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Process evaluation on the separation of ethyl acetate and ethanol using extractive distillation with ionic liquid
Accepted Manuscript Process Evaluation on the Separation of Ethyl acetate and Ethanol using Extractive Distillation with Ionic Liquid Zhaoyou Zhu, Yongsaeng Ri, Hui Jia, Xin Li, Yong Wang, Yinglong Wang PII: DOI: Reference:
S1383-5866(17)30026-6 http://dx.doi.org/10.1016/j.seppur.2017.03.011 SEPPUR 13597
To appear in:
Separation and Purification Technology
Received Date: Revised Date: Accepted Date:
3 January 2017 5 March 2017 9 March 2017
Please cite this article as: Z. Zhu, Y. Ri, H. Jia, X. Li, Y. Wang, Y. Wang, Process Evaluation on the Separation of Ethyl acetate and Ethanol using Extractive Distillation with Ionic Liquid, Separation and Purification Technology (2017), doi: http://dx.doi.org/10.1016/j.seppur.2017.03.011
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Process Evaluation on the Separation of Ethyl acetate and Ethanol using Extractive Distillation with Ionic Liquid Zhaoyou Zhua, Yongsaeng Ria, b, Hui Jiaa, Xin Lia, Yong Wanga, Yinglong Wanga, * a
College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao
266042, China b
Faculty of Chemistry, Kim Il Sung University, Pyongyang 999093, DPR of Korea
Corresponding Author *E-mail: [email protected]. Abstract This paper provides process design and simulation methodology for the separation of ethyl acetate and ethanol by extractive distillation using ionic liquids (ILs) as solvents and obtains the design parameters of extractive distillation process with solvent recovery system on the basis of a suitable IL solvent. The feasibility of IL-extractive distillation was examined via process simulation for the separation of ethyl acetate and ethanol in Aspen Plus. Four ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] were created in Aspen Plus database with several thermodynamic and physical property parameters to allow the process to be simulated via the UNIFAC-Lei thermodynamic method. The results show that the separation process containing the hybrid regeneration system of flash tank and stripper with [EMIM][MeSO3] as a suitable solvent is the best option in extractive distillation for the separation of ethyl acetate and ethanol compared with other ILs from the analysis of relative volatilities. The separation process was optimized by sensitivity analysis and the optimal design parameters were verified by economic evaluation based on the total annual cost (TAC). Keywords: Extractive distillation; Ionic liquid; UNIFAC-Lei; VLE; Simulation 1. Introduction In the history of chemical separations, conventional distillation has been applied to more commercial processes than all other techniques combined [1]. However, for systems with close boiling point or azeotropic systems, a separation by conventional distillation process becomes difficult or even impossible [2-4]. Nowadays, extractive distillation with ILs as solvents has become a promising alternative for these systems such as ethanol dehydration [5-7]. Extractive distillation is known as an efficient technology in the separation of complex mixtures by the 1
addition of a solvent at the top of the column that modifies the activity coefficients in the liquid phase to increase the relative volatilities [8-10]. The effectiveness of an extractive distillation process relies on the choice of solvent [11-13]. Since Lei’s group [14] first reported the extractive distillation process with ILs as green solvents which possess the advantages of a liquid solvent (easy operation) and a solid salt (high separation ability) in 2003, ILs are under intensive investigation to determine their potential as replacement solvents for extractive distillation. ILs have the outstanding physicochemical properties of low vapor pressure, low melting point, and high chemical and thermal stability [15-19]. So ILs are suggested by many authors for a broad range of separations, but the experimental way to determine the thermodynamic properties of ILs and the capability of the preferential IL solvent is a time consuming task because a large number of ILs exist [20-23]. The application of predictive models to estimate the thermodynamic properties of ILs and their mixture with water or organic compounds is of great interest. Indeed, investigation on predicting ILs properties by UNIFAC-Lei and COSMO-RS method gained focus in recent [24-28]. Selection of property methods depends on the non-ideal degree and operating conditions of the system, and a suitable property method determines the accuracy of the simulation results such as operating temperature and pressure of the system. In the works that the conceptual design of separation processes using ILs was carried out [7, 29-32], most of the property methods selected to the separation of azeotropic mixtures is the NRTL model. Nancarrow et al. [33] reported the simulation method of IL-based desulfurization processes by UNIFAC-Lei model in Aspen Plus environment and the feasibility of industrial scale IL-extractive processing of the ultra low sulfur diesel. However, to our knowledge, there is no study on the separation process of the azeotropic mixture by IL-extractive distillation simulation using UNIFAC-Lei as property method. In this paper, the UNIFAC-Lei model was employed to separate the close-boiling mixture of ethyl acetate and ethanol mixture using ILs as solvents. Santacesaria et al. [34] described that the production process of ethyl acetate by the dehydrogenation of ethanol has become very attractive considering the growing production of bio-ethanol for the fuel market, and also this reaction is simple, noncorrosive, less-toxic process and requiring only an ethanol feedstock. But the separation of ethyl acetate and ethanol mixture produced from the reaction system is among the difficult separation processes due to their close-boiling points and formation of an azeotropic 2
mixture [35], as boiling points of ethyl acetate and ethanol are 78.31℃ and 77.20℃, respectively. The organic solvents currently used in extractive distillations for separating ethyl acetate from ethanol are an amine (e.g., diethylene triamine), an alkylated thiopene (e.g., ethyl thiopene), and paraffins (e.g., dodecane) [36]. Recently, the ability of extractive distillation by ILs in facilitating the separation of azeotropic mixture of ethyl acetate - ethanol has been experimentally illustrated in some works [37-41]. However, the separation process with the consideration of effects between IL solvent and design parameters and the ability of solvent recovery for the continuous process have not been studied. The purpose of this article is to separate ethyl acetate and ethanol mixture by extractive distillation process based on the selection of a suitable IL solvent. Firstly, the vapor-liquid equilibrium (VLE) behaviors of ethyl acetate and ethanol with different ILs were described by the UNIFAC-Lei model and the separation effects were compared by relative volatilities of ethyl acetate when using ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] as solvents. The four ILs were preliminary selected as a set of study object which owned the ability of eliminating ethyl acetate and ethanol azeotrope in the relatively low concentrations among ILs reported by several literatures [37-41]. In addition, an IL that has the best performance among the four ILs in separating the ethyl acetate and ethanol mixture was employed to seek the conditions and configuration for the separation process by sensitivity analysis. TAC was calculated for the different configurations producing ethyl acetate with the distillate purity of 99.9 mol % to confirm a novel separation process configuration. 2. Method description 2.1. The UNIFAC-Lei model for IL-based systems UNIFAC-Lei computes the activity coefficient of liquid mixtures in terms of constants which reflects the size and surface area of individual functional groups present in the mixture. If the necessary binary group interaction parameters exist for every pair of functional groups of the mixture, UNIFAC-Lei model is worthy of being considered to predict the phase behavior [26, 27, 42, 43]. The activity coefficient can be calculated by
ln i ln i C ln i R 3
(1)
where
represents the combinatorial contribution to the activity coefficient due to differences
in the sizes and shapes of the groups, and energetic interactions. The part
represents the residual contribution from
contains the group parameters Rk and Qk, which values for
the groups investigated in this work can be found in the literature, as listed in Table 1. The part is a function of binary interaction parameters between groups, which group interaction parameter
is expressed as follows:
nm exp anm / T where T is the absolute temperature, of functional groups n and m, and
(2)
is the temperature-independent parameter for each pair .
