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Systems Design, Modeling, and Thermo-Economic Analysis of Azeotropic Distillation Processes for Organic Waste Treatment and Recovery in Nylon Plants Hua Zhou, Yintian Cai, and Fengqi You Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b00275 • Publication Date (Web): 23 Mar 2018 Downloaded from http://pubs.acs.org on March 29, 2018

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Industrial & Engineering Chemistry Research

Systems Design, Modeling, and Thermo-Economic Analysis of Azeotropic Distillation Processes for Organic Waste Treatment and Recovery in Nylon Plants Hua Zhoua, b, Yintian Caia, Fengqi Youb,* a

Department of Chemical and Biochemical Engineering, National Engineering Laboratory

for Green Chemical Productions of Alcohols-Ethers-Esters, College of Chemistry & Chemical Engineering, Xiamen University, Xiamen 361005, PR China b

Robert Frederick Smith School of Chemical Engineering and Biomolecular Engineering, Cornell University, Ithaca, New York 14853, USA Submitted to 2017 European Symposium on Computer-Aided Process Engineering Special Issue of I&EC Research

Abstract Nylon-6 and nylon-6,6 processes produce considerable amount of organic waste (known as light oil) consisting of n-pentanol, cyclohexanone, and cyclohexene oxide, which are difficult to separate and recover. This paper proposes six novel process designs to separate the light oil into three value-added products based on azeotropic distillation using water as an entrainer. These azeotropic distillation process designs take into account direct sequence, indirect sequence, thermal coupled column, and three types of dividing wall columns (dividing wall at the top, bottom, and middle of columns, respectively) for entrainer recovery. A conventional distillation process design for separation of the same light oil is also modeled and analyzed for comparison. High-fidelity process simulations are performed for each of the seven process designs in Aspen Plus. We further conduct exergy analyses and techno-economic analyses to evaluate and compare the exergy efficiencies and economic performances of these seven process designs. The results indicate that the proposed azeotropic distillation process design with dividing wall at the middle of the column has the best performance in terms of both exergy efficiency and total annual cost.

*

Correspondence should be addressed. Email: [email protected].

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Introduction The oxidation products of cyclohexane are important intermediates in the manufacture of nylon-6 and nylon-6,6 polymers. However, due to the low selectivity of oxidation reaction, the corresponding industrial process produces considerable amount of wastes and pollutants.1 For example, this process produces large amounts of undesired by-product called light oil, which mainly contains n-pentanol, cyclohexanone, and cyclohexene oxide. The resulting light oil may lead to significant environmental impacts if it is discharged unsuitably. Waste treatment and utilization has been the focus in the chemical industry, which can not only generate value-added products but also mitigate environmental pollution.2 Additionally, high-purity n-pentanol, cyclohexanone, and cyclohexene oxide are important pharmaceutical intermediates or raw materials. Therefore, it is of significant importance to develop cost-effective and energy efficient processes to separate the light oil into value-added products, which are profitable and environmentally benign. Table 1. Properties of components in ternary mixture3. Items Name CAS Number Chemical formula Molar mass (g·mol−1) Density (g·cm−3) Melting point (°C) Boiling point (°C)

n-pentanol 71-41-0 C5H12O 88.15 0.811 -78 137 to 139

Ingredients cyclohexanone 108-94-1 C6H10O 98.15 0.948 -47 155.6

cyclohexene oxide 286-20-4 C6H10O 98.15 0.966 -40 130

Distillation is one of the most widely used separation techniques used in chemical industries.4 However, according to the chemical properties shown in Table 1,3 it is found that the mixture of n-pentanol, cyclohexanone, and cyclohexene oxide has close boiling-points of the species. This poses a great challenge to recover high purity products from the light oil through the conventional distillation strategy. The main reason is that the conventional distillation process often leads to considerable capital investments and high energy consumption for separating close-boiling-point mixtures.5 To reduce the cost and improve the energy efficiency, energy-efficient separation process with integrated and hybrid technologies should be considered.6 There are many recent publications actively addressing energy-efficient distillation column designs,7-8 analysis,9 and


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operations.10-11 Nowadays, various advanced technologies have been applied for process intensification to improve energy efficiency of distillation, such as heat integrated distillation,12-16 heat pump assisted distillation,9, 17-18 multi-effect distillation,19-20 distillation with dividing-wall column,21-24 and so on. Despite recent advances in distillation technologies, it remains a challenge to separate the light oil using conventional distillation sequences with process intensification.25 Specialized distillation processes or novel separation technologies should be developed and employed for separation of the light oil. These separation technologies such as extractive distillation,26 azeotropic distillation,27 reactive distillation,28 and pressure-swing distillation,29 have been used to separate close-to-boiling-point mixture or binary azeotropes. The formation of constant-boiling mixtures can be explained as a function of the nonideality of the mixture. Deviation from ideality can be attributed to the effects of hydrogen bonds or internal pressure.30 Hydrogen bonds, which is a more important factor, can frequently occur between organic molecules with hydroxyl group (-OH) and intramolecularly with a highly electronegative atom such as nitrogen (-N), oxygen (-O), or fluorine (-F).31 There is a hydroxyl group (-OH) in n-pentanol, so there is a potential opportunity to separate the light oil into value-added products by azeotropic distillation. In addition, the separation process with process integration could lead to 30% energy reduction, compared to the corresponding conventional approach.32 Therefore, the objective of this paper is to develop novel separation methods to recover the three value-added products (n-pentanol, cyclohexanone, and cyclohexene oxide) from light oil. In this paper, we propose six novel process designs to separate the light oil into three value-added products based on azeotropic distillation using water as an entrainer. These novel azeotropic distillation process designs take into account direct sequence, indirect sequence, thermal coupled column, and three types of dividing wall columns (DWC) for entrainer recovery. High-fidelity process simulations are performed for each of the proposed process designs. A conventional distillation process design for separation of the same light oil is also modeled and analyzed for comparison. We further conduct exergy analyses and techno-economic analyses to evaluate and compare the exergy efficiencies and economic performances of these seven process design alternatives. The novel contributions of this work are summarized below:

