Energy-Efficient ExtractionâDistillation Process for Separating

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Energy-Efficient Extraction-Distillation Process for Separating Diluted Acetonitrile-Water Mixture: Rigorous Design with Experimental Verification from Ternary Liquid-Liquid Equilibrium Data Bor-Yih Yu, Ray Huang, Xin-Yi Zhong, Ming-Jer Lee, and I-Lung Chien Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b04408 • Publication Date (Web): 06 Dec 2017 Downloaded from http://pubs.acs.org on December 7, 2017

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Industrial & Engineering Chemistry Research

Paper submitted for publication in Ind. Eng. Chem. Res.

Energy-Efficient Extraction-Distillation Process for Separating Diluted Acetonitrile-Water Mixture: Rigorous Design with Experimental Verification from Ternary Liquid-Liquid Equilibrium Data

Bor-Yih Yu1, Ray Huang1, Xin-Yi Zhong2, Ming-Jer Lee2,* and I-Lung Chien1,*

1

Department of Chemical Engineering National Taiwan University Taipei 10617, Taiwan

2 Department of Chemical Engineering National Taiwan University of Science and Technology Taipei 10607, Taiwan

Revised: November 30, 2017

*

Corresponding authors: I-Lung Chien, Tel: +886-3-3366-3063; Fax: +886-2-2362-3040; E-mail: [email protected] and Ming-Jer Lee, Tel: +886-2-2737-6626; Fax: +886-2-2737-6644, E-mail: [email protected]

1 Environment ACS Paragon Plus

Industrial & Engineering Chemistry Research 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

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ABSTRACT In this paper, an energy-efficient extraction-distillation process to separate diluted azeotropic acetonitrile-water mixtures is newly developed. Compared with the conventional azeotropic separation methods (i.e extractive distillation), the potential dominant benefit of this proposed method is that the main separation task can be achieved by an extraction column without needing of reboiler duty. In this work, an efficient solvent of n-propyl chloride is proposed to extract the organic compound into the extract phase and to let water to remain in the raffinate phase. Ternary liquid-liquid equilibrium experiments are also conducted to verify the separation performance in extraction column and decanter of the proposed process. It is found that significant savings of 40.3% in steam cost and 34.7% in total annual cost can be obtained by the proposed separation method as compared to that of a three-column extractive distillation system published in open literature.

Keywords:

Diluted

azeotropic

mixture;

extraction-distillation;

Acetonitrile

dehydration; n-propyl chloride; LLE experiments; Process design.


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1. Introduction There are plenty of applications in industry where a diluted azeotropic mixture needs to be separated to obtain both the valuable organic compound and also water at high-purity specifications. These separation processes are considered to be energy intensive because large amount of water needs to be separated out and also azeotropic separation cannot be achieved by conventional distillation. For these diluted azeotropic mixtures, it is often more economical to first purify the diluted mixture to near azeotropic composition and then goes into an azeotropic separation section by some commonly available separation method such as pressure-swing distillation, extractive distillation, or heterogeneous azeotropic distillation. By calculating the energy consumption of the whole separation process, the reboiler duty of the pre-concentration column is often comparable, if not greater, than that of the azeotropic separation section. Alternatively, extraction-distillation process is potentially a perfect candidate to fulfill the separation task for diluted azeotropic mixture. With an efficient solvent available, extraction column can be used to extract out the organic compound while leaving water in the raffinate phase. Further purification of the extract and raffinate phases are needed to purify the organic compound and also water at high-purity specifications and also to concentrate the solvent to be recycled back to the extraction



column. Since the main separation task is achieved by the extraction column without needing of reboiler, significant energy-saving can be gained by this separation method. The detailed design of extraction-distillation process is relatively scarce in open literature. For the acetic acid dehydration system, MTBE was proposed as an efficient solvent in Kürüm, et al.1 and Chilev and Lamari2, while ethyl acetate was proposed as suitable solvent in Lucia et al.3. For the acetone-methanol-butanol (ABE) fermentation broth, various solvents were proposed to extract out the organic compounds, such as mesitylene4, dual solvents of decanol and decane5, and n-hexyl acetate6-7. For the purification of the bio-based platform chemical γ–valerolactone, a massive solvent screening approach was performed using COSMO-RS8. For pyridine9 and n-propanol10 dehydration systems, diisopropyl ether was used as the solvent for these two separation systems. For the separation of ethanol/hexane or ethanol/heptanes11-13, various ionic liquids were used as solvent in liquid-liquid extraction. A review paper by Pereiro, et al.14 listed various separation systems by using ionic liquid as solvent. Other applications of using ionic liquid as solvent including: recovery of glycols from aqueous stream15; and selectively extract aromatics from naphtha feed to ethylene crackers by using an ionic liquid16 or mixed ionic liquids.17 In this paper, energy-efficient separation method for acetonitrile dehydration