ILs have to be transformed to functional group when using the UNIFAC-Lei model, and the fragmentations for ILs used in this work are listed in Table 2. ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] have been constructed from the new [MIM][MeSO3], [MIM][MeSO4], [MIM][CF3SO3] and [MIM][BF4] groups combined with the established CH2 and CH3 groups existed in the UNIFAC-Lei model. The group interaction parameters (
) between each pair of functional groups in their mixture with ethyl
acetate and ethanol taken from Lei et al. [26, 27] are listed in Table 3. 2.2. Process simulation For these three-component systems, the UNIFAC-Lei model was used to describe the nonideality of the liquid phase while the vapor phase is assumed to be ideal. The extractive distillation processes containing ILs were simulated using Aspen Plus v8.4 with the rigorous equilibrium stage model RadFrac. New component ILs were created as user-defined components in Aspen Plus by specifying their molecular weights, normal boiling temperatures and critical properties that were taken by Valderrama et al. [44]. Such properties only for IL [EMIM][MeSO4] were calculated by Lydersen-Joback-Reid method in this work, the calculation procedure is given in the Supporting Information. The process flowsheet is shown in Fig.1. The separation process consists of three parts, namely, an extractive distillation column (EDC), a flash tank and a stripper. The feed flowrate for all cases is 100 kmol/h, with the composition of 50 mol % of ethyl acetate and 50 mol % of ethanol taken from the reactive process by Santacesaria et al. [34]. In this work, the third stage of EDC was 4
selected as IL feed tray throughout the preliminary study to decide major factors which can affect the feasibility of IL-extractive distillation and the optimization using TAC. In the EDC of this separation process, IL increases the volatility of ethyl acetate with respect to ethanol and thus makes the separation easier. Since IL is much less volatile than either ethyl acetate or ethanol, it flows down the column to leave with the bottom product. The two ethyl acetate-ethanol mixture and the IL solvent streams are located at the EDC, with the stages above the IL solvent feed tray as the rectifying section, the stages between the feed stages of the IL solvent and the ethyl acetate-ethanol mixture as the extractive section, and the stages below the feed tray of the ethyl acetate-ethanol mixture as the stripping section, as shown in Fig. 2. The presence of IL alters the relative volatility between ethyl acetate and ethanol and makes ethyl acetate move toward the top part and ethanol move toward the bottom part in the column. In the rectifying section, there is essentially no ethanol, thus the simple separation between ethyl acetate and IL is performed with pure ethyl acetate going to the distillate and IL returning to the extractive section for solvent recovery. In the stripping section, ethyl acetate as the lightest component is stripping toward the extractive section of the column resulting in only negligible ethyl acetate in the column bottoms stream. The solvent recovery system including a flash tank and a stripping column removes ethanol from the IL. The solvent is then cooled and recycled to the EDC. The operating pressure of flash tank with the same size for all cases is 30 kPa and the stripper with 10 stages for all cases uses stripping agent with the condition of 160 ℃, 100 kmol/h air flow rate. The stripper was determined to satisfy the constraints of the final product purity of the IL solvent leaving the stripping column (specified at 99.9 mol %). Thus, the global optimization of the separation process is based on the optimization of the EDC design parameters and operating conditions, as shown in Fig. 3. To establish the operating conditions for the EDC with an IL solvent, a sensitivity analysis was performed by the main design variables, such as solvent-to-feed molar ratio (IL/F), the number of stages (NEDC), feed stage (NF), and reflux ratio (RR). 3. Results and discussion 3.1. Selection of a suitable IL For the selection of a suitable IL in the separation of ethyl acetate and ethanol, the ILs such as [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3], [EMIM][BF4], [EMIM][OAc] and 5
[EMIM][Cl] are investigated in several works [37-41]. These ILs are non-corrosive, non-toxic, and non-reactive with feed component and do not form additional azeotropes with the components in the mixture to be separated. However, as the solvent is the core of extractive distillation, more attention should be paid to the selection of a potential solvent. Kulajanpeng et al. [45] studied the screening of ILs for the design of ILs-based separation processes in various homogeneous binary azeotropic mixtures and concluded that [EMIM][OAc] was unfeasible for the separation process containing the recovery column due to its relative low degradation temperature (160℃). Zhu et al. [7] described that ILs containing the anion [Cl] group are not favorable for a continuous process because
of
its
high
viscosity.
Thus,
four
ILs
[EMIM][MeSO3],
[EMIM][MeSO4],
[BMIM][CF3SO3] and [EMIM][BF4] were preliminary collected as a set of feasible solvents for the separation process of ethyl acetate and ethanol. The property of relative volatility is important to select the suitable solvent in the extractive distillation. To verify the effect of the four ILs, the VLE behaviors of the ternary mixtures, on a solvent-free basis, were simulated and plotted in Fig. 4. We can see that IL reverses the volatility of the original mixture, that is, enhances the relative volatility of ethyl acetate to ethanol. All of the four ILs studied in this work were able to break the ethyl acetate and ethanol azeotrope. It also indicates that the UNIFAC-Lei model which can be used to predict the VLE behavior of mixtures containing IL based on the semi-empirical method and the NRTL model which can be used to reproduce the VLE behavior based on the experimental data yield similar results. The comparisons between relative volatilities of key component ethyl acetate when using different ILs are shown in Fig. 5. The values of relative volatility in the whole comparison region of ethyl acetate are greater than 1, which means separating of ethyl acetate and ethanol using these four ILs as solvents are feasible. [EMIM][MeSO3] has the largest separation factor values at the given concentration ranges of ethyl acetate. It can be seen that the separation ability of [EMIM][MeSO3] is superior to other ILs because preferential affinity of ethanol and [EMIM][MeSO3] by hydrogen bonding is greater than that of other ILs from the analysis of relative volatility values. The results on the relative volatility predicted by the UNIFAC-Lei model in this work are corresponding to the discussion based on experimental data reported by Andreatta et al. [37]. Thus, [EMIM][MeSO3] was used as the suitable solvent to investigate the feasibility of IL-extractive distillation and process optimization separating 6