Six novel azeotropic distillation process designs for separation of the light oil using water as the entrainer for waste recovery and environmental protection;



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A comprehensive investigation of energy saving strategies for azeotropic distillation including thermally coupled column and DWC;

Systematic comparisons of conventional distillation and different azeotropic distillation process designs for separation of light oil from the exergetic and economic perspectives. The rest of this article is organized as follows. The process designs and description of

conventional distillation and azeotropic distillation are presented in the following section. The Process Evaluation and Analysis section presents the results of process simulation, exergy analysis, and techno-economic analysis. The conclusion is provided at the end.

Process design, simulation and optimization Azeotropic distillation is employed to separate the light oil into value-added products. Because there is no azeotrope in the light oil, it is necessary to find an entrainer which could form an azeotrope with an ingredient in the light oil. In light of the knowledge of hydrogen bonds,30 water is a suitable entrainer for the separation. Furthermore, it is economical and environmentally friendly to select water as the entrainer. We propose six novel azeotropic distillation process designs. Besides, a conventional distillation process design is considered for systematic comparisons. Table 2. Binary parameters of components in the mixture. i

n-pentanol

n-pentanol

cyclohexanone

j

cyclohexanone

water

cyclohexene oxide

water

water

Aij

0

-5.221

0

0

0

Aji

0

2.3356

0

0

0

Bij

-427.508

2254.9058

-273.24458

280.3563

108.7611

Bji

890.0146

831.1645

383.245852

841.9608

1532.8239

Cij

0.3

0.39

0.3

0.3

0.3

cyclohexanone cyclohexene oxide

In this study, all process simulation models are built in Aspen Plus33 to systematically investigate the performance of different separation designs, and NRTL model is used to predict the thermodynamic properties of the vapor and liquid mixtures and binary NRTL parameters are Aspen Plus built-in as shown in Table 2. The azeotrope of water and n-pentanol can be seen in the


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vapor-liquid equilibrium (VLE) diagram as shown in Figure 1.

Figure 1. VLE diagram of water and n-pentanol.

Design specifications The capacity of the light oil separation process is set at 500kg/h according to the requirement of the plant.34 Based on the industrial data from Yueyang Branch of Sinopec Corp., the composition of the feedstock is given as follows: n-pentanol 35.4%, cyclohexanone 31.8%, and cyclohexene oxide 32.8% (mole fraction). According to the requirements of environmental protection and product quality, the design specifications for the separation process should satisfy the following constraints: 1) The recovery ratio of each product should be above 92.5%. 2) The purify of n-pentanol should be greater than 95 wt% and that of cyclohexanone should be above 95 wt%. 3) The recovery ratio of the entrainer should be more than 98% for the azeotropic distillation process. For separation of the light oil, a conventional distillation separation process design consists of two columns and three products. The azeotropic distillation process design can be divided into two parts: Part I combines an azeotropic distillation column with a cyclohexanone and cyclohexene oxide recovery column; and Part II is the entrainer and n-pentanol recovery system. The conventional separation process, as well as the novel separation schemes, are introduced in the following subsections. All the seven separation process designs are summarized in



Table 3.


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Table 3 Categories of different separation process design. No.

Scheme

1

Conventional distillation Azeotropic distillation with direct sequence Azeotropic distillation with indirect sequence Azeotropic distillation with thermally coupled Azeotropic distillation integrated DWC with the wall at upper part Azeotropic distillation integrated DWC with the wall at lower part Azeotropic distillation integrated DWC with the wall at middle part

2 3 4 5 6 7

Number of columns 2 4

Entrainer used N Y

Number of Number of condensers reboilers 2 2 4 4

4

Y

4

4

4

Y

3

3

3

Y

4

3

3

Y

3

4

3

Y

3

3

Design 1: Conventional distillation sequence To separate the ternary mixture into three products, at least two distillation columns are needed. The conventional distillation process design for the light oil is illustrated in Figure 2. Light oil (S1) is fed into the column C1, and the top stream (S3) of C1 is used as the feedstock of the subsequent column C2. The bottom stream (S2) of C1 is the cyclohexanone product. After that, the top stream (S4) and the bottom stream (S5) of C2 are cyclohexene oxide and n-pentanol by the separation of column C2, respectively. There are 30 trays in Column C1, and 70 trays in the column C2, excluding the condenser because of its total condenser. The feed position in column C1 is 14th tray, and the feed position in Column C2 is 51st tray. This process is modeled by ASPEN Plus.35 The information of feedstock, products, and columns is reported in Table 4. From Table 4, we can see that high purify products can be obtained by the conventional distillation sequences. However, this design leads to relatively high capital and operating costs because a large number of trays are required, and the reflux ratio of the column C2 (RR = 28.0) is very high.

