Conceptual Design of an Energy-Efficient Process for Separating

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Conceptual design of energy efficient process for separating aromatic compounds from naphtha with high concentration of aromatic compounds using 4-methyl-N-butylpyridinium tetrafluoroborate ionic liquid Tae Hoon Oh, Se-Kyu Oh, Hosoo Kim, Kyungmoo Lee, and Jong Min Lee Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b00021 • Publication Date (Web): 02 Jun 2017 Downloaded from http://pubs.acs.org on June 13, 2017

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Industrial & Engineering Chemistry Research

Conceptual design of energy efficient process for separating aromatic compounds from naphtha with high concentration of aromatic compounds using 4-methyl-N-butylpyridinium tetrafluoroborate ionic liquid Tae Hoon Oh,† Se-Kyu Oh,†,‡ Hosoo Kim,¶ Kyungmoo Lee,¶ and Jong Min Lee∗,† †School of Chemical and Biological Engineering, Institute of Chemical Processes, Seoul National University, 1 Gwanak-ro, Gwanak-gu, Seoul, 08826, Republic of Korea ‡Co-first author ¶LG CHEM R&D Campus, Yuseong-gu, Daejeon, 34122, Republic of Korea E-mail: [email protected] Phone: +82-2-880-1878 Abstract In order to obtain aromatic compounds from a crude mixture such as reformate or pyrolysis gasoline, three different processes are simulated with the realistic composition of reformate and product specification. Simulations were performed by Aspen Plus supported with COSMO-RS method to predict the physical and thermodynamic properties of ionic liquid. Furthermore, utility analysis and economic evaluation are presented. Conventionally, aromatic compounds are extracted from a crude mixture either by extraction or by extractive distillation using a solvent such as sulfolane and separated by

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a series of distillation columns. In this study, the sulfolane-based commercial process is firstly introduced and two novel processes which use 4-methyl-N-butylpyridinium tetrafluoroborate ionic liquid as solvent are proposed. The Second process shows that energy consumption has successfully decreased, but the high price of ionic liquid offset the cost advantage. The Third process is therefore proposed to reduce the amount of ionic liquid using two extractions. A Similar level of energy saving is achieved with reduced costs.

Introduction Separating aromatic compounds, especially benzene, toluene, ethylbenzene and xylene (BTEX) from naphtha is an interesting issue in the petrochemical industry. However, the separation is not simple because BTEX have boiling points close to those of the C4 to C10 aliphatic compounds. This makes it difficult to use distillation alone in order to separate BTEX from aliphatic compounds. For this reason, instead of using distillation, extraction has been conventionally used. Usually for the range of 20 to 65 wt.% of aromatic compounds in naphtha stream, extraction column is used. For the range of 65 to 90 wt.%, extractive distillation column is suitable and for the range higher than 90 wt.%, azeotropic distillation column is recommended. 1,2 In industrial scale, UOP proposed the most conventional extraction process which uses sulfolane as a solvent to extract BTEX in 1977. 3 Although this process can achieve the desired separation, there are several drawbacks. The major issue is the high energy consumption. Since regenerating the solvent is inevitable for an extraction method, high boiling point of the sulfolane (560K) requires huge energy consumption. Furthermore, the process is not suitable for low concentration of aromatic compounds in the naphtha feed stream, and additional separation is often needed to recover the solvent from raffinate stream. 4 Ionic liquids (ILs) have been studied as alternative solvent for separating aromatic compounds from aliphatic compounds. 5–7 ILs are salt in liquid state which typically has melting 2


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point below 373K. Compared with sulfolane, ILs have some advantages as a solvent. The most notable property of ILs is a very low volatility. This is critical in reducing the energy consumption when regenerating the solvent using vacuum distillation column. Therefore, the process using ILs is expected to require less energy than any existing processes. 7–9 Another attractive property is that ILs are salt, which is made by combining cations and anions. By choosing different cations or anions, there are hundreds of possible ILs with distinct performance. Therefore, many options can be chosen depending on the type of process units, operating condition, naphtha feed stream, and economic values. 10 Various ILs have been studied for separating aromatic compounds. Among them are imidazolium, pyridinium and pyrrolidinium combined with several anions such as [TF2 N]− , [CH3 SO4 ]− , [BF4 ]− , etc. With the aromatic property of the cations, these have high aromatic distribution coefficient and selectivity and thus are suitable for the separation. 10 A comprehensive review of research work on ILs for the past decade can be found in meindersma. 11 It is reported that mixture of ILs can increase separation efficiency. 12,13 Studies on low-viscous ILs as a solvent have been conducted. 14 Despite a large number of studies on ILs, only a small number of ILs show higher distribution coefficient and selectivity than those of sulfolane: 1-butyl-3-methylimidazolium dicyanamide[bmim][DCA], 1-butyl-3-methylimidazolium thiocyanate [bmim][SCN] 1-butyl-3-methylimidazolium tricyanomethanide [bmim][TCM] and 4methyl-N-butylpyridinium tetrafluoroborate [mebupy][BF4 ]. 11,14,15 A great part of the ILs research is focusing on the experiment or the modeling to find out basic properties such as distribution coefficient, selectivity and equilibrium conditions which are all based on thermodynamics knowledge. 16–28 There is research on the conceptual process design level and pilot plant scaled experiments. 29–34 Most of the studies on process level are dealing with a naphtha feed stream which contains about 10 to 20 wt.% of aromatic compounds. Since the process using sulfolane is not suitable for this range, process using ILs shows good results. However, it is also expected that ILs can show better results than sulfolane for the naphtha feed containing high concentration of aromatic compounds. There-