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process is investigated. From the published studied of this separation process18, separation method via extractive distillation was used to obtain high-purity products. The chosen heavy entrainer was ethylene glycol (EG). A further energy-saving design method was also proposed in that paper to combine pre-concentration column and entrainer recovery column. The main contribution of this paper is to newly develop the proposed extraction-distillation processes for this separation process. The economics and steam cost of our proposed separation method will be compared with published results via extractive distillation system. To the best of our knowledge, none of the published literatures have studied the detailed design flowsheets of this dehydration system via extraction-distillation process. To make sure separation performance in the extraction column and a decanter can be achievable; liquid-liquid equilibrium experiments are also conducted with the experimental data used to refine the thermodynamic model in the simulation investigations. 2. Previous Published Extractive Distillation Process for Comparison Purpose In Liang, et al.18, the same process which separated diluted azeotropic mixture of acetonitrile (ACN) and water was presented. The feed contains 80 mol% water and 20 mol% ACN. The feed was a saturated liquid with feed rate of 500 kmol/h. The two product specifications were set at high-purity of 99.99 mol%. In this paper, an



alternative more energy-efficient separation method via extraction-distillation process will be proposed. In order for direct comparison, all the feed condition and two product specifications are set to be exactly the same as the ones in Liang, et al.18. In Liang, et al.18, ethylene glycol (EG) was used as a heavy entrainer to enhance the relative volatility of ACN over water. For the process simulation study of that process, NRTL is selected as the thermodynamic model. All the NRTL binary parameters used in the simulation of this ACN case study in this and the next sections can be found in Table 1. Figure 1 duplicated the simulation result in Liang, et al.18 with only some minor differences. The differences might be due to lack of full information from that paper, such as pressure drop in the columns, etc., for us to completely replicate their results. From this design flowsheet in Fig. 1, a pre-concentration column is used to first purify the diluted feed to 65 mol% ACN and then goes into a two-column extractive distillation system. The design of this pre-concentration column is intuitive to remove much of water as earlier as possible. Otherwise, large amount of water will travel through the extractive distillation column and then goes out of the system from the distillate of entrainer recovery column. Note that the reboiler duty of pre-concentration column (1659 kW) is significant as compared to the total reboiler duty of the overall


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extractive distillation system. In Liang, et al.18, an improved design flowsheet was proposed to combine pre-concentration column and entrainer recovery column into one column. The design of the combined column can be considered as a stack column with pre-concentration column on top and entrainer recovery column at the bottom. A side reboiler (functioning as the reboiler for preconcentration column) is needed and the overall withdrawal of water is from a side stream of this combined column. With this design, part of the reboiler duty needed for the pre-concentration column can be saved. The saved reboiler duty can be roughly estimated as the condenser duty of the original entrainer recovery column because in the combined design flowsheet this condenser is no longer needed. The detailed design of this energy-saving flowsheet can be seen in Liang, et al.18. Since significant amount of reboiler duty was needed to concentrate the diluted azeotropic mixture, a much improved design flowsheet with respect to energy consumption will be proposed next for this ACN dehydration system. In the proposed extraction-distillation process, main separation task will be fulfilled by an extraction column without needing of reboiler. The economics and total steam cost of the original design flowsheets in Figure 1 and also the stacked two-column design in Liang, et al.18 will be compared with that of a newly proposed design flowsheet via