Table 4. Conventional distillation sequence: physical specification and simulation results. Items Total number of tray Feed flowrate (kg/hr) Feed composition Feed tray position Reflux ratio

Column C1 30 500 0.354/0.318/0.328 a 14 5.5


Column C2 70 --51 28.0


S2 (cyclohexanone, mole fraction) 0.992 -S4 (cyclohexene oxide, mole fraction) -0.969 S5 (n-pentanol, mole fraction) -0.972 Condenser duty (kW) 258 483 Reboiler duty (kW) 273 483 a The mole fractions of n-pentanol, cyclohexanone, and cyclohexene oxide, respectively.

Figure 2. Conventional distillation consequence with two distillation columns.

Design 2: Direct azeotropic distillation sequences To alter the relative volatility of the two species in the feedstock, azeotropic distillation is adopted to separate the light oil using water as an entrainer. From the azeotropic distillation column, an azeotrope (including water, n-pentanol, and part of cyclohexanone) and a mixture of cyclohexanone and cyclohexene oxide can be obtained. There are two sequences for azeotrope separation, namely, direct and indirect separation sequences, according to the literature.36 Figure 4 presents the flowsheet of direct azeotropic distillation process. The light oil (S1) and entrainer (S2) are first introduced into the azeotropic distillation column C1. The mixture of cyclohexanone and cyclohexene oxide (S3) from the bottom of column C1 is then fed into column C2, where it is separated into cyclohexanone (S8) and cyclohexene oxide (S7). After being cooled in a heat


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exchanger HX1, the azeotrope from the top of column C1 enters a decanter to remove part of water. The resulting stream (S6) is sent to column C3 as feedstock. In column C3, water is recovered from the azeotrope as steam (S9). The mixture of n-pentanol and cyclohexanone (S10) from the bottom of column C3 is then separated into n-pentanol (S11) and cyclohexanone (S12) in column C4. Water recovered from the decanter (stream S9) and column C3 (stream S13) can be transported to the vessel and heated up as the entrainer for column C1. In the direct separation sequence, water is first removed from the azeotrope in column C3, and n-pentanol and cyclohexanone are then separated in another column C4. The role of process section in the red dash box of Figure 4 is used to separate the light oil into cyclohexanone, cyclohexene oxide, and azeotrope. This section is the same as those in the direct and indirect azeotropic distillation process designs. In order to understand the process better, a ternary diagram of n-pentanol/water/cyclohexanone is presented in Figure 3.

Figure 3. Ternary diagram of n-pentanol/water/cyclohexanone.



Figure 4. Flowsheet of the direct azeotropic distillation process design. To optimize this separation process design, major design parameters, such as flowrate of the entrainer and recovery of n-pentanol, are further investigated through detailed process simulation model in Aspen Plus. Analysis of flowrate of the entrainer The VLE diagram of the water and n-pentanol in Figure 1 suggests that there is a suitable composition for the mixture of water and n-pentanol to form the azeotrope. In addition, the variation of entrainer flowrate has influences on the operating cost and product recovery ratio. Thus, the influences of the variation of entrainer flowrate are investigated, and the result is presented in Figure 5. The horizontal axis in Figure 5 is the flowrate of the entrainer. The left vertical axis and right vertical axis in Figure 5 are mass fraction of water and mass fraction of C6 components (including cyclohexanone and cyclohexene oxide), respectively. To reduce the loss of cyclohexanone and cyclohexene oxide, as well as energy consumption, the optimal flowrate of the entrainer is determined as 242kg/hr based on Figure 5. In addition, the parameters of columns C1 and C2 are given in Table 5. It is noteworthy that the parameters of columns C1 and C2 in Table 5 are used in the following separation schemes (indirect azeotropic distillation, thermal couple distillation, and DWC separation process). Table 5 shows that the reflux ratios of the azeotropic distillation for columns C1 and C2 are 2.4 and 3.1, respectively.


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Column C1 Total number of tray 15 Feed flowrate (kg/hr) 500 Feed composition -(mole fraction) S1 (light oil) 0.354/0.318/0.328/0a S2 (entrainer, water) 0/0/0/1a Feed tray position -S1 1 S2 9 Reflux ratio 2.4 Product composition (mole fraction) S3 0.010/0.426/0.510/0.017a S4 0.119/0.014/trace/0.867a S7 -S8 -Condenser duty 456 (kW) Reboiler duty (kW) 442

Column C2 30 318.511 ---15 --3.1

--0.019/0.002/0.947/0.032a trace/0.992/0.008/tracea 73 72

Table 5. Simulation results of columns C1 and C2 in the azeotropic distillation process. The mole fractions of n-pentanol, cyclohexanone, cyclohexene oxide, and water, respectively. mass fraction of water in the bottom stream of column C1 mass fraction of organic component in the top stream of column C1

0.012

0.22

0.010

0.20

0.008 0.18 0.006 0.16 0.004 0.14

0.002

0.12

0.000 234

236

238

240

242

244

246

248

flowrate of entrainer (kg/hr)

Figure 5. Effects of the variation of entrainer flowrate.