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fore, the aim of this paper is to show that the process using ILs can achieve the product specification with much less energy consumption and more economic benefit even with the realistic feed stream containing high concentration of aromatic compounds. Most of existing works on processes using ILs 4,7,29 have been done with commercial process simulators (mostly Aspen Plus and/or Aspen HYSYS). However, there are some issues with simulating process using ILs. 35 Since ILs are not commonly used compounds, no data for ILs are embedded in commercial process simulators. Therefore, all physical parameters should be embedded by experiment or estimated values. Another problem occurs from using activity models. To calculate the activity models such as NRTL or UNIQAC, binary interaction parameters between ILs and organic compounds are necessary. Mostly, there are not enough data for such parameters. Therefore, a feasible method with proper assumptions and algorithms is needed to estimate the binary interaction parameters. The Conductor like Screening Model for Real Solvent (COSMO-RS) method have drawn attentions to deal with these problems. 36–40 Based on the quantum chemical calculations and statistical thermodynamics, the COSMO-RS method predicts physical properties of mixed components and other thermodynamics parameters only by using the data of molecular structure. Combing Aspen Plus and COSMO-RS method has already been used to simulate and analyze separation process using ILs, including sensitivity analysis, operating conditions, energy consumption, operating cost and capital cost. 41–43 Furthermore, it is already suggested that using the COSMO-RS model gives reliable results for the feed of multicomponent realistic mixture. 35 Compared to the other ILs, 4-methyl-N-butylpyridinium tetrafluoroborate ([4-methyl][BF4 ]) has a fair amount of studies and data including activity coefficients at infinite dilution gamma, 44–46 densities, and viscosities. 47 There are also available data for liquid-liquid equilibrium of various ternary mixtures and for binary and ternary systems predicted by NRTL model. 48–52 It is also claimed that [4-methyl][BF4 ] can serve as a suitable solvent for industrial extraction process for separating aromatic and aliphatic hydrocarbons due to its higher selectivity and partition coefficients than those of sulfolane. 27,48 In addition, recent

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study shows that a COSMO-RS estimation method can provide reliable property estimates of [4-methyl][BF4 ]. 35 Hence, this study selects [4-methyl][BF4 ] as an IL solvent for extracting aromatic compounds expecting that [4-methyl][BF4 ] can exhibit higher separation performance than sulfolane and its further analysis can be validated in a reliable, quantitative manner. This work presents design and simultion study of three different processes using Aspen Plus with COSMO-RS estimation method. Realistic feed data for a stream with a high concentration of aromatic compounds were used. The total feed stream flow is about 47 tons/h containing 16 components with 68 wt.% of aromatic compounds which are mostly benzene and toluene. With the same feed stream of naphtha, each process separates them according to the desired specifications. Energy analysis and economic evaluation were also done. The first process is a base case, which describes a real existing plant. Simulation results of the proposed processes are compared with this base case. The base case contains an extractive distillation column, because the feed composition of aromatic compounds is over 65 wt.%. For the second and third processes, IL is used for liquid-liquid extraction. The second process only differs from the first one with extraction units and solvent. The third process is a new process proposed to reduce the amount of IL. Reducing a solvent flow has advantages on equipment sizing and raw material cost.

Computational details Component property Components except for the C10+NA (Exo-tetrahydrodicyclopentadiene) and IL were defined as conventional components in Aspen Plus v8.8. Because C10+NA in the feed is not included in Aspen Plus, its binary parameters were obtained using NRTL model with the values in the literature. 53,54 The IL [4-methyl][BF4 ] could be defined in Aspen Plus as conventional components or as pseudo-components. However, adding the IL as conventional components 5



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requires huge amount of data, 35 thus pseudo-component method was used to handle the IL. Basic properties such as molecular weight, normal boiling point and density of IL were added. For activity coefficient model, NRTL was selected. To calculate the binary coefficients of NRTL, a commercial program called COSMOtherm was used to obtain the LLE data between IL and other components. This work used COSMOtherm (ver. C3.0 release 15.01) to calculate the ternary liquid-liquid equilibria for mixtures. 40,55–57 COSMOtherm can specify IL by selecting cation and anion respectively. In particular, COSMOtherm includes [mebupy]+ and BF− 4 , cation and anion of the IL of our interest, respectively. Ternary LLE data for a mixture at a given temperature was first obtained using COSMOtherm. This data was then regressed in Aspen Plus to obtain the binary parameters of NRTL. For the verification purpose, we have also generated LLE data of other mixtures, Cyclohexane-Benzene[4-methyl][BF4 ], Cyclohexane-Toluene-[4-methyl][BF4 ] , n-Heptane-Benzene-[4-methyl][BF4 ] and n-Heptane-Toluene-[4-methyl][BF4 ], using COSMOtherm whose experimental data are available in the literature. Furthermore, the experimental data of pilot plant scale 4 and the simulation data under same condition were compared.