extraction-distillation process. 3. Conceptual Design of Extraction-Distillation Process 3.1. Solvent selection The key point for the success of this proposed separation method is the selection of an efficient solvent which can extract ACN into extract phase while let most of the water to be in raffinate phase. The conceptual design flowsheet of this proposed process can be seen in Figure 2. The overall process contains an extraction column, two strippers, and another decanter. The extract phase is further purified in a C-1 stripper to obtain bottom ACN product. The top of this C-1 stripper, after condensation, is designed to go to a decanter. The organic phase of the decanter should majorly be the solvent component, thus, is recycled back to the extraction column. The aqueous phase of the decanter together with the raffinate phase of extraction column will send to another C-2 stripper to obtained bottom water product satisfying purity specification. The top of this C-2 stripper, after condensation, is combined with the fresh feed to the extraction column. From the conceptual design flowsheet in Figure 2(a), it is clear that a large liquid-liquid envelop is favorable by adding the solvent into the system. In the ternary diagram, the liquid-liquid tie-lines should be tilted to point into the water end so that the


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composition of raffinate phase can contain mostly water. Also because the ACN product is withdrawn from the overall process at a bottom stream, Thus, the solvent should be lighter (lower n.b.p.) than ACN in order to have feasible separation in C-1 stripper. The ternary diagrams of four possible solvent candidates are given in Figure 3. All the NRTL binary parameters used in obtaining the ternary diagrams can be found in Table 1. The four candidate solvents are: n-propyl chloride (NPC); di-ethyl ether (DEE); diisopropyl ether (DIPE); and n-Propyl Formate (NPF). These candidates are taken from tables of commonly used solvents in Kürüm et al.1 and Gmehling and Möllmann19. It is found that using NPC, DEE and NPF as solvent can have nice property of the tilted tie-lines. Also, both DIPE and NPF exhibit additional distillation boundary prohibiting enrichment of solvent concentration at the top of C-1 stripper. As for DEE, the composition of the possible extract phase will be too close to the distillation boundary to hinder the separation at C-1 column top. Based on the above-mentioned descriptions, it is found that NPC is the most appropriate solvent for this dehydration system. Other factors such as heat of vaporization and toxicity property of the solvent are also important in determining the suitable solvent. Lower heat of vaporization would reduce the reboiler duty of C-1 stripper. In Table S1 of the supporting information, properties for the solvent candidates are listed, including heat of vaporization, lethal



dose 50% (LD50), and IARC category. It is found that NPC gives a low heat of vaporization and also is not carcinogenic to humans. 3.2. Thermodynamic Parameters In this paper, all the simulation work is performed in Aspen Plus V8.8, and NRTL is selected as the global thermodynamic model. When performing simulation, the binary interaction parameter pairs between different components play a big role, as they determine the phase behavior, which influences the description of separation performance. Thus, selecting the correct parameters for simulation is always vital. In the Aspen Plus database, there are built-in parameter pairs which were regressed from the experimental data in the literature. Those parameters pairs are usually reliable. However, sometimes there are lacking of binary parameters, and the typical way to solve this issue is to estimate them through UNIFAC group contribution method. Or, more accurately, ternary liquid-liquid equilibrium experiments can be conducted and to regress the required parameters from the experimental data. In this work, the extraction-distillation process for ACN dehydration with NPC as solvent was investigated. There are three components in this process, namely, ACN, H2O and NPC. Thus, three binary interaction pairs are needed to describe the behavior among these components. There are built-in parameters for the ACN-H2O and


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H2O-NPC pairs, but not for the ACN-NPC pair. Here, firstly in the previous solvent selection in Section 3.1, this pair was estimated by UNIFAC method. Since the actual ternary liquid-liquid equilibrium behavior of these three components is most important in the design study of an extraction column and a decanter, in the following, sets of experiments were conducted to validate the thermodynamic model parameters in binary pairs. Regression of the parameters in Section 4 was made to more accurately represent the phase equilibrium behavior. After that, an optimal process flowsheet is developed based on these regressed parameter settings, and all the detailed simulation results regarding this proposed process flowsheet can be found in Section 5.