ethyl acetate from ethanol. 3.2. Feasibility of the extractive distillation using [EMIM][MeSO3] The residue curve map (RCM) for this three-component system including ethyl acetate-ethanol mixture and [EMIM][MeSO3] is shown in Fig. 6. From the RCM, we can see that the ethyl acetate-ethanol azeotrope is the unstable node, [EMIM][MeSO3] is the stable node, and both ethyl acetate and ethanol are the saddles. For the isovolatility curve, the relative volatility between ethyl acetate and ethanol was calculated in the presence of [EMIM][MeSO3] and the unit values of the relative volatility at different compositions were collected. The separation of ethyl acetate from its mixture with ethanol adding [EMIM][MeSO3] illustrates the case where the isovolatility curve reaches the binary side of ethyl acetate-ethanol. It is noticed that the ethyl acetate-ethanol mixture can be separated easier in the presence of [EMIM][MeSO3] because of higher values of the relative volatility. Also, the feed composition of 50 mol % of ethyl acetate taken from the reactive system is located at feasible region overcoming the unstable node and the part under the isovolatility curve as seen in Fig. 6. As shown in Fig. 7, the essential features of feasibility for the separation of ethyl acetate and ethanol can be observed from the liquid or vapor phase behavior in the EDC. For a feasible column, the profiles of the rectifying and the stripping section must be connected by the profile of the extractive middle section. Figs. 7a and 7b show the composition profiles along the EDC. The profile of ethyl acetate composition in the liquid phase displays a pronounced slope in the stripping zone and is smooth in the rectifying zone. And ethyl acetate obtained throughout the first stage as the condenser is maintained at almost 1 mole fraction of the purity. Between stage 37 and 38, there is a slope change due to the introduction of the azeotropic mixture. The profile of ethanol composition in the liquid phase quickly decreases along the stripping section and turns out to be smoother in the rectifying section. The profile of [EMIM][MeSO3] composition in the liquid phase decrease between stage 37 and 38 due to the mixture feed stream; in most of the extractive section, it is kept approximately constant, since the IL vapor pressure was neglected. As shown in Fig. 7b, the vapor composition of IL is approximately zero in all stages of the EDC. On the other hand, the ethyl acetate and ethanol vapor phase composition displays monotonic increasing and decreasing behavior, respectively. The temperature profile (Fig. 7c) along the EDC displays an increase between stage 3 and 4 7
due to mixing heats which are originated by the contact between the azeotropic mixture fed in stage 3 and the IL coming down from the stage above. It is clear that the presence of the IL in the reboiler increases the boiling temperature of the ethanol and IL mixture. Since the IL is not present in the condenser because of its negligible vapor pressure, the temperature in this part is close to the ethyl acetate normal boiling temperature. Consequently, it can be inferred that the extractive separation of ethyl acetate from ethanol in the presence of [EMIM][MeSO 3] is successfully carried out in the extractive section. From the analysis of RCM and liquid/vapor behaviors of each components in the EDC, it is obvious that the extractive distillation using [EMIM][MeSO3] as solvent is sufficient to purify ethyl acetate with the desired purity. 3.3. Inflence of different design parameters on the purity of ethyl acetate and energy requirement The influence of the number of stages and the feed stage on the molar composition of ethyl acetate in the distillate and the reboiler duties were analyzed. Fig. 8a shows the relationship between the ethyl acetate purity in the distillate and the number of stages and feed stages, and results indicate that some cases with the special number stages and feed stages can be obtained ethyl acetate with 99.9 mol % purity in the overhead product. Fig. 8b shows the influence of the number of stages and feed stage on reboiler duties, results show that reboiler duties are immediately increased when the feed stage is close to the bottom. The influence of the IL/F and reflux ratio on the molar composition of ethyl acetate in the distillate and the reboiler duties were also analyzed, as shown in Figs. 8c and 8d. The IL/F is one of important parameters for the improvement of the ethyl acetate purity in the separation process. Fig. 8c shows that IL/F effects directly on the distillate purity without considerably affecting the energy consumption. Moreover, Fig. 8d shows that there is no significant change in energy consumption at a constant reflux ratio for different IL/F values within the interval 0.15 - 0.40. Similarly, at a constant IL/F the reflux ratio increases until a distillate composition equivalent to the one obtained in the previous variation is reached. From Fig. 8d the increase in energy consumption is 22.9 %, so the reboiler energy consumption in the extractive distillation for the separation of ethyl acetate and ethanol is obviously related to the reflux ratio. Thus, the reflux ratio must be set to the lowest possible value so that the IL/F ratio can be manipulated to reach the desire distillate purity without high energy consumption. 3.4. Process design verification through economic anaylsis 8
According to sensitivity analysis results shown in Fig. 8, it is obvious that the separation of ethyl acetate and ethanol mixture with [EMIM][MeSO3] is possible in some cases with different design parameters when the product purity specifications are fixed. Therefore, to determine the best option of these process design parameters it is necessary to carry out an economic evaluation based on the minimum TAC. The cost models for TAC calculation by Luyben [46] are shown in Appendix. The economics of the extractive distillation system are evaluated in terms of capital investment and energy costs. Fig. 9 shows the comparison of TACs of the EDC with different design parameters while the ethyl acetate purity is fixed at 99.9 mol %. TACs are no significant change along the different feed locations providing the desired purity. TACs of the cases with 50 stages are least in comparison to the other cases as seen in Fig. 9. Fig. 10 shows the optimal conditions and results of the ethyl acetate + ethanol azeotropic separation process when RR is fixed at 2.0. Solvent recovery process required a flash tank with the operating temperature of 117 ℃ and the heat duty of 511 kW at the condition of 30 kPa, and a stripper with 10 stages and the pressure of 100 kPa. Ethanol can be recovered from air + ethanol stream by condensation or absorption using water for the amount of ethanol is not high. Ethanol obtained from the flash tank and the stripper can be reused by sending to the reactive system for the dehydrogenation of ethanol. Furthermore, the design parameters of three cases and the corresponding costs of all units are summarized in Table 5. Solvent recovery system consisting of flash tank and stripper occupies 22 % ~ 24 % of the total TAC, so it can be sure that the process design is decided by EDC optimization in the separation process for ethyl acetate purification using [EMIM][MeSO3] as the solvent. 4. Conclusion This work focused on the extractive distillation process for the separation of ethyl acetate and ethanol mixture using ILs as solvents. Four ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] were employed as a set of feasible solvents and the UNIFAC-Lei thermodynamic model was applied to find the favorable IL solvent. The whole extractive distillation process including solvent recovery system of a flash tank and a stripper was conducted in Aspen Plus. The sensitivity analysis was performed to investigate the influence of various operating parameters on the separation target with the selection of a suitable IL. From the results, [EMIM][MeSO3] was selected as the proposed solvent in the separation process for the 9
ethyl acetate purification. Ethanol withdrawn from the solvent recovery system can be reused by sending to the reactive system for the dehydrogenation of ethanol. The separation process configuration and operating parameters were confirmed on the basis of the EDC optimization by TAC. The results show that the TAC of the optimal process is $745,380/y and the solvent recovery system occupies 24 % of the total cost of this process. Consequently, the optimization problem of the extractive distillation system by the selection of a favorable IL solvent is the main strategy that can allow economically efficient separation of close-boiling mixtures such as the ethyl acetate-ethanol system. Acknowledgement Financial support from National Natural Science Foundation of China (Project 21676152) is gratefully acknowledged. References [1] A.A. Kiss, Distillation technology–still young and full of breakthrough opportunities, J. Chem. Technol. Biotechnol. 89 (2014) 479–498. [2] N.V.D. Long, M. Lee, Optimal retrofit design of extractive distillation to energy efficient thermally coupled distillation scheme, AlChE J. 59 (2013) 1175–1182. [3] S. Liang, Y. Cao, X. Liu, X. Li, Y. Zhao, Y. Wang, Y. Wang, Insight into pressure-swing distillation from azeotropic phenomenon to dynamic control, Chem. Eng. Res. Des. 117 (2017) 318-335. [4] Y. Cao, J. Hu, H. Jia, G. Bu, Z. Zhu, Y. Wang, Comparison of pressure-swing distillation and extractive distillation with varied-diameter column in economics and dynamic control, J. Process Contr. 