Conceptual process designs Feed Conditions and Product Specifications All three cases have the same naphtha feed stream which has 47 tons/h of mass flow. This naphtha feed stream goes into the separation process as 80.2◦ C and 1.2 bar. The specific conditions and composition of the naphtha feed stream are given in Table 1. About 68 wt.% of the feed stream are aromatic compounds, mostly composed of benzene and toluene. Components heavier than toluene rarely exist because an upstream process separates them ahead of extraction stage. Therefore, the main goal of the process is to separate benzene and toluene at reasonable cost. A product specification is determined as same as required for the real BTEX separating plant. Table 2 shows the product specifications. The purity required for toluene is slightly lower than that for benzene because ethylbenzene is included 6


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Table 1: Naphtha Feed Composition Component Cyclopentane Hexane Cyclohexane Benzene Methylcyclohexane Toluene Octane Ethylcyclohexane Ethylbenzene p-Xylene m-Xylene o-Xylene Nonane 1,2,4-Trimethylbenzene Exo-Tetrahydrodicyclopentadiene 1,2,3,5-Tetramethylbenzene

Wt.% 3.0 13.5 12.1 45.8 2.7 22.8 < 0.1 < 0.1 < 0.1 < 0.1 < 0.1 < 0.1 < 0.1 < 0.1 < 0.1 < 0.1

Table 2: Product Specification Product Stream Specification C6 Raffinate Benzene < 0.3 wt.% Benzene purity > 99.9 wt.% Toluene < 500ppm Benzene Non-aromatics < 700ppm Recovery of benzene > 99.9wt.% Toluene purity > 99.4 wt.% Toluene Recovery of toluene > 99.8 wt.% in the toluene product stream. Configuration 1 (Fig. 1) This configuration imitates a process that actually exists to separate aromatics from the above feed stream. However, this process is not the same as the UOP process that is currently used commercially to separate aromatics. Most of the extraction process goes through two stages. Separating the desired substance from the feed stream using a solvent, and regenerating the desired substance from the solvent. In case one, extraction distillation columns were used because of the high ratio of aromatic compounds in the feed stream, and 7



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sulfolane was used as a solvent. The naphtha feed (FEED) and sulfolane (SOL) are inserted into the extractive distillation column (EXT-COL) which correspond to the separation stage. In this operation, raffinate stream (RAF) is aliphatic rich. The extract stream (EXT-BOT), mainly containing sulfolane, benzene and toluene, goes into a next vacuum column (COL1) in order to recover the sulfolane as an extract stream (SOL-RE). This corresponds to the regeneration stage. Recovered sulfolane is mixed with a make-up stream (MAKE-UP) and recycled back to the extractive distillation column. Top stream (B-T-C) which contains benzene, toluene and a few aliphatic compounds moves forward to the distillation column which separates benzene (COL2) and toluene (COL3) from remaining aliphatic compounds. Note that the final product is a single component rather than a mixture of benzene and toluene. An additional distillation column (COL3) was used to adjust the purity of toluene, which is an additional disadvantage of using sulfolane. The recycle loop was modeled by using the mass balance tool of Aspen Plus. This makes the solvent stream have a constant mass flow rate by controlling the flow rate of make-up stream. However, since the objective is to compare the three processes, the operating conditions are designed such that the flow rate of the make up stream is close to zero. Configuration 2 (Fig. 2) In this case, IL [4-methyl][BF4 ] solvent is used. A liquid-liquid extraction column is used instead of an extractive distillation column to reduce the energy consumption. Since IL shows better performance than sulfolane, the extraction column alone is enough to achieve the desired product specification. The same naphtha feed stream (FEED) is going into the extraction column (EXT) with IL (SOL-P-H) and recycled stream (STR-R-H). The extract stream (EXT-BOT) of the extractor moves into the stripper (STR) so that remaining aliphatic compounds are recycled. The stripper (STR) is designed with no condenser since its purpose is just boiling the feed stream. Most of the light materials are boiled to follow the recycle loop. The inner recycle loop together with extractor and stripper corresponds to