4. Ternary Liquid-Liquid Equilibrium Experimental Verification In this section, ternary liquid-liquid equilibrium experiments were conducted to obtain a reliable set of thermodynamic parameters for the design study of this extraction-distillation process. 4.1. Materials n-Propyl chloride (99 %) and acetonitrile (99.9 %) were purchased from ACROS Organics (USA) and used without further purification. Double distilled deionized water was prepared by NANO pure-Ultra pure water system with resistivity of 18.3 MΩ-cm. 4.2. Apparatus and procedure



A glass-made LLE apparatus, similar to that reported in Peschke and Sandler22 is employed in this study for LLE tie-line data measurement. The internal volume of the equilibrium cell is about 10 cm3. Thermostatic water is circulated through the jacket, which is surrounded the cell, to maintain cell’s temperature to within ±0.1 K. A precision thermometer (Model-1560, Hart Scientific Co., USA) is used for measuring the temperature of solution inside the cell to an uncertainty of 0.02 K. The loaded ternary solution in the cell is agitated vigorously with a magnetic stirrer at least four hours and then settled at least eight hours to completely separate the organic-rich and the water-rich phases in the cell. A sample of aqueous phase is taken from the bottom sampling port and that of organic-rich phase from the sampling port at the top of the cell. At least five replicated samples (about one µL of each) are taken from each coexisting liquid phase. The composition of each sample is determined by gas chromatography (Model 8700, China Chromatography Co., Taiwan) equipped with a thermal conductivity detector. A stainless steel packed column (Porapak Q, 50/80 mesh, and 2 m of length) is utilized for separating the constituent compounds of samples. Helium (99.99 %) is used as a carrier gas. The area fraction detected by the gas chromatography is converted into mole fraction via the calibration equations of water + acetonitrile and n-propyl chloride + acetonitrile. Calibrations are made with gravimetrically prepared


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standard samples and checked frequently. The uncertainty of composition analysis is estimated to be 0.0005 in mole fraction. 4.3. Experimental data The reliability of the above mentioned LLE measurement has been checked with literature data previously.23 In the present study, the same method is used to measure the LLE tie-line data for the ternary system of n-propyl chloride + acetonitrile + water at 298.15 K and 308.15 K under atmospheric pressure. Table 2 reports the compositions of two coexistence liquid phases for this ternary system. 4.4. Othmer-Tobias Correlation The phase compositions of the two ends of the tie-lines were correlated with the empirical equation of Othmer-Tobias24 as below:

[(

)

]

[(

) ]

ln 1 − w3II / w3II = a + b ln 1 − w1I / w1I II

(1) I

where w3 is the mass fraction of water in the aqueous phase and w1 is the mass fraction of n-propyl chloride in the organic-rich phase. Table S2 of the supporting information lists the correlated results, indicating that the square of correlation coefficients (R2) is greater than 0.99. Figure S1 of the supporting information illustrates the results, which well follows the linear relationship as given by eqn 1. This illustrates that the experimental LLE tie-line data measured in the present study are very



consistent. 4.5. Comparisons of experimental data with previous model prediction The originally predicted liquid-liquid equilibrium (LLE) behavior in Section 3.1 is firstly compared with the experimental data at 298.15 K and 308.15 K, which is included in Table 2. Figure 4(a) and 4(c) show the outcomes of comparison at these two operating temperatures. It is clearly shown that the originally predicted LLE behaviors in Section 3.1 do not match the experimental data well, because of the following two aspects. Firstly, the actual measured liquid-liquid envelope is smaller than using the model parameters in original Table 1. Secondly, the slopes of the tie lines from model prediction are slightly overly estimated. These reasons lead to the fact that the design study based on the model parameters in original Table 1 may be too optimistic. Hence, the parameters in the binary pairs should be regressed from the given experimental data. 4.6. Refitting of thermodynamic model parameters When reevaluating the binary pairs to fit the experimental results, it may be intuitive to firstly start from refitting the NPC-ACN pair, which was originally estimated through UNIFAC. This is because the remaining two parameter pairs (ACN-H2O and NPC-H2O) were from the Aspen built-in database and may be more reliable. However, if only the NPC-ACN pair is reevaluated while the other two pairs are kept unchanged, the