49 (2017) 9-25. [5] Z. Lei, C. Dai, J. Zhu, B. Chen, Extractive distillation with ionic liquids: A review, AlChE J. 60 (2014) 3312–3329. [6] H. Rodríguez, Ionic Liquids in the Context of Separation Processes, Springer Berlin Heidelberg, (2016) 1-9. [7] Z. Zhu, Y. Ri, M. Li, H. Jia, Y. Wang, Y. Wang, Extractive distillation for ethanol dehydration using imidazolium-based ionic liquids as solvents, Chem. Eng. Proc. 109 (2016) 190-198. [8] Z. Lei, R. Zhou, Z. Duan, Process improvement on separating C4 by extractive distillation, Chem. Eng. J. 85 (2002) 379-386. [9] S. Li, H. Liu, Entrainer Screening for Separation of Isoparaffins and Naphthenes by Extractive Distillation, Ind. Eng. Chem. Res. 48 (2009) 5427-5430. [10] Z.G. Zhang, D.H. Huang, M. Lv, P. Jia, D.Z. Sun, W.X. Li, Entrainer selection for separating tetrahydrofuran/water azeotropic mixture by extractive distillation, Sep. Purif. Technol. 122 (2014) 73-77. [11] N. Medinaherrera, I.E. Grossmann, M.S. Mannan, A. Jiménezgutiérrez, An Approach for Solvent Selection in Extractive Distillation Systems Including Safety Considerations, Ind. Eng. Chem. Res. 53 (2014) 12023-12031. [12] W.B. Ramos, M.F.D. Figueirêdo, K.D. Brito, S. Ciannella, L.G.S. Vasconcelos, R.P. Brito, Effect 10
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[31] I. Díaz, J. Palomar, M. Rodríguez, J.D. Riva, V. Ferro, E.J. González, Ionic Liquids as Entrainers for the Separation of Aromatic-Aliphatic Hydrocarbon Mixtures by Extractive Distillation, Chem. Eng. Res. Des. 115 (2016) 382-393. [32] G.W. Meindersma, A.B.D. Haan, Conceptual process design for aromatic/aliphatic separation with ionic liquids, Chem. Eng. Res. Des. 86 (2008) 745-752. [33] P. Nancarrow, N. Mustafa, A. Shahid, V. Varughese, U. Zaffar, R. Ahmed, N. Akther, H. Ahmed, I. Alzubaidy, S. Hasan, Technical Evaluation of Ionic Liquid-Extractive Processing of Ultra Low Sulfur Diesel Fuel, Ind. Eng. Chem. Res. 54 (2015) 10843-10853. [34] E. Santacesaria, G. Carotenuto, R. Tesser, M.D. Serio, Ethanol dehydrogenation to ethyl acetate by using copper and copper chromite catalysts, Chem. Eng. J. 179 (2012) 209-220. [35] L. Berg, Separation of ethyl acetate from ethanol by azeotropic distillation, US 1999. [36] I. Nieuwoudt, B. Van Dyk, Separation of ethanol mixtures by extractive distillation, US 2002. [37] A.E. Andreatta, M.P. Charnley, J.F. Brennecke, Using Ionic Liquids to Break the Ethanol - Ethyl Acetate Azeotrope, ACS Sustain. Chem. Eng. 3 (2015) 3435-3444. [38] R. Li, X. Cui, Y. Zhang, T. Feng, J. Cai, Vapor-liquid equilibrium and liquid-liquid equilibrium of ethyl acetate + ethanol +1-ethyl-3-methylimidazolium acetate, J. Chem. Eng. Data 57 (2012) 911-917. [39] D.L. Zhang, Y. F. Deng, C.B. Li, J. Chen, Separation of Ethyl Acetate−Ethanol Azeotropic Mixture Using Hydrophilic Ionic Liquids, Ind. Eng. Chem. Res. 47 (2008) 1995-2001. [40] Q. Li, J. Zhang, Z. Lei, J. Zhu, F. Xing, Isobaric Vapor−Liquid Equilibrium for Ethyl Acetate + Ethanol + 1-Ethyl-3-methylimidazolium Tetrafluoroborate, J. Chem. Eng. Data 54 (2009) 193-197. [41] Q. Li, J. Zhang, Z. Lei, J. Zhu, J. Zhu, X. Huang, Selection of Ionic Liquids as Entrainers for the Separation of Ethyl Acetate and Ethanol, Ind. Eng. Chem. Res. 48 (2009) 9006-9012. [42] R. Kato, J. Gmehling, Systems with ionic liquids: Measurement of VLE and γ∞ data and prediction of their thermodynamic behavior using original UNIFAC, mod. UNIFAC(Do) and COSMO-RS(Ol), J. Chem. Thermodyn. 37 (2005) 603-619. [43] E.I. Alevizou, G.D. Pappa, E.C. Voutsas, Prediction of phase equilibrium in mixtures containing ionic liquids using UNIFAC, Fluid Phase Equilib. 284 (2009) 99-105. [44] J.O. Valderrama, R.E. Rojas, Critical Properties of Ionic Liquids. Revisited, Ind. Eng. Chem. Res. 48 (2009) 6890-6900. [45] K. Kulajanpeng, U. Suriyapraphadilok, R. Gani, Systematic screening methodology and energy efficient design of ionic liquid-based separation processes, J. Clean. Prod. 111 (2015) 93-107. [46] W.L. Luyben, C.C. Yu, Reactive Distillation Design and Control, Jhon Wiley & Sons, 2008. [47] D.N. Kozlov, J. Kiefer, T. Seeger, A.P. Fröba, A. Leipertz, Simultaneous measurement of speed of sound, thermal diffusivity, and bulk viscosity of 1-ethyl-3-methylimidazolium-based ionic liquids using laser-induced gratings, J. Phys. Chem. B 118 (2014) 14493-14501. [48] T. Kavitha, T. Vasantha, P. Venkatesu, R.S.R. Devi, T. Hofman, Thermophysical properties for the mixed solvents of N-methyl-2-pyrrolidone with some of the imidazolium-based ionic liquids, J. Mol. Liq. 198 (2014) 11-20. [49] M. Shamsipur, A.A.M. Beigi, M. Teymouri, S.M. Pourmortazavi, M. Irandoust, Physical and electrochemical
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of
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liquids
1-ethyl-3-methylimidazolium
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1-butyl-1-methylpyrrolidinium
bis(trifluoromethylsulfonyl)imide, J. Mol. Liq. 157 (2010) 43-50. [50] X. Li, W. Sun, G. Wu, L. He, H. Li, H. Sui, Ionic Liquid Enhanced Solvent Extraction for Bitumen Recovery from Oil Sands, Energy Fuels 25 (2011) 5224-5231. 12
Fig. 1. Conceptual process flowsheet of the separation for ethyl acetate and ethanol mixture using IL as solvent.
13
Fig. 2. Schematic drawing to the EDC with IL as solvent.
14
Fig. 3. Flowsheet of the EDC optimization by sensitivity analysis and TAC calculation.
15
Fig. 4. Effect of different ILs on the ethyl acetate - ethanol VLE at xIL = 0.2. The dash line represents VLE curve without IL.
16
Fig. 5. Relative volatilities for ethyl acetate to ethanol using different ILs: , [EMIM][MeSO 3]; , [EMIM][MeSO4]; , [BMIM][CF3SO3]; , [EMIM][BF4].
17
Fig. 6. RCM with the isovolatility curve of ethyl acetate and ethanol using [EMIM][MeSO 3] as solvent.
18
Fig. 7. Composition and temperature profiles of an EDC for the separation of ethyl acetate and ethanol using [EMIM][MeSO3]: (a) liquid; (b) vapor; (c) temperature.
19
Fig. 8. Influence of several process design parameters on distillate composition and reboiler duties: (a) influence of the number of stages and feed stages on the ethyl acetate purity; (b) influence of the number of stages and feed stage on reboiler duties; (c) influence of IL solvent flow rate and reflux ratio on the ethyl acetate purity; (d) influence of IL solvent flow rate and reflux ratio on reboiler duties.
20
Fig. 9. Influence of number of stages and feed stages on TAC.
21
Fig. 10. The optimal conditions and results of the ethyl acetate + ethanol azeotropic separation process using [EMIM][MeSO3] at RR = 2.0.
22
Table 1. Group division of ILs ILs
Group division
[EMIM][MeSO3]
[MIM][MeSO3] + CH2 + CH3
[EMIM][MeSO4]
[MIM][MeSO4] + CH2 + CH3
[BMIM][CF3SO3]
[MIM][CF3SO3] + 3CH2 + CH3
[EMIM][BF4]
[MIM][BF4] + CH2 + CH3
23
Table 2. Parameters of Rk and Qk Group
Rk
Qk
CH3a
0.9011
0.8480
CH2a
0.6744
0.5400
1.0000
1.2000
a
OH
a
CH3COO
1.9031
1.7280
a
7.9863
5.5761
b
3.7481
3.0714
9.5357
5.0500
6.5669
4.0050
[MIM][MeSO3] [MIM][MeSO4]
b
[MIM][CF3SO3] b
[MIM][BF4] a
Reference [26]; b Reference [27]
24
Table 3. Group interaction parameters for the UNIFAC model m
n
CH2
[MIM][MeSO3]
661.05a
544.93a
CH2
[MIM][MeSO4]
575.45b
473.80b
CH2
[MIM][CF3SO3]
405.39b
284.37b
CH2
[MIM][BF4]
1108.51b
588.74 b
OH
[MIM][MeSO3]
1.14a
-389.86a
OH
[MIM][MeSO4]
13.83b
-570.26b
OH
[MIM][CF3SO3]
34.90b
-305.24b
OH
[MIM][BF4]
-13.77b
131.24b
a
-41.76a
CH3COO
[MIM][MeSO3]
767.31
CH3COO
[MIM][MeSO4]
n.a.
CH3COO
[MIM][CF3SO3]
n.a.
CH3COO a
[MIM][BF4]
32.09
n.a. n.a. a
112.17a
Reference [26]; b Reference [27]
25
Table 4. Several properties of ILs, which were used to operate the simulation with IL component in Aspen Plus Mw
Tb(K)
Tc(K)
Pc(bar)
Vc(cm3/mol)
zc
ω
η(cP)
ρ(g/cm3)
[EMIM][MeSO3]
206.3a
667.4a
1026.0a
48.1a
587.1a
0.3312a
0.3307a
86.0 (28℃)c
1.24c
[EMIM][MeSO4]
222.3b
602.7b
1053.6b
45.9b
602.7b
0.3115b
0.3401b
84.3 (25℃)d
1.28d
[BMIM][CF3SO3]
288.3a
707.7a
1023.5a
29.5a
750.7a
0.2600a
0.4046a
77.0 (25℃)e
1.31e
[EMIM][BF4]
198.0a
449.5a
596.2a
23.6a
540.8a
0.2573a
0.8087a
66.5 (20℃)f
1.25f
ILs
a
Reference [44]; b this work; c Reference [47]; d Reference [48]; e Reference [49]; f Reference [50].