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the separating stage. The bottom stream of the stripper (STR-BOT) moves further to the vacuum distillation column (COL1) to recover IL. The recovered IL (IL-RE) is recycled back to the mixer while benzene and toluene move forward to the next distillation column (COL2). There is another outer recycle loop which originated from the raffinate of extraction column (EXT-RAF). Since the amount of IL in the raffinate stream is small, outer recycle loop might not be needed in actual process. However, to compare the three cases, additional distillation column (COL3) and recycle loops were used to achieve zero mass flow rate of make-up stream. These two outer recycle loops correspond to the regenerating stage. Contrast to case one, all the aliphatic compounds are fully separated by extractor and stripper. Therefore, the product specification can be achieved without using an additional column, corresponding to (COL3) in case one, for separating toluene. Configuration 3 (Fig. 3) This case was proposed to reduce the amount of IL. The key to this process is to change the order of the extraction and separation of benzene and toluene. Therefore, the feed stream is first separated into a benzene-rich(COL-B) stream and a toluene-rich(COL-T) stream using a distillation column(COL1). The benzene rich stream goes to extraction column (EXT1) with the solvent stream (SOL1) and recycle stream (STR1-RE). Similar to case two, the bottom stream of extraction column (EXT1) moves to the stripper (STR1). Remaining aliphatic compounds are recycled back to the extraction column and separated to raffinate stream (EXT1-RAF). The vacuum distillation column (COL2) recovers the remaining IL in (EXT1-RAF) and recycle them back to the mixer. Again, since the amount of IL in (EXT1-RAF) stream is small so that this column can be removed in actual process. The bottom stream of stripper (STR1-B) mainly contains benzene and IL. Using vacuum column (COL3), benzene rich product stream (B) can be obtained. IL stream (ILRE) is then split into two streams, one of which goes directly back to the mixer (ILRE1) for recycle and the other of which goes to the extraction column (EXT2) with a toluene rich stream (COL-T)

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and recycled stream (STR2-R-H). Since the amount of benzene is larger than that of toluene in this case, only half of the IL is needed to extract toluene. This way, the recycling effects could be maximized. The procedure for extracting toluene is the same as extracting benzene. The raffinate stream mostly contains aliphatic compounds. The extract stream goes to the stripper (STR2) to recycle the remaining aliphatic compounds. The bottom stream of the stripper (STR2) again moves to the vacuum distillation column (COL4) to recover the IL. The toluene rich product stream (T) is obtained. The remaining IL (ILRE2) is recycled back to the mixer. Model and operating conditions of units All multi-stage columns except for extractor were modeled by RadFrac model in Aspen Plus. The RadFrac model is a rigorous model for simulating multistage operations. Extractor model was used for the extraction column. Specific information on the model used in each case is given in Table 3. In Case one, the design and operating values of the extractive distillation column (EXT-COL) were set up similar to the commercial process and benzene it was confirmed that benzene and toluene could be separated from the naphtha feed stream. For the remaining columns, the short-cut model was used to determine the number of stages. In general, as the number of stages in the distillation column increases, the required reflux ratio decreases. Therefore, large number of stages can reduce operating costs, but also increases the initial installation cost. A proper balance must be found between the two variables. In this study, the number of stage was selected to minimize the objective function, Jobs = [ reflux ratio × number of stages ]. Figure 5 shows the value of the objective function according to the number of stages in each column. After setting the model with RadFrac, input stage was determined to minimize the total duty of the column within the range showing the required separation performance. The operating pressure of the distillation columns is set at 1 bar and the total pressure drop of the column is set at 0.2 bar. For the vaccum distillation columns, the pressure is set to be 0.08 bar with a total pressure drop of

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Table 3: Information of design variables used in each case Number of stages Input stage reflux ratio (mole) Case one EXT-COL COL1 COL2 COL3 Case two EXT STR COL1 COL2 COL3 Case three EXT1 EXT2 STR1 STR2 COL1 COL2 COL3 COL4

Utilityb

29 13 46 43

3/19a 7 23 4

1.231 0.195 1.196 1.271

Air/MP REF/HP Air/MP REF/HP

30 19 13 52 26

9c 1 7 26 13

None 14.144 0.270 1.209 0.350

None/None None/LP REF/MP Air/MP REF/MP

30 20c None None/None 30 9c None None/None 13 1 12.740 None/LP 6 1 76.560 None/MP 41 23 0.965 Air/MP 26 6 0.490 Air/MP 12 8 0.230 REF/MP 13 9 0.520 REF/MP a Each stage is the input stage of solvent and naphtha. b REF stands for refrigerant, LP for MP and HP for low pressure stream, medium pressure stream, and high pressure stream, respectively. c Shows the input stage of the recycle stream. The solvent stream is fed to the first stage and the naphtha stream to the last stage. 0.1 bar. The reason for choosing 0.08 bar is related to the degradation of sulfolane. In the case of sulfolane, solvent degradation occurs at temperatures above 220◦ C. 58 Therefore, the temperature of the reboiler in the vacuum distillation column needs to be less than 220◦ C. The pressure of the column and the temperature of the reboiler have a positive correlation. When the column pressure is 1 bar, the temperature of the reboiler is 287◦ C, whereas for 0.08 bar the temperature of the reboiler is 217◦ C. Therefore, 0.08 bar was chosen to eliminate the risk of solvent degradation. It is worth to note that the UOP process, which was mentioned ahead, was able to operated at mild pressure condition by using stripping steam. However the purpose of case one is to compare with the processes using IL. Therefore, the pressure and pressure drop of all vacuum distillation columns are set as same and stripping steam 11