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outcome of fitting was still far from satisfactory. This means that the two built-in pairs may accurately describe the binary vapor-liquid equilibrium behavior, but not satisfactory in describing the ternary liquid-liquid equilibrium behavior under the sub-cooled temperature. In order to deal with this issue, all three binary pairs should be reevaluated. In all three pairs, aij, aji, bij and bij are regressed. For the ACN-H2O and the ACN-NPC pairs, cij is set at 0.3 as the original value. For the NPC-H2O pair, cij is also regressed in order to more accurately fit the LLE boundary and the slopes of the tie lines. The data used for regression includes the LLE data shown in Table 2, as well as another ACN-water vapor-liquid equilibrium (VLE) data at normal pressure from the ASPEN NIST databank. This ensures the re-estimated parameters to accurately describe both LLE and VLE. The new parameters of all binary pairs from regression are listed in Table 3, while the refitting ternary LLE results are illustrated in Figure 4(b) and 4(d). From these two sub-figures, the fitting becomes much better. Also, the predicted temperature and composition of the azeotropes are shown in Table S3 of the supporting information, which show highly correspondence to the original information. Thus, the new parameters can be concluded to be capable of accurately predicting both VLE and LLE



behaviors.

5. Proposed extraction-distillation design flowsheet 5.1. Development of the proposed flowsheet based on model parameters in Table 3 The conceptual design flowsheet of the proposed extraction−distillation process can be found in previous Figure 2(a). The pressures of the extraction column and the decanter were all set at atmospheric pressure for easier operation. The top vapor temperatures of the two strippers were all set to be greater than 320 K so that cooling water can be used as cooling medium in the top condenser of the two strippers. If the top vapor temperature of the stripper is lower than 320 K, the stripper operating pressure can be elevated to avoid using too much of the more expensive chilled water. The inlet temperatures of all liquid streams into extraction column and also into decanter were also needed to be determined. Since n.b.p. of the NPC solvent is only at 46.52 ºC (319.67 K), inlet temperatures of all liquid streams into extraction column and decanter were conservatively set at 35 ºC (308.15 K) to make sure no vapor loss of the solvent is possible. Thus, some extra chillers are needed in the design flowsheet to achieve this goal. The remaining design variables in the conceptual design flowsheet include: solvent-to-feed ratio (FS/FF); extraction column total stages (EXT); and total stages of


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C-1 stripper (NT1); and total stages of C-2 stripper (NT2). Sequential iterative optimization procedure outlined in Figure S2 of the supporting information is used to obtain the optimal design flowsheet by minimization of TAC. In all the simulation runs, product purity specifications of ACN and water are set at 99.99 mol% by varying the respective reboiler duties. The TAC of the overall system includes the annualized capital costs and the operating costs. The annualized capital costs include: costs of extraction column, two strippers, decanter, reboilers, condensers and chillers. The payback period was assumed to be three years in the calculation. The operating costs include: steam for two reboilers, cooling water for two condensers, chilled water for chillers and also a small solvent makeup cost. The bases for calculations of all terms in TAC were mainly adopted from Luyben.20 The diameter and height of the extraction column are adopted from Seider et al.21 Table S4 in the supporting information summarizes the formulas for all the TAC calculations. The TAC plots for obtaining the optimal design flowsheet is shown in Figure S3. The top part of this figure shows the TACs by varying total stages of the extraction column and two strippers. It is shown that the minimum TAC was occurred at EXT=4, NT1=15 and NT2=10. If NT1 decreases further, ACN starts to come out more at the top



vapor stream, which negatively impacts the separation performance in both the decanter and the extraction column. Thus, this accumulation finally causes the failure of reaching the separation target. The summary plot of minimized TAC at each FS/FF ratio is also shown in the bottom plot of Figure S3. It is intuitive that the optimal FS/FF ratio reaches a low constraint of the system. This means that a minimum solvent flow rate is needed to satisfactorily perform the liquid-liquid separation task. The optimal design flowsheet with the optimal FS/FF at 0.36 is shown in Figure 5. This design flowsheet can be considered as a base case for further dynamic study investigating operation and control of the proposed process. It is worth mentioning that a suitable overall control strategy for the extraction-distillation process has been developed in Chang and Chien10 with design flowsheet at this minimum FS/FF ratio. Even with unmeasured feed composition changes, amount of solvent can be properly adjusted upward/downward by a ratio scheme to the reboiler duty of C-1 stripper. Since the control study is just a repetitive work of that in Chang and Chien10, the control study is not included in this paper. 5.2. Results comparisons Table 4 compares the itemized TAC of this proposed design with previously published three-column extractive distillation process and stacked two-column process