26
Table 5. TAC of separation process for ethyl acetate purification Case 1
Case 2
Case 3
45/3/37
50/3/38
55/3/45
ID1 (m)
1.12
1.03
1.03
RR
2.5
2.0
2.0
1568
1343
1344
1789
1565
1566
1075.01
1032.08
1081.35
413.83
361.86
362.06
628.84
568.28
578.33
0.90
0.90
0.90
511
511
511
capital cost (×10 $)
169.13
169.13
169.13
3
115.98
115.98
115.98
149.80
149.80
149.80
10
10
10
0.652
0.652
0.652
136.46
136.46
136.46
27.3
27.3
27.3
805.94
745.38
755.43
EDC NEDC/NIL/NF
QC (kW) QR (kW) 3
capital cost (×10 $) 3
operating cost (×10 $) 3
TAC1 (×10 $) flash tank ID2 (m) QH (kW) 3
operating cost (×10 $) 3
TAC2 (×10 $) stripper Nstripper ID3 (m) 3
capital cost (×10 $) TAC3 (×103 $) 3
TAC (×10 $)
27
Appendix. Basis of economics Column vessel (diameter and length in meters) Length: L = 1.2 × 0.61 × (N – 2) Capital cost = 17640 (D)1.066 (L)0.802 Cooler (area in m2) Heat-transfer coefficient (KC) = 0.852kW K-1 m-2 Capital cost = 7296 (area)0.65 Reboiler (area in m2) Heat-transfer coefficient (KR) = 0.568 kW K-1 m-2 Capital cost = 7296 (area)0.65 Flash tank Liquid volume fraction: 0.5 m3 Residence time: 20min
Volume = D=
3
1 moleflow 20 2 60
2 volume , L = 2D
Capital cost = 17640 (D)1.066 (L)0.802
TAC =
capital cos t + operating cost payback period
Payback period = 5 years
28
Nomenclature ETAC
Ethyl acetate
ETOH
Ethanol
[EMIM][MeSO3]
1-Ethyl-3-methylimidazolium methanesulfonate
[EMIM][MeSO4]
1-Ethyl-3-methylimidazolium methylsulfate
[BMIM][CF3SO3]
1-Butyl-3-methylimidazolium trifluoromethanesulfonate
[EMIM][BF4]
1-Ethyl-3-methylimidazolium tetrafluoroborate
[EMIM][OAc]
1-Ethyl-3-methylimidazolium acetate
[EMIM][Cl]
1-Ethyl-3-methylimidazolium chloride
Mw
Molecular weight
Tb
Normal boiling temperature (K)
Tc
Critical temperature (K)
Pc
Critical pressure (bar)
Vc
Critical molecular volume (cm3/mol)
Zc
Critical compressibility factor
w
Acentric factor
A.P.
Azeotropic point
αi/j
Relative volatility of ethyl acetate to ethanol
NEDC
Number of stages of extractive distillation column
NIL
Solvent feed stage of extractive distillation column
NF
Feed stage of extractive distillation column
Nstripper
Number of stages of stripper
RR
Reflux ratio of extractive distillation column
ID1
Diameter of extractive distillation column
ID2
Diameter of flash tank
ID3
Diameter of stripper
QC
Heat duty of cooler (kW)
QR
Heat duty of reboiler (kW)
QH
Heat duty of flash tank (kW)
29
Highlights VLE behaviors of ternary mixtures ethyl acetate-ethanol-IL are predicted by UNIFAC model. The separation process using [EMIM][MeSO3] as solvent is studied. The optimization of the extractive distillation column using IL becomes a main strategy.
30
S1383-5866(17)30026-6 http://dx.doi.org/10.1016/j.seppur.2017.03.011 SEPPUR 13597
To appear in:
Separation and Purification Technology
Received Date: Revised Date: Accepted Date:
3 January 2017 5 March 2017 9 March 2017
Please cite this article as: Z. Zhu, Y. Ri, H. Jia, X. Li, Y. Wang, Y. Wang, Process Evaluation on the Separation of Ethyl acetate and Ethanol using Extractive Distillation with Ionic Liquid, Separation and Purification Technology (2017), doi: http://dx.doi.org/10.1016/j.seppur.2017.03.011
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.
Process Evaluation on the Separation of Ethyl acetate and Ethanol using Extractive Distillation with Ionic Liquid Zhaoyou Zhua, Yongsaeng Ria, b, Hui Jiaa, Xin Lia, Yong Wanga, Yinglong Wanga, * a
College of Chemical Engineering, Qingdao University of Science and Technology, Qingdao
266042, China b
Faculty of Chemistry, Kim Il Sung University, Pyongyang 999093, DPR of Korea
Corresponding Author *E-mail: [email protected]. Abstract This paper provides process design and simulation methodology for the separation of ethyl acetate and ethanol by extractive distillation using ionic liquids (ILs) as solvents and obtains the design parameters of extractive distillation process with solvent recovery system on the basis of a suitable IL solvent. The feasibility of IL-extractive distillation was examined via process simulation for the separation of ethyl acetate and ethanol in Aspen Plus. Four ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] were created in Aspen Plus database with several thermodynamic and physical property parameters to allow the process to be simulated via the UNIFAC-Lei thermodynamic method. The results show that the separation process containing the hybrid regeneration system of flash tank and stripper with [EMIM][MeSO3] as a suitable solvent is the best option in extractive distillation for the separation of ethyl acetate and ethanol compared with other ILs from the analysis of relative volatilities. The separation process was optimized by sensitivity analysis and the optimal design parameters were verified by economic evaluation based on the total annual cost (TAC). Keywords: Extractive distillation; Ionic liquid; UNIFAC-Lei; VLE; Simulation 1. Introduction In the history of chemical separations, conventional distillation has been applied to more commercial processes than all other techniques combined [1]. However, for systems with close boiling point or azeotropic systems, a separation by conventional distillation process becomes difficult or even impossible [2-4]. Nowadays, extractive distillation with ILs as solvents has become a promising alternative for these systems such as ethanol dehydration [5-7]. Extractive distillation is known as an efficient technology in the separation of complex mixtures by the 1
addition of a solvent at the top of the column that modifies the activity coefficients in the liquid phase to increase the relative volatilities [8-10]. The effectiveness of an extractive distillation process relies on the choice of solvent [11-13]. Since Lei’s group [14] first reported the extractive distillation process with ILs as green solvents which possess the advantages of a liquid solvent (easy operation) and a solid salt (high separation ability) in 2003, ILs are under intensive investigation to determine their potential as replacement solvents for extractive distillation. ILs have the outstanding physicochemical properties of low vapor pressure, low melting point, and high chemical and thermal stability [15-19]. So ILs are suggested by many authors for a broad range of separations, but the experimental way to determine the thermodynamic properties of ILs and the capability of the preferential IL solvent is a time consuming task because a large number of ILs exist [20-23]. The application of predictive models to estimate the thermodynamic properties of ILs and their mixture with water or organic compounds is of great interest. Indeed, investigation on predicting ILs properties by UNIFAC-Lei and COSMO-RS method gained focus in recent [24-28]. Selection of property methods depends on the non-ideal degree and operating conditions of the system, and a suitable property method determines the accuracy of the simulation results such as operating temperature and pressure of the system. In the works that the conceptual design of separation processes using ILs was carried out [7, 29-32], most of the property methods selected to the separation of azeotropic mixtures is the NRTL model. Nancarrow et al. [33] reported the simulation method of IL-based desulfurization processes by UNIFAC-Lei model in Aspen Plus environment and the feasibility of industrial scale IL-extractive processing of the ultra low sulfur diesel. However, to our knowledge, there is no study on the separation process of the azeotropic mixture by IL-extractive distillation simulation using UNIFAC-Lei as property method. In this paper, the UNIFAC-Lei model was employed to separate the close-boiling mixture of ethyl acetate and ethanol mixture using ILs as solvents. Santacesaria et al. [34] described that the production process of ethyl acetate by the dehydrogenation of ethanol has become very attractive considering the growing production of bio-ethanol for the fuel market, and also this reaction is simple, noncorrosive, less-toxic process and requiring only an ethanol feedstock. But the separation of ethyl acetate and ethanol mixture produced from the reaction system is among the difficult separation processes due to their close-boiling points and formation of an azeotropic 2
mixture [35], as boiling points of ethyl acetate and ethanol are 78.31℃ and 77.20℃, respectively. The organic solvents currently used in extractive distillations for separating ethyl acetate from ethanol are an amine (e.g., diethylene triamine), an alkylated thiopene (e.g., ethyl thiopene), and paraffins (e.g., dodecane) [36]. Recently, the ability of extractive distillation by ILs in facilitating the separation of azeotropic mixture of ethyl acetate - ethanol has been experimentally illustrated in some works [37-41]. However, the separation process with the consideration of effects between IL solvent and design parameters and the ability of solvent recovery for the continuous process have not been studied. The purpose of this article is to separate ethyl acetate and ethanol mixture by extractive distillation process based on the selection of a suitable IL solvent. Firstly, the vapor-liquid equilibrium (VLE) behaviors of ethyl acetate and ethanol with different ILs were described by the UNIFAC-Lei model and the separation effects were compared by relative volatilities of ethyl acetate when using ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] as solvents. The four ILs were preliminary selected as a set of study object which owned the ability of eliminating ethyl acetate and ethanol azeotrope in the relatively low concentrations among ILs reported by several literatures [37-41]. In addition, an IL that has the best performance among the four ILs in separating the ethyl acetate and ethanol mixture was employed to seek the conditions and configuration for the separation process by sensitivity analysis. TAC was calculated for the different configurations producing ethyl acetate with the distillate purity of 99.9 mol % to confirm a novel separation process configuration. 2. Method description 2.1. The UNIFAC-Lei model for IL-based systems UNIFAC-Lei computes the activity coefficient of liquid mixtures in terms of constants which reflects the size and surface area of individual functional groups present in the mixture. If the necessary binary group interaction parameters exist for every pair of functional groups of the mixture, UNIFAC-Lei model is worthy of being considered to predict the phase behavior [26, 27, 42, 43]. The activity coefficient can be calculated by
ln i ln i C ln i R 3
(1)
where
represents the combinatorial contribution to the activity coefficient due to differences
in the sizes and shapes of the groups, and energetic interactions. The part
represents the residual contribution from
contains the group parameters Rk and Qk, which values for
the groups investigated in this work can be found in the literature, as listed in Table 1. The part is a function of binary interaction parameters between groups, which group interaction parameter
is expressed as follows:
nm exp anm / T where T is the absolute temperature, of functional groups n and m, and
(2)
is the temperature-independent parameter for each pair .