250

252

mass fraction of organic component (C6)

a

mass fraction of water

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60



Optimization of n-pentanol and entrainer recovery process To maximize the profit of the recovery process (corresponding to maximizing the recovery ratio of the n-pentanol and entrainer and minimizing the energy consumption), tray number, feedstock position, reflux ratio, and distillation ratio the columns C3 and C4 are investigated. Figure 6 shows entrainer and n-pentanol recovery ratio over different tray number (from 5 to 15) of column C3 in the direct distillation sequence. The horizontal axis represents tray number and the vertical axis represents recovery ratio of entrainer and n-pentanol. The graph indicates that the recovery ratio of entrainer and n-pentanol varies sharply under 6 trays, and then keeps stable over 6 trays. Hence, the required number of trays is 6 for column C3. Figure 7 illustrates entrainer and n-pentanol recovery ratio with different reflux ratio of column C3 in the direct distillation sequence. Although the analysis results show that the optimal reflux ratio is 1.49, we suggest the reflux ratio as 0.93 because the design specifications can be achieved in this condition.

Figure 6. Relationship of tray number of C3 with mass fraction of n-pentanol and cyclohexanone in the direct azeotropic separation sequence.


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Figure 7. Relationship of reflux ratio of C3 with recovery ratio of entrainer and n-pentanol in the direct azeotropic distillation sequence.

Figure 8. Relationship of distillate ratio of C3 with recovery ratio of entrainer and n-pentanol in the direct azeotropic separation sequence.

Figure 8 presents the relationship of distillate ratio of column C3 with the recovery ratio of entrainer and n-pentanol in the direct azeotropic sequence. The analysis suggests that the optimal distillate ratio for column C3 should be 0.44. To satisfy the design specification (2) on product quality, a sensitivity analysis for tray number of column C4 is performed. Figure 9 shows the result of sensitivity analysis of the tray number ranging from 25 to 40. The results suggest that compositions of the top stream and bottom stream can be satisfied when the tray number is 39.



Feed tray position is analyzed from 2 to 38, and the result is given in Figure 10. The optimal feed tray position of column C4 is the 25th tray. At the same time, the reflux ratio of column C4 is optimized to satisfy the quality of the product and minimize the energy consumption. The relationship of different reflux ratio with the compositions of top stream and bottom stream is provided in Figure 11. It can be seen that a suitable reflux ratio of column C4 is 10.

Figure 9. Effects of changing the tray number of column C4 on the mass fraction of n-pentanol and cyclohexanone in the direct azeotropic separation sequence.

Figure 10. Result of the feed tray position versus mass fraction of the product of C4 in direct azeotropic distillation sequence.


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Figure 11. Effects of the reflux ratio on the quality of the product of column C4 in the direct azeotropic sequence.

Furthermore, the ratio of the distillate to feed ranging from 0.90 to 0.92 is analyzed, and the result is given in Figure 12. After balancing energy consumption and product purity, the ratio of the distillate to the feed is set as 0.914. The resulting optimal design parameters of columns C3 and C4 are summarized in Table 6. In addition, all information of the direct azeotropic distillation is listed along with the flowsheet in Figure 13. Table 6. Design and operating parameters of columns C3 and C4 in the direct azeotropic distillation Total number of tray Feed flowrate (kg/hr) Feed composition Feed tray position Reflux ratio S9 S10 S11 S12 Condenser duty (kW) Reboiler duty (kW) a

Column 3 (C3) 6 201.887 0.547/0.054/trace/0.399a 2 0.93 0.121/0.013/0/0.867a 0.884/0.083/0/0.033a --32 31

Column 4 (C4) 39 161.364 0.884/0.083/0/0.033a 25 10 --0.963/0.001/0/0.036a 0.040/0.960/0/tracea 236 236

process.

The mole fractions of n-pentanol, cyclohexanone, cyclohexene oxide, and water, respectively.



Figure 12. Relationship of the purity of the products with the ratio of distillate to feed in direct azeotropic separation sequence.

Figure 13. Flowsheet of the optimized design for the direct azeotropic distillation scheme.

Design 3: Indirect azeotropic distillation sequences Besides direct azeotropic distillation, another distillation sequence for azeotropic separation is


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indirect azeotropic distillation. In the indirect sequence process, cyclohexanone is first recovered from the azeotrope in column C3 as stream S10 as shown in Figure 14. The mixture of n-pentanol and water (S9) from the bottom of column C3 is then fed into column C4, where water is recovered from the mixture, and high-purity n-pentanol is obtained from the bottom. By using the same optimization method as the direct azeotropic distillation process, the indirect azeotropic distillation process is developed. Figure S1 in the Supporting information gives the recovery ratio for n-pentanol and cyclohexanone with different tray numbers of column C3 in the indirect azeotropic distillation sequence. Depending on the information of Figure S1, 31 trays are required in column C3. The optimal feed tray location should be 18th tray according to the sensitivity analysis of Figure S2. The reflux ratio is further set as 7.4 to achieve the desired product recovery as shown from Figure S3. In addition, different distillate ratio for column C3 is considered, and the results are provided in Figure S4, which indicates that the optimal distillate ratio of column C3 is 0.9475. Furthermore, the reflux ratio and distillate ratio of column C4 are optimized as 1.8 and 0.44 as shown in Figure S5 and Figure S6, respectively. According to the above information, the whole flowsheet with detailed information is presented in Figure 14. The optimal design and operating parameters of columns C3 and C4 are summarized in Table 7.