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was not used in all cases. For the extraction column, the temperature of the stream and the input stage have a greater effect than the number of stages. Therefore, in case two, the number of stages is set to 30, and the influence of the input streams on the temperature is investigated. When the temperature of the solvent stream is high, the separation efficiency is further improved and the amount of solvent required can be reduced. In addition, utility duties that occur during cooling of recycled streams can also be reduced. On the other hand, if too high a temperature is used, the stripper placed behind the extraction column can not be operated. Since the composition of the entire process including recycling may vary slightly, The temperature of the solvent stream was set at 100◦ C with a margin. In case of Stripper, we modeled to minimize recycling within the range of product specification. Only 3.2% of the stream entering the stripper is recycled and the minimum number of stages in which aliphatic compounds do not enter the extract is used. Likewise, to satisfy the specification, the optimum amount of solvent was determined by increasing the amount of solvent until benzene and toluene exiting into the raffinate of the extraction column became appropriate values. The input stage of the stream recycled from the stripper and fed into the extraction column has more influence on the convergence than the separation efficiency. Therefore, the input stage was determined considering the robustness of the process. Three distillation columns were modeled similar to Case one. The numerical values and results for modeling in Case two can be found in Fig. 6 and Table 3. Case three is also designed similar to the previous cases, and the optimal number of distillation columns is shown in Fig. 7

Results and discussion Validation of COSMO-RS Method As mentioned previously, the properties of ionic liquid were estimated by COSMO-RS method. There are several experimental LLE data between aromatic compounds, aliphatic 12


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Table 4: Comparing the mass flow rate of extraction column of pilot plant and that of simulation results Pilot planta Raffinate Extract 0.12 0.75 0.58 0.29 0 9.25 9.22 0.03 19.88 0 0.03 19.85 20 10 9.83 20.17 Pilot plant data are taken from. 4

❵❵❵ ❵❵❵ kg/hr ❵❵❵ ❵❵❵ Composition ❵❵ ❵

Toluene Heptane Ionic liquid Total

a

Solvent Feed

Simulation Raffinate Extract 0.73 0.14 9.18 0.07 0.07 19.81 9.98 20.02

compounds and [4-methyl][BF4 ]. 59,60 For the validity of using COSMO-RS method, comparisons between the experimental data and predicted data were represented in four cases. Figure 4 shows the results for Cyclohexane-Benzene-[4-methyl][BF4 ], Cyclohexane-Toluene[4-methyl][BF4 ] , n-Heptane-Benzene-[4-methyl][BF4 ] and n-Heptane-Toluene-[4-methyl][BF4 ]. The root mean square deviations (RMSD) of each case were 0.0108, 0.0112, 0.0112, and 0.0128, respectively. The difference between the experimental data and the estimated data was considered to be acceptable. Therefore, NRTL parameters were estimated by Aspen Plus Data Regression System using the LLE data from COSMOtherm. Since the most important part of the process in this paper is the extraction column, the validity of extraction column using IL needs to be considered. For validation, extraction column was modeled as same as the one used in pilot plant experiments. 4 The comparison between pilot plant data and simulation data are shown in Table 4. According to Table 4, the amount of IL and toluene were higher in the extract of the pilot plant than the simulation results, and that of heptane was lower. This indicates that actual plant results will show better performance than what the simulation predicts. Configuration 1 Using the realistic multi-components naphtha feed stream, it was successfully modeled to produce desired product streams. Extractive distillation column, vacuum distillation column, two distillation columns, one pump and two coolers were modeled. The specific stream

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conditions such as temperature, pressure, stream enthalpy and mass flow rate of the major components are listed in Table S1 (Supplementary Material). This process has one recycle loop which recovers the solvent by vacuum distillation. The result shows that the amount of make-up stream is negligible. The satisfaction of product specification can be easily checked from Table S1. The recovery of benzene is over 99.9 wt.% and the recovery of toluene is over 99.8 wt.%. The purity of product streams is over 99.9 wt.% for benzene rich product stream and over 99.4 wt.% for toluene. The mass fraction of both benzene and toluene in the raffinate stream is less than 0.1 wt.%. The solvent to feed ratio (S/F) is about 4.5 (211.96 tons/h of sulfolane). Contrast to conventional extraction process, no additional column was needed to separate the solvent from the raffiante stream. This shows that the extractive distillation column could lower the process complexity. However, the energy consumption of the extractive distillation column is excessive. This is shown in Fig. 8, where the utility duty for all cases are presented. In case one, both total heating and cooling duties are about 31Gcal/h which is highest of all cases. In more detail, most of the heating duties are required for the extractive distillation column rather than the vacuum distillation column. For the extractive distillation column, about 43% of total heating duty is consumed while 33% is required for the vacuum distillation column. On the other hand, cooling the recycled solvent from 217◦ C to 25◦ C account for 56% of the total cooling duty. Without extractive distillation columns, reduction in total utility duty is expected. Economic evaluation was done using Aspen Plus economic analyzer. The purpose of economic evaluation is to see if the process using IL can achieve an economic benefit, compared with the process using sulfolane, even with the naphtha feed stream containing high concentration of aromatic compounds. The summary of results is in Table S4. For the solvent cost, sulfolane price is specified as 3.7 US dollars per kilogram which is the typical price in industry. 61 The naphtha feed cost and product price are not calculated since all three pro-