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in Liang, et al.18 Significant reductions of 40.3% in steam cost and 34.7% in TAC can be obtained by the proposed design as compared with previous published extractive distillation process. By checking the total stages of the proposed design flowsheet in Figure 5 with that of the extractive distillation system in Figure 1, it is quite clear that the needed total stages of the proposed design is much less than that of the previous published separation system. As compared to the more complex stacked two-column process in Liang, et al.18, the proposed design can still significantly save steam cost and TAC by 29.8% and 25.0%, respectively. 5.3. Possible Further heat-integration of the proposed design flowsheet As the operating temperature inside the decanter and the extraction column is low, heat integration can be performed to preheat the feed stream of stripper to save some energy requirement in the stripper. Note that the fresh feed temperature is high; which needs to be cooled before sending into extraction column. Therefore, the removed energy in the cooler can be recovered by preheating the feed stream into strippers. Also the temperatures of two bottom product streams are high, which can be used to preheat the feed streams of strippers as well. However, special attention should be paid when we intend to apply heat integration strategy to preheat the feed stream to the C-1 stripper. The top vapor stream of C-1



stripper contains mostly NPC. From Figure 5, it is shown that the temperature difference of the C-1 feed stream and the top vapor stream is only about 35 K. If this C-1 feed stream is preheated, a greater trend of vaporization can be expected, and much more amounts of ACN may come out from the top vapor stream. This is an undesired situation, as it can negatively influences the separation performance in decanter and extraction column which is already addressed in previous section 5.1. The possible heat integration is to preheat the C-2 stripper feed to save the steam requirement in this reboiler. From Figure 5, it is shown that the temperature difference between the C-2 stripper feed and top vapor is about 52 K. Thus there will be more room to save energy. Here, the authors only discuss the possible way of heat integration but not investigate in details, because the further improvement will not be that significant as compared to that of changing the separation method as proposed in this paper. 5.4. Design flowsheet based on model parameters in original Table 1 It is interested to know the discrepancy of the simulation results in the design flowsheet when the missing thermodynamic model parameters were estimated by UNIFAC method. The process studied in previous section 5.1 was investigated again using those original thermodynamic parameters in Table 1. For optimization, the same


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algorithm shown in Figure S2 and the same objective function are used here again. The optimal design flowsheet is shown in Figure 6, in which the optimal design variables to minimize TAC are found to be: FS/FF=0.21; EXT=4; NT1=15; and NT2=10. To support these findings of results, the detailed TAC plots are included in Figure S4 of the supporting information. The itemized TAC for this design flowsheet is also included in the middle column of Table 4. For this ACN dehydrating system, the optimal FS/FF reduces from 0.36 to 0.21 when UNIFAC estimated parameters in Table 1 were used. This is because the slopes of the tie lines in LLE envelope predicted by the original parameters in Table 1 were steeper together with too large of the LLE envelope. The TAC and steam savings were more optimistic than the results of using more reliable parameter set in Table 3. This demonstrates the importance of experimental verification for this design study.

6. Conclusions Separation of diluted azeotropic mixture is known to be energy intensive because large amount of water needs to be separated out at high-purity. In this paper, energy-efficient process for separation of ACN and water are presented. The proposed design flowsheet for the separation of ACN and water via extraction-distillation process is newly developed. It is found that n-propyl chloride is a suitable solvent for this



dehydration system. Apart from performing simulation study, separation behavior in this system is validated by conducting ternary liquid-liquid equilibrium experiments. Since the main separation task is achieved by an extraction column without needing of reboiler, significant savings in steam cost and TAC of the proposed design can be obtained as compared to that of conventional azeotropic separation systems published in open literature. With the correct thermodynamic description, significant reductions of 40.3% in steam cost and 34.7% in TAC can be obtained by comparing with a three-column extractive distillation process. Furthermore, the proposed design can also provide 29.8% savings in steam cost and 25.0% in TAC by comparing the economics with that of more complex stacked two-column extractive distillation system.

Acknowledgements The research funding from the Ministry of Science and Technology of R. O. C. under grant no. MOST 106-2221-E-002-176 is greatly appreciated.

Supporting Information This part includes some detailed information in tables and figures, and it is available free of charge via the Internet at http://pubs.acs.org/.