ILs have to be transformed to functional group when using the UNIFAC-Lei model, and the fragmentations for ILs used in this work are listed in Table 2. ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] have been constructed from the new [MIM][MeSO3], [MIM][MeSO4], [MIM][CF3SO3] and [MIM][BF4] groups combined with the established CH2 and CH3 groups existed in the UNIFAC-Lei model. The group interaction parameters (
) between each pair of functional groups in their mixture with ethyl
acetate and ethanol taken from Lei et al. [26, 27] are listed in Table 3. 2.2. Process simulation For these three-component systems, the UNIFAC-Lei model was used to describe the nonideality of the liquid phase while the vapor phase is assumed to be ideal. The extractive distillation processes containing ILs were simulated using Aspen Plus v8.4 with the rigorous equilibrium stage model RadFrac. New component ILs were created as user-defined components in Aspen Plus by specifying their molecular weights, normal boiling temperatures and critical properties that were taken by Valderrama et al. [44]. Such properties only for IL [EMIM][MeSO4] were calculated by Lydersen-Joback-Reid method in this work, the calculation procedure is given in the Supporting Information. The process flowsheet is shown in Fig.1. The separation process consists of three parts, namely, an extractive distillation column (EDC), a flash tank and a stripper. The feed flowrate for all cases is 100 kmol/h, with the composition of 50 mol % of ethyl acetate and 50 mol % of ethanol taken from the reactive process by Santacesaria et al. [34]. In this work, the third stage of EDC was 4
selected as IL feed tray throughout the preliminary study to decide major factors which can affect the feasibility of IL-extractive distillation and the optimization using TAC. In the EDC of this separation process, IL increases the volatility of ethyl acetate with respect to ethanol and thus makes the separation easier. Since IL is much less volatile than either ethyl acetate or ethanol, it flows down the column to leave with the bottom product. The two ethyl acetate-ethanol mixture and the IL solvent streams are located at the EDC, with the stages above the IL solvent feed tray as the rectifying section, the stages between the feed stages of the IL solvent and the ethyl acetate-ethanol mixture as the extractive section, and the stages below the feed tray of the ethyl acetate-ethanol mixture as the stripping section, as shown in Fig. 2. The presence of IL alters the relative volatility between ethyl acetate and ethanol and makes ethyl acetate move toward the top part and ethanol move toward the bottom part in the column. In the rectifying section, there is essentially no ethanol, thus the simple separation between ethyl acetate and IL is performed with pure ethyl acetate going to the distillate and IL returning to the extractive section for solvent recovery. In the stripping section, ethyl acetate as the lightest component is stripping toward the extractive section of the column resulting in only negligible ethyl acetate in the column bottoms stream. The solvent recovery system including a flash tank and a stripping column removes ethanol from the IL. The solvent is then cooled and recycled to the EDC. The operating pressure of flash tank with the same size for all cases is 30 kPa and the stripper with 10 stages for all cases uses stripping agent with the condition of 160 ℃, 100 kmol/h air flow rate. The stripper was determined to satisfy the constraints of the final product purity of the IL solvent leaving the stripping column (specified at 99.9 mol %). Thus, the global optimization of the separation process is based on the optimization of the EDC design parameters and operating conditions, as shown in Fig. 3. To establish the operating conditions for the EDC with an IL solvent, a sensitivity analysis was performed by the main design variables, such as solvent-to-feed molar ratio (IL/F), the number of stages (NEDC), feed stage (NF), and reflux ratio (RR). 3. Results and discussion 3.1. Selection of a suitable IL For the selection of a suitable IL in the separation of ethyl acetate and ethanol, the ILs such as [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3], [EMIM][BF4], [EMIM][OAc] and 5
[EMIM][Cl] are investigated in several works [37-41]. These ILs are non-corrosive, non-toxic, and non-reactive with feed component and do not form additional azeotropes with the components in the mixture to be separated. However, as the solvent is the core of extractive distillation, more attention should be paid to the selection of a potential solvent. Kulajanpeng et al. [45] studied the screening of ILs for the design of ILs-based separation processes in various homogeneous binary azeotropic mixtures and concluded that [EMIM][OAc] was unfeasible for the separation process containing the recovery column due to its relative low degradation temperature (160℃). Zhu et al. [7] described that ILs containing the anion [Cl] group are not favorable for a continuous process because
of
its
high
viscosity.
Thus,
four
ILs
[EMIM][MeSO3],
[EMIM][MeSO4],
[BMIM][CF3SO3] and [EMIM][BF4] were preliminary collected as a set of feasible solvents for the separation process of ethyl acetate and ethanol. The property of relative volatility is important to select the suitable solvent in the extractive distillation. To verify the effect of the four ILs, the VLE behaviors of the ternary mixtures, on a solvent-free basis, were simulated and plotted in Fig. 4. We can see that IL reverses the volatility of the original mixture, that is, enhances the relative volatility of ethyl acetate to ethanol. All of the four ILs studied in this work were able to break the ethyl acetate and ethanol azeotrope. It also indicates that the UNIFAC-Lei model which can be used to predict the VLE behavior of mixtures containing IL based on the semi-empirical method and the NRTL model which can be used to reproduce the VLE behavior based on the experimental data yield similar results. The comparisons between relative volatilities of key component ethyl acetate when using different ILs are shown in Fig. 5. The values of relative volatility in the whole comparison region of ethyl acetate are greater than 1, which means separating of ethyl acetate and ethanol using these four ILs as solvents are feasible. [EMIM][MeSO3] has the largest separation factor values at the given concentration ranges of ethyl acetate. It can be seen that the separation ability of [EMIM][MeSO3] is superior to other ILs because preferential affinity of ethanol and [EMIM][MeSO3] by hydrogen bonding is greater than that of other ILs from the analysis of relative volatility values. The results on the relative volatility predicted by the UNIFAC-Lei model in this work are corresponding to the discussion based on experimental data reported by Andreatta et al. [37]. Thus, [EMIM][MeSO3] was used as the suitable solvent to investigate the feasibility of IL-extractive distillation and process optimization separating 6
ethyl acetate from ethanol. 3.2. Feasibility of the extractive distillation using [EMIM][MeSO3] The residue curve map (RCM) for this three-component system including ethyl acetate-ethanol mixture and [EMIM][MeSO3] is shown in Fig. 6. From the RCM, we can see that the ethyl acetate-ethanol azeotrope is the unstable node, [EMIM][MeSO3] is the stable node, and both ethyl acetate and ethanol are the saddles. For the isovolatility curve, the relative volatility between ethyl acetate and ethanol was calculated in the presence of [EMIM][MeSO3] and the unit values of the relative volatility at different compositions were collected. The separation of ethyl acetate from its mixture with ethanol adding [EMIM][MeSO3] illustrates the case where the isovolatility curve reaches the binary side of ethyl acetate-ethanol. It is noticed that the ethyl acetate-ethanol mixture can be separated easier in the presence of [EMIM][MeSO3] because of higher values of the relative volatility. Also, the feed composition of 50 mol % of ethyl acetate taken from the reactive system is located at feasible region overcoming the unstable node and the part under the isovolatility curve as seen in Fig. 6. As shown in Fig. 7, the essential features of feasibility for the separation of ethyl acetate and ethanol can be observed from the liquid or vapor phase behavior in the EDC. For a feasible column, the profiles of the rectifying and the stripping section must be connected by the profile of the extractive middle section. Figs. 7a and 7b show the composition profiles along the EDC. The profile of ethyl acetate composition in the liquid phase displays a pronounced slope in the stripping zone and is smooth in the rectifying zone. And ethyl acetate obtained throughout the first stage as the condenser is maintained at almost 1 mole fraction of the purity. Between stage 37 and 38, there is a slope change due to the introduction of the azeotropic mixture. The profile of ethanol composition in the liquid phase quickly decreases along the stripping section and turns out to be smoother in the rectifying section. The profile of [EMIM][MeSO3] composition in the liquid phase decrease between stage 37 and 38 due to the mixture feed stream; in most of the extractive section, it is kept approximately constant, since the IL vapor pressure was neglected. As shown in Fig. 7b, the vapor composition of IL is approximately zero in all stages of the EDC. On the other hand, the ethyl acetate and ethanol vapor phase composition displays monotonic increasing and decreasing behavior, respectively. The temperature profile (Fig. 7c) along the EDC displays an increase between stage 3 and 4 7