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cesses have same streams so that same value will be added or subtracted for each process. The total annual cost was calculated by adding 30% of each total capital cost and solvent cost to total operating cost. The results of the economic evaluation show that utility duty occupies a large portion of the cost. About 41% of the total annual cost is utility cost. In addition, 58% of the total annual cost is the operating cost. On the other hand, the solvent cost is less than 2% of total annual cost. This analysis implies that less energy consumption is needed to reduce the total annual cost. Configuration 2 Because of the high energy consumption of configuration 1, reducing the utility cost by using the IL [4-methyl][BF4 ] as solvent was done in this process. With the same naphtha feed data and product specification, only a few differences exist. Although the aromatic compound rate is high, the extractive distillation column is changed to liquid-liquid extraction to reduce the heating duty. Additional heating unit was added by using the stripper model in Aspen Plus. With no condenser in the stripper, its role is to boil the stream and recycle the light compounds. In this way, sharp separation could be achieved with relatively less energy consumption. Furthermore, this sharp separation also helps to get the product specification with one less distillation column since heavy compounds was already separated. However, some of IL is remaining in the raffinate stream of the extraction column. Therefore, additional distillation column was used to recover the remaining IL. The solvent recycle loop was designed similarly as case one. The process successfully produced the desired streams which satisfy the product specification. In Table S2, the specific data are listed. The product stream satisfies the specification. The recovery of benzene and toluene are over 99.9 wt.% and 99.8 wt.%. The purity of each product stream is over 99.9 wt.% for benzene rich product stream and over 99.4 wt.% for toluene. The mass fraction of both benzene and toluene in the raffinate stream is less than 0.1 wt.%. The solvent to feed ratio (S/F) is 5.6 (263.00 tons/h of IL) which is higher than

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sulfolane. This is mainly attributed to the molecular weight of [4-methyl][BF4 ] which is about two times of sulfolane. In Figure 8, the total heating and cooling duties in case two are 23Gcal/h and 24Gcal/h, respectively. The overall utility duty is about 76% of case one. Since there is no extractive distillation column, the largest heating duty consumption occurs at the vacuum column, which is about 44% of total heating duty. Similar to case one, most of the cooling duties are required for cooling the solvent, which is 40% of total cooling duty. The results show that using the IL as solvents successfully reduces the energy consumption. This is also directly shown in the economic evaluation. In Table S5, the utility cost is about 68% of case one which is lower than the ratio of utility duty. This is because separation using sulfolane required to use middle pressure or high pressure steam which is more expensive than low pressure stream which is usually used in this case. The capital cost is a little bit higher than case one, it is not enough to offset the benefits of reducing the energy consumption. However, the solvent cost is much higher than case one and this does offset the benefits of reducing the energy consumption. This is attributed to the high price of IL [4-methyl][BF4 ] which is taken as 30 US dollar per kilogram. 5,62 Figure 9 shows that the total solvent cost and Figure 10 shows the share of each item in total annual cost. The total solvent cost for case two is about ten times higher than that of case one. This high cost of solvent accounts for about 17% of total annual costs and it sets off the economic benefit of energy consumption. For this reason, total annual cost is 11% higher than case one. Configuration 3 In case two, reducing the amount of solvent seems to be critical since the price of [4-methyl][BF4 ] is too high. Therefore, this process was proposed to use much less amount of solvent. As mentioned ahead, this process separates the feed naphtha to benzene rich stream and toluene rich stream before the extraction. After the separation, the extraction of benzene takes place at the upper extraction column (EXT1) using IL. The cycle from extractor is similar to case

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two, but the feed mass flow is less than that of case two. Hence only 60% of IL is required. The solvent to feed ratio (S/F) for a benzene extraction column is 4.5 (158.02tons/h of IL). For toluene rich stream, the portion of the recycled IL is used to separate the toluene. About half of recycled IL (78.21tons/h of IL) is used to separate toluene from naphtha stream. The separating process is the same as that of separating benzene. The solvent to feed ratio is 6.7. Since the S/F of the toluene stream does not affect the total amount of IL, a relatively high value was set to increase the separation efficiency. The specific data are listed in Table S3. The results also satisfied the product specification. Both recovery of benzene and toluene are over 99.9 wt.%. The purity of each product stream is over 99.9 wt.% for benzene rich product stream and over 99.4 wt.% for toluene. The mass fraction of both benzene and toluene in the raffinate stream is less than 0.1 wt.%. Fig.8 shows that the utility duty is bit higher than case two. The total heating duty is 24Gcal/h and cooling duty is 25Gcal/h. Unlike the previous cases, the most energy consuming units for heating duty are newly added distillation column (COL1) 26%, followed by the stripper (STR1) 15%. For cooling duty, 46% of total cooling duty are the highest energy consumption, which occurs at cooling the solvent coming from the first extraction column. Only 27% is needed for cooling the solvent coming from the second extraction column. Although the utility duty is similar to case two Table S6 shows that the cost is lower than case two. This is because, unlike case 2, the amount of solvent used is small. Therefore, the rate of cooling using air can be increased. Figure 10 shows the decrease in total annual cost due to the decrease in the total operating cost and the amount of solvent, whereas the capital cost has increased compared to Case 2. Particularly, it can be seen that the total cost is slightly lower than that of the sulfolane process because the portion of the solvent cost is reduced. As mentioned in Configuration 2, the price of IL was set as 30 US dollars per kilogram. This was based on the prospect that the price of IL can reach the level of 10–30 Euros per kilogram with production on a large industrial scale. 5,62 Therefore, given this projection of