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from the Azeotropic Mixture Ethanol + Hexane with an Ionic Liquid as Solvent. Chem. Eng. J. 2009, 153, 80-85. (12) Pereiro, A. B.; Rodríquez, A. Azeotrope-breaking Using [BMIM][MeSO4] Ionic

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from naphtha feed to ethylene crakers using 1-ethyl-3-methylimidazolium thiocyanate ionic liquid. Chem. Eng. Res. Des. 2016, 120, 102-112. (17) Navarro, P.; Larriba, M.; García, J.; Rodríguez, F. Design of the recovery section

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Page 26 of 40

Page 27 of 40 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


Table 1. NRTL model parameters used in Sections 2 and 3 of the paper. Comp. i

ACN

Water

ACN

Water

ACN

Water

ACN

Comp. j

Water

EG

EG

DIPE

DIPE

NPF

NPF

Source

ASPEN VLE-IG

ASPEN VLE-IG

ASPEN VLE-IG

ASPEN LLE

UNIFAC

ASPEN LLE

UNIFAC

aij

-0.1164

0.3479

0

8.0209

0

146.2329

0

aji

1.0567

-0.0567

0

0.035

0

24.3719

0

bij (K)

256.4588

34.8234

536.542

bji (K)

283.4087

-147.1373

130.1648

cij

0.3

0.3

0.3

eij

0

0

0

0

0

-20.6911

0

eji

0

0

0

0

0

-4.1008

0

-766.4165 509.063591 -7166.7568 45.6194046 422.9778 313.297148 88.2652 73.2160055 0.2 0.2 0.3 0.3

Comp. i

Water

ACN

Water

ACN

Comp. j

DEE

DEE

NPC

NPC

Source

ASPEN LLE

Aspen VLE-LIT

ASPEN LLE

UNIFAC

aij

107.1788

0

8.3833

0

aji

-43.365

0

8.5029

0

bij (K)

-6697.9443

292.9133

-848.3878 200.436606

bji (K)

3154.5229

43.8872

-883.1837 174.919183

cij

0.2

0.2982

0.2

0.3

eij

-14.3108

0

0

0

eji

6.0307

0

0

0

Aspen Plus NRTL:

∑ x jτ jiG ji ln γ i =

j

∑ xk Gki k

  +∑ τ ij −  x G ∑ k kj j  k  x j Gij

∑ xmτ mj Gmj  m

∑ xk Gkj k

  

Where : Gij = exp( −α ijτ ij )

τ ij = aij +

bij T

+ eij ln T

α ij = cij , τ ii = 0 , Gii = 1


Industrial & Engineering Chemistry Research 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

Page 28 of 40

Table 2. LLE tie-line data for n-propyl chloride (1) + acetonitrile (2) + water (3) at 1 atm* Organic-rich phase

T/K

298.15

308.15

*

Water-rich phase

‫ݔ‬ଵூ

‫ݔ‬ଶூ

‫ݔ‬ଷூ

‫ݔ‬ଵூூ

‫ݔ‬ଶூூ

‫ݔ‬ଷூூ

0.7169

0.2620

0.0211

0.0065

0.0680

0.9255

0.4822

0.4544

0.0634

0.0026

0.0944

0.9030

0.2903

0.5774

0.1323

0.0051

0.1212

0.8737

0.2134

0.5997

0.1869

0.0060

0.1348

0.8592

0.1494

0.6103

0.2403

0.0046

0.1554

0.8400

0.6950

0.2773

0.0277

0.0019

0.0690

0.9291

0.5026

0.4333

0.0641

0.0033

0.0935

0.9032

0.2668

0.5653

0.1679

0.0042

0.1305

0.8653

0.2054

0.5802

0.2144

0.0049

0.1476

0.8475

0.1775

0.6018

0.2207

0.0061

0.1659

0.8280

xi: mole fraction of component i.; standard uncertainties: u(T) = 0.1 K; u(P) = 2 kPa; u(xi) = 0.0005.