Table 7. Design and operating parameters of the column C3 and C4 in the indirect azeotropic distillation process.

Total number of tray Feed flowrate (kg/hr) Feed composition Feed tray position Reflux ratio S9 S10 S11 S12 Condenser duty (kW) Reboiler duty (kW) a

Column C3 31 201.887 0.547/0.054/trace/0.399a 18 7.4 0.578/trace/0/0.422a 0.008/0.992/0/tracea --358.8 355.8

Column C4 15 184.485 0.578/trace/0/0.422a 2 1.8 --0.106/trace/0/0.894a 0.980/0.001/0/0.019a 46.6 51.7

The mole fractions of n-pentanol, cyclohexanone, cyclohexene oxide, and water, respectively.



Figure 14. Flowsheet with information for the indirect azeotropic distillation.

Design 4: Thermally coupled distillation sequences In both direct and indirect azeotropic distillation processes, there are two columns with four heat exchangers (two condensers and two reboilers) for each recovery system to obtain the entrainer and products. To reduce the capital cost, we can intensify the process by heat integration such as thermally coupled column. There are generally three types of thermally coupled columns: side rectifier, side stripper, and fully coupled. The fully coupled configuration is often named as the Petlyuk configuration.32 It has been reported in the literature that the fully coupled configuration results in the lowest energy consumption for separating ternary mixtures into high-purity products.37 Therefore, only the fully thermally coupled distillation sequence is investigated in this study. Compared with the direct and indirect azeotropic distillation sequences, the condenser and reboiler of the column C3 is not required in the fully thermally coupled distillation sequence. The flowsheet of fully thermally coupled configuration distillation process is shown in Figure 15. In this so-called Petlyuk configuration, the vapor and liquid streams leaving column C3 are directly


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connected with column C4. Column C3 performs a sharp split between water and cyclohexanone, whereas n-pentanol is distributed naturally between the top and bottom products. A further separation towards high-purity components takes place in column C4. For fully thermally coupled distillation sequence, side draw tray position, flowrate of side draw, heat duty of condenser, and reflux ratio are important parameters in the main column. The effects of the important parameters are shown in Figure S7 to Figure S11 in the Supporting information, and the optimal results of these parameters are given in Figure 15. From Figure S11 in the Supporting information, we see that the effect of reflux ratio on mass fraction of n-pentanol is stronger than the effect of reflux ratio on mass fraction of cyclohexanone. The reason is that cyclohexanone is the main component of the effluent at the bottom of column C4. However, n-pentanol is the main component of side draw, and the side draw tray is close to the top of column C4. The composition of liquid at the top of the distillation column is easily influenced by the reflux ratio. Hence, side draw flowrate is the most important factor for the mass fraction of n-pentanol in the product of side draw. Variation of the side draw flowrate is further investigated as shown in Figure S12, so the optimal value of the side draw flowrate is determined to be 147.58kg/h. According to above results, the optimal parameters for columns C3 and C4 in the fully thermally coupled azeotropic distillation sequence are provided in the Table 8.

Table 8. Design and operating parameters of the column C3 and C4 in the fully thermally coupled configuration.

Total number of tray Feed flowrate (kg/hr) S6 (kg/hr) S9 (kg/hr) S10 (kg/hr) S7 (kg/hr) S8 (kg/hr) Feed composition S6 S9 S10 S7 S8 Feed tray position

Column C3 25

Column C4 42

201.887 444.853 658.447 ---

---521.917 782.983

0.547/0.054/trace/0.399a 0.975/0.008/0/0.017a 0.870/0.130/0/tracea ---

---0.789/0.010/0/0.201a 0.874/0.126/0/tracea



S6 3 -S9 1 -S10 25 -S7 -4 S8 -30 Reflux ratio -10.3 Product composition S11 -0.110/trace/0/0.890a S12 -0.002/0.998/0/tracea S13 -0.991/0.008/0/0.001a Condenser duty (kW) 0 193.7 Reboiler duty (kW) 0 196.5 a The mole fractions of n-pentanol, cyclohexanone, cyclohexene oxide, and water, respectively.

Figure 15. Flowsheet of azeotropic distillation integrated with thermally coupled column.

DWC distillation sequences To avoid the remixing effects, DWC is one of the promising technologies for a ternary mixture separation. According to the location of the dividing wall, there are three different kinds of DWCs for a three-component separation, namely, partition wall located at the upper part of a column,


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partition wall located at the lower part of a column, and partition wall located in the middle (between the feed and the side draw) of a column. We analyze each of these three options in the following subsections for Designs 5-7. Design 5: Dividing wall at the top of the column In the direct azeotropic distillation process, the bottom stream from the reboiler of column C3 is fed into column C4 as shown in Figure 4. From the viewpoint of heat integration, we can save the reboiler of column C3 to reduce the capital cost. That is, the bottom stream of column C3 can be directly used as the feedstock of column C4. Meanwhile, a vapor stream from column C4 is extracted back to column C3. Further integration and cost saving can be achieved, if the two columns are integrated into a single shell. This alternative scheme is a DWC, and the partition wall is located at the upper part of the column as shown in Figure 16.