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the IL price, it is possible to achieve economic benefits from the IL process. Figure 11 shows the total annual cost calculated according to the IL price. In the figure, the horizontal line represents the total annual cost of the sulfolane process. In case two, it is difficult to obtain more economic benefit than than the process using sulfolane, except for the case where IL price is 10 US dollars per kilogram. On the other hand, in Case 3, the cost is lower than the process using sulfolane, regardless of the price of IL and the reduction of total annual cost up to 9% could be achieved. In addition, processes using ILs have advantages in running processes over a long period of time because they have lower operating costs than those of the sulfolane process.

Conclusion It is expected that the BTEX extraction process using IL can be more efficient in terms of energy consumption than the conventional process using sulfolane. Studies on the separation of naphtha with a low proportion of aromatic compounds, which are unsuitable for separation using sulfolane, have been carried out, but there have been few studies on the case where the proportion of aromatic compounds is high. Therefore, this paper shows that even if the ratio of aromatic compounds in naphtha crude oil is high, IL can still be competitive compared to sulfolane. For this purpose, the conceptual design of the extraction process using sulfolane, IL and a small amount of IL were done through the commercial program Aspen Plus, respectively. First, [4-methyl][BF4 ] was selected among the ILs expected to be more competitive than sulfolane, and physical properties were estimated through the COSMO-RS method. Then, each process was modeled. The design variable and operating condition of process were determined through sensitivity analysis. Finally, each process was evaluated through utility and economic analysis. As a result of economical analysis, it was shown that the cost of utility for extraction process using IL was much lower than that of sulfolane. The total annular cost of the process using IL as a solvent instead of sulfolane

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was increased due to the high price of IL. However, in the newly proposed process, the total annular cost was lower than that of sulfolane because the use of IL could be reduced even though the installation cost was increased. From the above analysis, It has roughly shown that even when the ratio of aromatic compounds in the naphtha feed stream is high, IL still have an advantage over conventional sulfolane-based processes. Especially, as shown in previous studies, 10,11,29,35 energy consumption is reduced and the utility cost is decreased by the mild operation condition. However, this study is at a conceptual stage, and more sophisticated simulation and analysis are required to take into account various factors for practical applications. In particular, there are several factors that can complement the claim of this paper. One of those factors is the use of more effective and less expensive ILs. As mentioned earlier, studies on IL are underway and various ILs other than [4-methyl][BF4 ] can be used. Since the cost of operating the process can be quite dependent on the price of the solvent, in selecting the new IL, not only the efficiency of separating the BTEX but also the production cost per unit weight should be considered. It is also necessary to analyze the duration of continuous use of IL. In addition, a more rigorous comparison of energy can be made by adding heat integration that was not covered in this paper. Finally, the analysis of the various proportions of aromatic compounds in the naphtha will be able to support the claim more solidly.

Author Information Corresponding author *Phone: +82-2-880-1878. E-mail: [email protected] Notes ‡ Co-first author.

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Acknowledgement This work was supported by the Brain Korea 21 Plus Project in 2016 (5261-20150020).

Supporting Information Available The following files are available free of charge. • Table S1–S3: Mass flow rates and conditions of streams in three cases. • Table S4–S6: The results of economic analysis for three cases. This information is available free of charge via the Internet at http://pubs.acs.org/.

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Figure 1: The first configuration: Separation of aromatic and aliphatic compounds by extractive distillation followed by a vacuum distillation column and two more a series of distillation columns.

Figure 2: The second configuration: Separation of aromatic and aliphatic compounds by extraction using ionic liquid and stripper. The vacuum distillation column and two more distillations column are followed.

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Figure 3: The third configuration: Separation of aromatic and aliphatic compounds by two extractions each followed by a vacuum distillation column and one more distillation column.

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80 10

60

40

70 20

50

50

60 30

40

60

50 40

30

70

40

60

50

20

80

30

60

10

90

20 70

40

Toluene 0

0

90

Cyclohexane

20

90 0

Toluene 0

IL

(c) IL + Heptane + Toluene 313.2K

10 10

20

30

40

50

60

70

80

90

0

IL

(d) IL + Heptane + Toluene 348.2K

Figure 4: Comparing various LLE data between experiment (large circle) and COSMO-RS method prediction (filled square). The data values in (a) and (b) are weight percentage based and the data values in (c) and (d) are based on mole percentage. Experimental data are taken from following references. 59,60

31



Column 1

2.2

Objective function

2 1.8 1.6 1.4 1.2 1

8

9

10

11

12

13

14

15

16

17

18

47

48

49

50

47

48

49

50

Number of stage

Column 2

62

Objective function

61 60 59 58 57 56 55 54 40

41

42

43

44

45

46

Number of stage

Column 3

59

58

Objective function

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

57

56

55

54 40

41

42

43

44

45

46

Number of stage

Figure 5: The value of the objective function according to the number of stages in case one. The objective function is Jobs = [ reflux ratio × number of stages ] 32