28

ACS Paragon Plus Environment

Page 29 of 40 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


Table 3. Regressed NRTL Model Parameters of the ACN Dehydration System Comp. i

ACN

Water

NPC

Comp. j

Water

NPC

ACN

Source

Regression Regression Regression

aij

-0.5375

22.5775

-5.4753

aji

0.7547

14.5988

3.6972

bij (K)

420.5127

-5186.4659 1611.0853

bji (K)

326.1765

-1979.0303 -495.5049

cij

0.3

0.2689

0.3

eij

0

0

0

eji

0

0

0



Page 30 of 40

Table 4. Itemized TAC terms for ACN dehydration system (Unit for costs: 1000 USD/yr) [*: data compared with three-column extractive distillation process; **: data compared with stacked two-column process] Extractive Distillation (three-column system) in Liang, et al.18 C1 C2 C3 Annualized Capital Cost

Extractive Distillation (stacked two-column) in Liang, et al.18 C1/C3 C2

Extraction-Distillation Process Original Parameters in Table 1

Proposed Extraction-Distillation Process Regressed Parameters in Table 3

C1

C2

C1

C2

column

42.29

131.80

37.16

77.65

131.80

33.95

16.69

46.21

18.95

reboiler

31.43

45.06

25.59

156.81

45.06

24.86

23.60

32.84

24.37

condenser

32.02

25.01

14.21

32.02

25.01

33.39

15.24

42.44

15.71

4.90

9.67

7.40

9.78

chiller cooler

13.51

13.51

21.73

22.38

decanter

-

-

23.51

32.99

Extractor

-

-

27.78

31.37

steam cost

379.34

409.70

326.97

538.91

409.70

275.37

233.75

419.64

246.21

cooling water cost

16.04

12.54

7.62

16.04

12.54

12.98

5.77

16.83

4.65

chilled water cost

-

-

-

-

-

6.59

18.78

12.45

19.09

reboiler duty (MW)

1.69

1.73

1.15

2.10

1.73

1.23

1.04

1.87

1.10

total reboiler duty (MW)

4.57

3.83

2.27

2.97

makeup cost

0.34

0.35

31.55

12.11

total steam cost

1116.01

948.61

509.12(-54.4%* / -46.3%**)

665.84 (-40.3%* / -29.8%**)

annualized total capital cost

389.10

372.20

235.32

284.45

Total operating cost

1170.9

986.87

584.79

734.69

TAC

1560.00

1359.07

820.11(-47.4%* / -39.7%**)

1019.14 (-34.7%* / -25.0%**)

30


Page 31 of 40 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


Figure Captions Figure 1

Duplicated design flowsheet for the ACN dehydration system in Liang, et al.18.

Figure 2

(a) Conceptual design flowsheet of extraction−distillation process. (b) Material balance lines of extraction−distillation process.

Figure 3

Ternary diagrams of ACN dehydration system using several candidate solvents.

Figure 4

Validation of liquid-liquid equilibrium for the ACN-H2O-NPC system. (a) Original parameters at T=298.15 K; (b) New parameters at T=298.15 K; (c) Original parameters at T=308.15 K; (d) New parameters at T=308.15 K.

Figure 5

Proposed

design

flowsheet

of

ACN

dehydration

system

via

extraction-distillation process (regressed NRTL model parameters in Table 3). Figure 6

Design flowsheet of ACN dehydration system via extraction-distillation process (original NRTL model parameters in Table 1).



Page 32 of 40

Figure 1

Figure 1 Duplicated design flowsheet for the ACN dehydration system in Liang, et al.18.


Page 33 of 40 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


Figure 2



Figure 2

Page 34 of 40

(a) Conceptual design flowsheet of extraction−distillation process. (b) Material balance lines of extraction−distillation process.

Figure 3

Figure 3

Ternary diagrams of ACN dehydration system using several candidate

solvents.


Page 35 of 40 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


Figure 4

Figure 4

Validation of liquid-liquid equilibrium for the ACN-H2O-NPC system. (a) Original parameters at T=298.15 K; (b) New parameters at T=298.15 K; (c) Original parameters at T=308.15 K; (d) New parameters at T=308.15 K.



Page 36 of 40

Figure 5

36


Page 37 of 40 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


Figure 5

Proposed design flowsheet of ACN dehydration system via extraction-distillation process (regressed NRTL model parameters in Table 3).

37



Page 38 of 40

Figure 6

38


Page 39 of 40 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


Figure 6

Design flowsheet of ACN dehydration system via extraction-distillation process (original NRTL model parameters in Table 1).

39



Table of Contents (TOC) Graphic

40


Page 40 of 40

Energy-Efficient ExtractionâDistillation Process for Separating

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