Figure 16. Flowsheet of azeotropic distillation integrated DWC with partition wall at the upper part of the column.



Figure 17. Effect of feed tray position of DWC C3 with the wall at the upper part of column.

Similarly, sensitivity analysis is performed for feedstock tray location, reflux ratio of the right partition in DWC, and heat duty of the condenser in the right partition of DWC. Based on the sensitivity analysis results shown in Figure 17, the optimal feed tray position of column C3 is identified as the 18th tray. Figure 18 shows the simulation results when the reflux ratio of the right partition in the DWC C3 changes from 7.0 to 10.0. It can be seen that the mass fraction of n-pentanol varies obviously with the reflux ratio increases from 7.0 to 8.0. However, the mass fraction of cyclohexanone decreases greatly as the reflux ration increases from 8.0 to 10.0. The optimal reflux ratio for the right partition of the DWC C3 is identified as 8.0 according to the simulation results.


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Figure 18. Simulation result of the variation of reflux ratio of DWC C3 with the wall at the upper part of column.

Furthermore, heat duty of the right condenser in DWC C3 is adjusted to optimize the entrainer recovery system, and the result is depicted in Figure 19. The total heat duty of the two condensers in column C3 changes conspicuously when heat duty of the right condenser in column C3 varies from -180kW to -184kW. In addition, there are different trends for the mass fraction of n-pentanol and the mass fraction of cyclohexanone on the variation of heat duty of the right condenser in column C3. When the heat duty of the right condenser in column C3 is lower than -183.0kW, the mass fraction of cyclohexanone drops down with the decrease of heat duty of the right condenser in column C3. However, the mass fraction of n-pentanol increases with the decrease of heat duty of the right condenser in column C3 when the heat duty is larger than -182.5kW. According to the design specifications, the optimal heat duty of the right condenser in column C3 is determined as -182.5kW.



Figure 19. Effect of heat duty of right condenser in the DWC C3 with the wall at the upper part of column. Based on the above results, we can conclude that there are 42 trays in total for column C3. The partition wall is located at the upper part of the column. In the partition zone of the column C3, both the left part and the right part have the same number of 25 trays. The feed position is on the 18th tray of the left partition. In addition, there are two condensers and one reboiler for the DWC C3. The reflux ratio of the right partition is 8.0, and the reflux ratio of the left partition is 0.25. Detailed information of this scheme is given in Figure 16. Design 6: Dividing wall at the bottom of the column For indirect azeotropic distillation process as shown in Figure 14, column C3 can be integrated with column C4 to eliminate the top condenser of column C3. Further integration can be implemented and the two columns can be integrated into one shell column as shown in Figure 22. In this scheme, the top stream S4 of column C1 is cooled and fed into a decanter to remove part of the entrainer (i.e. water). Then, the stream S6 is introduced into the DWC C3 to recover cyclohexanone and n-pentanol. In this design, sensitivity analysis of distillate ratio and heat duty of condenser for column C3 are investigated. Figure 20 shows the mass fraction of n-pentanol and cyclohexanone, and the heat duty of condenser over different distillate ratio. The horizontal axis represents different distillate ratios, and the vertical axis represents mass fraction of n-pentanol and cyclohexanone, and heat duty of condenser. The graph illustrates the changes of these parameters, as all the factors are highly connected in the DWC. In addition, variation of heat duty of condenser is considered, and the


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simulation results are given in Figure 21. According to the simulation results, detailed information of this process is given in Figure 22.

Figure 20. Results of distillate ratio in DWC C3 with the wall at the bottom part of the column.

Figure 21. Results of heat duty variation of condenser DWC C3 with the wall at the bottom part of the column.



Figure 22. Flowsheet of azeotropic distillation integrated DWC with partition wall at the lower part of the column. Design 7: Dividing wall at the middle of the column As shown in Figure 23, the separation of the light oil is implemented by azeotropic distillation using a DWC (C3) with the wall at the middle of the column. The DWC C3 mainly comprises of six compartments, namely, top section that discharges top product, bottom section that provides bottom product, left divided section (pre-fractionator) that receives the feedstock, right divided section that draws a side stream, and the remaining sections that are used for heat exchange (condenser and reboiler). Therefore, three products can be simultaneously obtained by a single DWC with only one condenser and one reboiler. We develop the DWC with the partition wall in the middle of the column to separate the light oil based on the information of fully thermally coupled distillation scheme. There are 4 trays in the top section, 13 trays in the bottom section, 25 trays in the left divided section, and 25 trays in the right divided section of the DWC. The splitting of liquid leaving the top section and the vapor coming from the bottom section are two key parameters for the DWC with the wall in the middle of the column. Thus, their effects for the DWC are investigated and the results are provided in Figure 24 and Figure 25.


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Figure 23. Flowsheet of azeotropic distillation integrated DWC with partition wall in the middle of the column.

Figure 24. Effect of variation of the splitting vapor flowrate for the DWC with partition wall at the middle part of the column.