Page 32 of 59

Page 33 of 59

Column 1

4.2

Objective function

4 3.8 3.6 3.4 3.2 3

9

10

11

12

13

14

15

Number of stage

Column 2

72

Objective function

70

68

66

64

62 45

46

47

48

49

50

51

52

53

54

55

27

28

29

30

Number of stage

Column 3

1.09 1.085

Objective function

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


1.08 1.075 1.07 1.065 1.06 1.055 1.05 20

21

22

23

24

25

26

Number of stage

Figure 6: The value of the objective function according to the number of stages in case two. The objective function is Jobs = [ reflux ratio × number of stages ] 33



Column 1

46

Column 2

2.55

2.5

44

Objective function

Objective function

45

43 42 41 40

2.45

2.4

2.35

39 38 35

36

37

38

39

40

41

42

43

44

2.3 20

45

21

22

Number of stage

24

25

26

27

28

29

30

Column 4

9.5 9

Objective function

2.4 2.3 2.2 2.1 2 1.9 10

23

Number of stage

Column 3

2.5

Objective function

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

Page 34 of 59

8.5 8 7.5 7

11

12

13

14

15

16

17

18

19

6.5 12

20

13

Number of stage

14

15

16

17

Number of stage

Figure 7: The value of the objective function according to the number of stages in case three. The objective function is Jobs = [ reflux ratio × number of stages ]

34


Page 35 of 59

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 8: Cooling duty (up) and heating duty (down) of three cases.

35



106 8

Solvent Cost ($)

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

6 4 2 0

1

2

3

Case Figure 9: Solvent costs for each case. Sulfolane is used in case one and IL is used in case two and three.

36


Page 36 of 59

Page 37 of 59

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 10: Percentage of each item that constitutes total annular cost

Figure 11: Comparison of total annual cost varying with IL price

37



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 ACS Paragon Plus Environment

Page 38 of 59

Page 39 of 59

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21




1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21


Page 40 of 59

Page 41 of 59

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



Benzene Industrial & Engineering Chemistry Research

Page 42 of 59

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23

10

90

20

80

30

70

40

60

50

50

60

40

70

30

80

20

90

IL 0

10 ACS Paragon Plus Environment

10

20

30

40

50

60

70

80

90

0

Cyclohexane

Toluene Industrial & Engineering Chemistry Research

Page 43 of 59

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23

10

90

20

80

30

70

40

60

50

50

60

40

70

30

80

20

90

IL 0


10

20

30

40

50

60

70

80

90

0

Cyclohexane

n-Heptane Industrial & Engineering Chemistry Research

Page 44 of 59

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21Toluene 22 23

10

90

20

80

30

70

40

60

50

50

60

40

70

30

80

20

90


0

10

20

30

40

50

60

70

80

90

0

IL

n-Heptane Industrial & Engineering Chemistry Research

Page 45 of 59

0 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21Toluene 22 23

10

90

20

80

30

70

40

60

50

50

60

40

70

30

80

20

90


0

10

20

30

40

50

60

70

80

90

0

IL


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


Page 46 of 59

Page 47 of 59

9

×10


8 7

Solvent Cost ($)

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

6

6 5 4 3 2 1 0

1


Case

3


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


Page 48 of 59

Page 49 of 59

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



Column 1


2.2 2

Objective function

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

Page 50 of 59

1.8 1.6 1.4 1.2 1

8

9

10

11ACS Paragon 12 Plus 13Environment 14 15

Number of stage

16

17

18

Column 2

Page 51 of 59


62 61

Objective function

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

60 59 58 57 56 55 54 40

41

42


Number of stage

48

49

50

Column 3


59

58

Objective function

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

Page 52 of 59

57

56

55

54 40

41

42


Number of stage

48

49

50

Column 1

Page 53 of 59


4.2 4

Objective function

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

3.8 3.6 3.4 3.2 3

9

10

11

12

13


Number of stage

14

15

Column 2


72

70

Objective function

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

Page 54 of 59

68

66

64

62 45

46

47


Number of stage

53

54

55

Column 3

Page 55 of 59


1.09

Objective function

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

1.085 1.08 1.075 1.07 1.065 1.06 1.055 1.05 20

21

22

ACS Paragon 23 24 Plus 25Environment 26 27

Number of stage

28

29

30

Column 1


46 45

Objective function

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

Page 56 of 59

44 43 42 41 40 39 38 35

36

37


Number of stage

43

44

45

Column 2

Page 57 of 59


2.55

Objective function

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

2.5

2.45

2.4

2.35

2.3 20

21

22


Number of stage

28

29

30

Column 3


2.5 2.4

Objective function

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

Page 58 of 59

2.3 2.2 2.1 2 1.9 10

11

12


Number of stage

18

19

20

Column 4

Page 59 of 59


9.5 9

Objective function

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

8.5 8 7.5 7 6.5 12

13

14

15


Number of stage

16

17

Conceptual Design of an Energy-Efficient Process for Separating

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