due to mixing heats which are originated by the contact between the azeotropic mixture fed in stage 3 and the IL coming down from the stage above. It is clear that the presence of the IL in the reboiler increases the boiling temperature of the ethanol and IL mixture. Since the IL is not present in the condenser because of its negligible vapor pressure, the temperature in this part is close to the ethyl acetate normal boiling temperature. Consequently, it can be inferred that the extractive separation of ethyl acetate from ethanol in the presence of [EMIM][MeSO 3] is successfully carried out in the extractive section. From the analysis of RCM and liquid/vapor behaviors of each components in the EDC, it is obvious that the extractive distillation using [EMIM][MeSO3] as solvent is sufficient to purify ethyl acetate with the desired purity. 3.3. Inflence of different design parameters on the purity of ethyl acetate and energy requirement The influence of the number of stages and the feed stage on the molar composition of ethyl acetate in the distillate and the reboiler duties were analyzed. Fig. 8a shows the relationship between the ethyl acetate purity in the distillate and the number of stages and feed stages, and results indicate that some cases with the special number stages and feed stages can be obtained ethyl acetate with 99.9 mol % purity in the overhead product. Fig. 8b shows the influence of the number of stages and feed stage on reboiler duties, results show that reboiler duties are immediately increased when the feed stage is close to the bottom. The influence of the IL/F and reflux ratio on the molar composition of ethyl acetate in the distillate and the reboiler duties were also analyzed, as shown in Figs. 8c and 8d. The IL/F is one of important parameters for the improvement of the ethyl acetate purity in the separation process. Fig. 8c shows that IL/F effects directly on the distillate purity without considerably affecting the energy consumption. Moreover, Fig. 8d shows that there is no significant change in energy consumption at a constant reflux ratio for different IL/F values within the interval 0.15 - 0.40. Similarly, at a constant IL/F the reflux ratio increases until a distillate composition equivalent to the one obtained in the previous variation is reached. From Fig. 8d the increase in energy consumption is 22.9 %, so the reboiler energy consumption in the extractive distillation for the separation of ethyl acetate and ethanol is obviously related to the reflux ratio. Thus, the reflux ratio must be set to the lowest possible value so that the IL/F ratio can be manipulated to reach the desire distillate purity without high energy consumption. 3.4. Process design verification through economic anaylsis 8
According to sensitivity analysis results shown in Fig. 8, it is obvious that the separation of ethyl acetate and ethanol mixture with [EMIM][MeSO3] is possible in some cases with different design parameters when the product purity specifications are fixed. Therefore, to determine the best option of these process design parameters it is necessary to carry out an economic evaluation based on the minimum TAC. The cost models for TAC calculation by Luyben [46] are shown in Appendix. The economics of the extractive distillation system are evaluated in terms of capital investment and energy costs. Fig. 9 shows the comparison of TACs of the EDC with different design parameters while the ethyl acetate purity is fixed at 99.9 mol %. TACs are no significant change along the different feed locations providing the desired purity. TACs of the cases with 50 stages are least in comparison to the other cases as seen in Fig. 9. Fig. 10 shows the optimal conditions and results of the ethyl acetate + ethanol azeotropic separation process when RR is fixed at 2.0. Solvent recovery process required a flash tank with the operating temperature of 117 ℃ and the heat duty of 511 kW at the condition of 30 kPa, and a stripper with 10 stages and the pressure of 100 kPa. Ethanol can be recovered from air + ethanol stream by condensation or absorption using water for the amount of ethanol is not high. Ethanol obtained from the flash tank and the stripper can be reused by sending to the reactive system for the dehydrogenation of ethanol. Furthermore, the design parameters of three cases and the corresponding costs of all units are summarized in Table 5. Solvent recovery system consisting of flash tank and stripper occupies 22 % ~ 24 % of the total TAC, so it can be sure that the process design is decided by EDC optimization in the separation process for ethyl acetate purification using [EMIM][MeSO3] as the solvent. 4. Conclusion This work focused on the extractive distillation process for the separation of ethyl acetate and ethanol mixture using ILs as solvents. Four ILs [EMIM][MeSO3], [EMIM][MeSO4], [BMIM][CF3SO3] and [EMIM][BF4] were employed as a set of feasible solvents and the UNIFAC-Lei thermodynamic model was applied to find the favorable IL solvent. The whole extractive distillation process including solvent recovery system of a flash tank and a stripper was conducted in Aspen Plus. The sensitivity analysis was performed to investigate the influence of various operating parameters on the separation target with the selection of a suitable IL. From the results, [EMIM][MeSO3] was selected as the proposed solvent in the separation process for the 9
ethyl acetate purification. Ethanol withdrawn from the solvent recovery system can be reused by sending to the reactive system for the dehydrogenation of ethanol. The separation process configuration and operating parameters were confirmed on the basis of the EDC optimization by TAC. The results show that the TAC of the optimal process is $745,380/y and the solvent recovery system occupies 24 % of the total cost of this process. Consequently, the optimization problem of the extractive distillation system by the selection of a favorable IL solvent is the main strategy that can allow economically efficient separation of close-boiling mixtures such as the ethyl acetate-ethanol system. Acknowledgement Financial support from National Natural Science Foundation of China (Project 21676152) is gratefully acknowledged. References [1] A.A. Kiss, Distillation technology–still young and full of breakthrough opportunities, J. Chem. Technol. Biotechnol. 89 (2014) 479–498. [2] N.V.D. Long, M. Lee, Optimal retrofit design of extractive distillation to energy efficient thermally coupled distillation scheme, AlChE J. 59 (2013) 1175–1182. [3] S. Liang, Y. Cao, X. Liu, X. Li, Y. Zhao, Y. Wang, Y. Wang, Insight into pressure-swing distillation from azeotropic phenomenon to dynamic control, Chem. Eng. Res. Des. 117 (2017) 318-335. [4] Y. Cao, J. Hu, H. Jia, G. Bu, Z. Zhu, Y. Wang, Comparison of pressure-swing distillation and extractive distillation with varied-diameter column in economics and dynamic control, J. Process Contr. 49 (2017) 9-25. [5] Z. Lei, C. Dai, J. Zhu, B. Chen, Extractive distillation with ionic liquids: A review, AlChE J. 60 (2014) 3312–3329. [6] H. Rodríguez, Ionic Liquids in the Context of Separation Processes, Springer Berlin Heidelberg, (2016) 1-9. [7] Z. Zhu, Y. Ri, M. Li, H. Jia, Y. Wang, Y. Wang, Extractive distillation for ethanol dehydration using imidazolium-based ionic liquids as solvents, Chem. Eng. Proc. 109 (2016) 190-198. [8] Z. Lei, R. Zhou, Z. Duan, Process improvement on separating C4 by extractive distillation, Chem. Eng. J. 85 (2002) 379-386. [9] S. Li, H. Liu, Entrainer Screening for Separation of Isoparaffins and Naphthenes by Extractive Distillation, Ind. Eng. Chem. Res. 48 (2009) 5427-5430. [10] Z.G. Zhang, D.H. Huang, M. Lv, P. Jia, D.Z. Sun, W.X. Li, Entrainer selection for separating tetrahydrofuran/water azeotropic mixture by extractive distillation, Sep. Purif. Technol. 122 (2014) 73-77. [11] N. Medinaherrera, I.E. Grossmann, M.S. Mannan, A. Jiménezgutiérrez, An Approach for Solvent Selection in Extractive Distillation Systems Including Safety Considerations, Ind. Eng. Chem. Res. 53 (2014) 12023-12031. [12] W.B. Ramos, M.F.D. Figueirêdo, K.D. Brito, S. Ciannella, L.G.S. Vasconcelos, R.P. Brito, Effect 10
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Fig. 1. Conceptual process flowsheet of the separation for ethyl acetate and ethanol mixture using IL as solvent.