Figure 25. Effect of variation of the splitting liquid flowrate for the DWC with the partition wall in the middle of the column. According to Figure 24 and Figure 25, we find that the suitable vapor and liquid flowrates are 590 kg/h and 424 kg/h, respectively. The composition of the mixture at each stage of the dividing parts (left and right of the DWC) is given in Figure S13. As shown in Figure S13 (b), the high purity n-pentanol can be obtained from the right part of the dividing zone. According to the simulation results, the location of side draw is set above the second tray of the right part of the dividing zone for producing n-pentanol. Correspondingly, the temperature distribution of dividing wall zones in the DWC is given in Figure S14. Because the feed tray in the left dividing wall is located above the third tray, we can see that the tray temperature varies sharply at the corresponding position from the result of Figure S14. Detailed information of the stream for the DWC with partition wall in the middle of the column is provided in Figure 26.


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Figure 26. Detail information of the stream for the DWC with partition wall at the middle part of the column.

Exergy and technoeconomic analyses To systematically evaluate and compare the proposed process designs, the exergy analysis and techno-economic analysis are performed.38-40

Exergy analysis Exergy of a system includes four parts: kinetic exergy, potential exergy, physical exergy, and chemical exergy. In a chemical separation process, the kinetic and the potential exergy of the streams are presumably negligible.41 Total exergy of streams can be expressed as:

(1)

where is the physical exergy and is the chemical exergy. The physical exergy of a mixture can be calculated by the enthalpy ( ) and entropy () of the mixture at its current state (current temperature and pressure), in combination with the enthalpy ( ) and entropy ( ) of the mixture at the standard state (1 atm and 298.15K):


(2)


Page 30 of 44

The chemical exergy of a mixture can be obtained by the standard chemical exergies of the component and their mole fraction: ∑ , ∑

(3)

Here is the gas constant, is the temperature of the standard state (298.15K), is mole fraction of component , and , is standard chemical exergy of component . Standard chemical exergy of hydrocarbons can be obtained according to the QSPR model.42 Hence, exergy efficiency can be expressed as the ratio of the output exergy of the system to the input exergy of the system:

∑ !"

(4)

Based on the above method, the exergy efficiency of each of the seven design schemes are calculated, and the results are summarized in Table 9. It is observed that exergy efficiency of fully thermally coupled distillation (Design 4) is the same as exergy efficiency of DWC with the wall at the middle part of the column (Design 7). This is because Design 4 is thermodynamically equivalent to Design 7.23 In addition, the exergy efficiencies of fully thermally coupled distillation (Design 4) and DWC with the wall in the middle part of the column (Design 7) is higher than other schemes. Note that the indirect azeotropic distillation consequence (Design 3) has the lowest exergy efficiency as shown in Table 9. However, the exergy efficiency for some schemes (design 1, 2, 3, and 6) are relatively similar, the reason is that exergy loss is focused on the condenser and reboilers of the column. Table 9. Exergy efficiency of the seven proposed separation process designs. No.

Scheme

Exergy efficiency (%)

1

Conventional distillation

75.32

consequence 2

Direct azeotropic distillation

75.28

consequence 3

Indirect azeotropic distillation

75.25

consequence 4

Fully thermally coupled

77.01

distillation 5

DWC with the wall at the upper


76.78

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part of the column 6

DWC with the wall at the lower

75.26

part of the column 7

DWC with the wall at the

77.02

middle part of the column

Process economic analysis Total annualized cost (TAC) is a useful indicator for evaluating the economic performance of a process system. TAC includes capital investment and operating cost,43 and it can be expressed as:44 #$

%%& ()*+,*( + %-.% / *01

2345678 92:7

(5)

Here, payback period is used as 5 years. For the separation process, capital investment cost includes the cost of distillation column vessels, trays, and heat exchangers (cost of condenser and reboiler). The distillation column parameters can be obtained by tray sizing in Aspen Plus, and the value of Marshall and Swift (M&S) index is 1,672.43 The costs of the distillation columns, trays and heat exchangers are estimated based on the equations 45-46. Costs of the columns and trays for different schemes are summarized in Table 10. Compared with the conventional distillation scheme, the scheme of azeotropic distillation integrated DWC with the dividing wall in the middle of the column (Design 7) can save 40.71% capital investment cost in the columns and the trays. Costs of condensers and reboilers for different schemes are provided in

Table 11 and Table 12. The results show that costs of condensers and reboilers of the conventional distillation sequence are better than other schemes. The reason is that the conventional distillation scheme requires less capital investment than other separation schemes.



Page 32 of 44

Table 10. Capital costs of the column and tray for the seven designs.

No.

