAC (105 US$)

12.96

A

(%)

A

/

N

efficiency

332.6

SC R

Stage number

IP T

74