13
Fig. 2. Schematic drawing to the EDC with IL as solvent.
14
Fig. 3. Flowsheet of the EDC optimization by sensitivity analysis and TAC calculation.
15
Fig. 4. Effect of different ILs on the ethyl acetate - ethanol VLE at xIL = 0.2. The dash line represents VLE curve without IL.
16
Fig. 5. Relative volatilities for ethyl acetate to ethanol using different ILs: , [EMIM][MeSO 3]; , [EMIM][MeSO4]; , [BMIM][CF3SO3]; , [EMIM][BF4].
17
Fig. 6. RCM with the isovolatility curve of ethyl acetate and ethanol using [EMIM][MeSO 3] as solvent.
18
Fig. 7. Composition and temperature profiles of an EDC for the separation of ethyl acetate and ethanol using [EMIM][MeSO3]: (a) liquid; (b) vapor; (c) temperature.
19
Fig. 8. Influence of several process design parameters on distillate composition and reboiler duties: (a) influence of the number of stages and feed stages on the ethyl acetate purity; (b) influence of the number of stages and feed stage on reboiler duties; (c) influence of IL solvent flow rate and reflux ratio on the ethyl acetate purity; (d) influence of IL solvent flow rate and reflux ratio on reboiler duties.
20
Fig. 9. Influence of number of stages and feed stages on TAC.
21
Fig. 10. The optimal conditions and results of the ethyl acetate + ethanol azeotropic separation process using [EMIM][MeSO3] at RR = 2.0.
22
Table 1. Group division of ILs ILs
Group division
[EMIM][MeSO3]
[MIM][MeSO3] + CH2 + CH3
[EMIM][MeSO4]
[MIM][MeSO4] + CH2 + CH3
[BMIM][CF3SO3]
[MIM][CF3SO3] + 3CH2 + CH3
[EMIM][BF4]
[MIM][BF4] + CH2 + CH3
23
Table 2. Parameters of Rk and Qk Group
Rk
Qk
CH3a
0.9011
0.8480
CH2a
0.6744
0.5400
1.0000
1.2000
a
OH
a
CH3COO
1.9031
1.7280
a
7.9863
5.5761
b
3.7481
3.0714
9.5357
5.0500
6.5669
4.0050
[MIM][MeSO3] [MIM][MeSO4]
b
[MIM][CF3SO3] b
[MIM][BF4] a
Reference [26]; b Reference [27]
24
Table 3. Group interaction parameters for the UNIFAC model m
n
CH2
[MIM][MeSO3]
661.05a
544.93a
CH2
[MIM][MeSO4]
575.45b
473.80b
CH2
[MIM][CF3SO3]
405.39b
284.37b
CH2
[MIM][BF4]
1108.51b
588.74 b
OH
[MIM][MeSO3]
1.14a
-389.86a
OH
[MIM][MeSO4]
13.83b
-570.26b
OH
[MIM][CF3SO3]
34.90b
-305.24b
OH
[MIM][BF4]
-13.77b
131.24b
a
-41.76a
CH3COO
[MIM][MeSO3]
767.31
CH3COO
[MIM][MeSO4]
n.a.
CH3COO
[MIM][CF3SO3]
n.a.
CH3COO a
[MIM][BF4]
32.09
n.a. n.a. a
112.17a
Reference [26]; b Reference [27]
25
Table 4. Several properties of ILs, which were used to operate the simulation with IL component in Aspen Plus Mw
Tb(K)
Tc(K)
Pc(bar)
Vc(cm3/mol)
zc
ω
η(cP)
ρ(g/cm3)
[EMIM][MeSO3]
206.3a
667.4a
1026.0a
48.1a
587.1a
0.3312a
0.3307a
86.0 (28℃)c
1.24c
[EMIM][MeSO4]
222.3b
602.7b
1053.6b
45.9b
602.7b
0.3115b
0.3401b
84.3 (25℃)d
1.28d
[BMIM][CF3SO3]
288.3a
707.7a
1023.5a
29.5a
750.7a
0.2600a
0.4046a
77.0 (25℃)e
1.31e
[EMIM][BF4]
198.0a
449.5a
596.2a
23.6a
540.8a
0.2573a
0.8087a
66.5 (20℃)f
1.25f
ILs
a
Reference [44]; b this work; c Reference [47]; d Reference [48]; e Reference [49]; f Reference [50].
26
Table 5. TAC of separation process for ethyl acetate purification Case 1
Case 2
Case 3
45/3/37
50/3/38
55/3/45
ID1 (m)
1.12
1.03
1.03
RR
2.5
2.0
2.0
1568
1343
1344
1789
1565
1566
1075.01
1032.08
1081.35
413.83
361.86
362.06
628.84
568.28
578.33
0.90
0.90
0.90
511
511
511
capital cost (×10 $)
169.13
169.13
169.13
3
115.98
115.98
115.98
149.80
149.80
149.80
10
10
10
0.652
0.652
0.652
136.46
136.46
136.46
27.3
27.3
27.3
805.94
745.38
755.43
EDC NEDC/NIL/NF
QC (kW) QR (kW) 3
capital cost (×10 $) 3
operating cost (×10 $) 3
TAC1 (×10 $) flash tank ID2 (m) QH (kW) 3
operating cost (×10 $) 3
TAC2 (×10 $) stripper Nstripper ID3 (m) 3
capital cost (×10 $) TAC3 (×103 $) 3
TAC (×10 $)
27
Appendix. Basis of economics Column vessel (diameter and length in meters) Length: L = 1.2 × 0.61 × (N – 2) Capital cost = 17640 (D)1.066 (L)0.802 Cooler (area in m2) Heat-transfer coefficient (KC) = 0.852kW K-1 m-2 Capital cost = 7296 (area)0.65 Reboiler (area in m2) Heat-transfer coefficient (KR) = 0.568 kW K-1 m-2 Capital cost = 7296 (area)0.65 Flash tank Liquid volume fraction: 0.5 m3 Residence time: 20min
Volume = D=
3
1 moleflow 20 2 60
2 volume , L = 2D
Capital cost = 17640 (D)1.066 (L)0.802
TAC =
capital cos t + operating cost payback period
Payback period = 5 years
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Nomenclature ETAC
Ethyl acetate
ETOH
Ethanol
[EMIM][MeSO3]
1-Ethyl-3-methylimidazolium methanesulfonate
[EMIM][MeSO4]
1-Ethyl-3-methylimidazolium methylsulfate
[BMIM][CF3SO3]
1-Butyl-3-methylimidazolium trifluoromethanesulfonate
[EMIM][BF4]
1-Ethyl-3-methylimidazolium tetrafluoroborate
[EMIM][OAc]
1-Ethyl-3-methylimidazolium acetate
[EMIM][Cl]
1-Ethyl-3-methylimidazolium chloride
Mw
Molecular weight
Tb
Normal boiling temperature (K)
Tc
Critical temperature (K)
Pc
Critical pressure (bar)
Vc
Critical molecular volume (cm3/mol)
Zc
Critical compressibility factor
w
Acentric factor
A.P.
Azeotropic point
αi/j
Relative volatility of ethyl acetate to ethanol
NEDC
Number of stages of extractive distillation column
NIL
Solvent feed stage of extractive distillation column
NF
Feed stage of extractive distillation column
Nstripper
Number of stages of stripper
RR
Reflux ratio of extractive distillation column
ID1
Diameter of extractive distillation column
ID2
Diameter of flash tank
ID3
Diameter of stripper
QC
Heat duty of cooler (kW)
QR
Heat duty of reboiler (kW)
QH
Heat duty of flash tank (kW)
29
Highlights VLE behaviors of ternary mixtures ethyl acetate-ethanol-IL are predicted by UNIFAC model. The separation process using [EMIM][MeSO3] as solvent is studied. The optimization of the extractive distillation column using IL becomes a main strategy.
30