1

2

3

4

5

6

7

Scheme Conventional distillation Azeotropic distillation with direct sequence Azeotropic distillation with indirect sequence Azeotropic distillation with thermally coupled Azeotropic distillation integrated DWC with the wall at upper part Azeotropic distillation integrated DWC with the wall at lower part Azeotropic distillation integrated DWC with the wall at middle part

Column

Number of trays

Diameter (m)

C1 C2 C1 C2 C3 C4 C1 C2 C3 C4 C1 C2 C3

30 70 15 30 6 39 15 30 31 15 15 30 25

0.5 0.65 0.55 0.25 0.15 0.45 0.55 0.25 0.55 0.2 0.55 0.25 0.25

Height of the column (m) 19.14 45.54 9.24 19.14 3.3 25.08 9.24 19.14 19.8 9.24 9.24 19.14 15.84

C4

42

0.4

27.06

1197.26

63.58

C1 C2

15 30

0.55 0.25

9.24 19.14

710.62 549.87

11.23 10.57

C3

42

0.45

27.06

1357.42

76.32

C1 C2

15 30

0.55 0.25

9.24 19.14

710.62 549.87

11.23 10.57

Cost of column (103¥)

Cost of trays (103¥)

1151.19 3051.59 710.62 549.87 77.89 1277.17 710.62 549.87 1308.63 241.56 710.62 549.87 472.14

21.23 65.44 11.23 10.57 1.09 70.73 11.23 10.57 76.22 7.41 11.23 10.57 17.96

Total cost (103¥) 4289.45

2709.17

2916.11

3033.23

2716.03

3067.64 C3

42

0.55

27.06

1681.19

104.16

C1 C2

15 30

0.55 0.25

9.24 19.14

710.62 549.87

11.23 10.57 2543.13

C3

42

0.4

27.06


1197.26

63.58

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Table 11. Capital cost of the condensers in the seven designs.

No.

1

2

3

4

5

6

7

Scheme Conventional distillation Azeotropic distillation with direct sequence Azeotropic distillation with indirect sequence Azeotropic distillation with thermally coupled Azeotropic distillation integrated DWC with the wall at upper part Azeotropic distillation integrated DWC with the wall at lower part Azeotropic distillation integrated DWC with the wall at middle part

Column

LMTD (°C)

Heat duty (kW)

C1 C2 C1 C2 C3 C4 C1 C2 C3 C4 C1 C2

110.80 107.80 71.65 106.75 71.43 111.97 71.65 106.75 85.97 70.90 71.65 106.75

257.95 482.61 443.64 72.34 32.44 235.47 443.64 72.34 358.79 88.35 443.64 72.34

Heat transfer area (m2) 2.33 4.48 6.19 0.68 0.45 2.10 6.19 0.68 4.17 1.25 6.19 0.68

C4

71.12

193.70

2.72

198.23

C1 C2 C3 (Left) C3 (Right) C1 C2

71.65 106.75

443.64 72.34

6.19 0.68

338.10 80.26

71.11

21.00

0.30

467.77

115.28

182.50

1.58

139.32

71.65 106.75

443.64 72.34

6.19 0.68

338.10 80.26

Cost (103¥) 178.73 273.39 338.10 80.26 61.87 167.56 338.10 80.26 261.60 119.24 338.10 80.26

Total cost (103¥) 452.12

647.79

799.2

616.59

1025.45

713.48 C3

71.26

358.00

5.02

295.12

C1 C2

71.65 106.75

443.64 72.34

6.19 0.68

338.10 80.26 616.67

C3

71.05

193.64

2.73


198.31


Page 34 of 44

Table 12. Capital costs of the reboilers of the seven process designs.

No.

1

2

3

4

5

6

7

Scheme

LMTD (°C)

Heat duty (kW)

C1 C2 C1 C2 C3 C4 C1 C2 C3 C4 C1 C2

9.70 27.20 27.85 9.85 42.86 14.20 27.85 9.85 11.09 42.92 27.85 9.85

272.58 483.01 467.32 136.85 34.17 235.66 467.32 136.85 355.78 51.71 467.32 136.85

Heat transfer area (m2) 28.10 17.76 16.78 13.90 0.80 16.60 16.78 13.90 32.09 1.21 16.78 13.90

C4

10.09

196.47

19.47

711.93

C1 C2

27.85 9.85

467.32 136.85

16.78 13.90

898.20 794.73

Column

Conventional distillation Azeotropic distillation with direct sequence Azeotropic distillation with indirect sequence Azeotropic distillation with thermally coupled Azeotropic distillation integrated DWC with the wall at upper part Azeotropic distillation integrated DWC with the wall at lower part Azeotropic distillation integrated DWC with the wall at middle part

Cost (103¥) 1253.99 930.53 898.20 794.73 89.20 641.67 898.20 794.73 985.02 116.66 898.20 794.73

Total cost (103¥) 2184.52

2423.80

2794.61

2404.86

2327.13 C3

12.66

206.37

16.29

634.20

C1 C2 C3 (L) C3 (R) C1 C2

27.85 9.85

467.32 136.85

16.78 13.90

898.20 794.73

14.02

113.64

8.11

402.77

27.16

247.19

9.10

434.23

27.85 9.85

467.32 136.85

16.78 13.90

898.20 794.73

2127.16

2394.87 C3

10.31

196.50

19.05

701.94

The main contributors of the operating costs for the distillation separation processes are the reboiler heating and condenser cooling duties. To evaluate the cost of the utility of the separation process, the yearly operating time is set as 8000 hours.47-48 In addition, the costs of steam and coolant are taken as ¥275 per ton and ¥2 per ton,48 respectively. The flowrate of the steam (;+ ) and


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coolant (; ) can be estimated by the following equations: ?

;+

(10)

@A,"

Systems Design, Modeling, and Thermoeconomic Analysis of

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