Alcohol Dehydration by Extractive Distillation with Use of Aminoethers of Boric Acid

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Article Alcohol Dehydration by Extractive Distillation with Use of Aminoethers of Boric Acid

Alexander V. Klinov *, Alexander V. Malygin , Alina R. Khairullina, Sergey E. Dulmaev and Ilsiya M. Davletbaeva * Department of Chemical Process Engineering, Kazan National Research Technological University, 68 Karl Marx St., 420015 Kazan, Russia; [email protected] (A.V.M.); [email protected] (A.R.K.); [email protected] (S.E.D.) * Correspondence: [email protected] (A.V.K.); [email protected] (I.M.D.)

Received: 14 October 2020; Accepted: 13 November 2020; Published: 16 November 2020

Abstract: Aminoethers of boric acid (AEBA) were studied as potential extractants for the separation of aqueous–alcoholic azeotropic mixtures by extractive distillation. The conditions of vapor–liquid equilibrium in aqueous solutions of ethanol and isopropanol in the presence of AEBA were studied. The division of AEBA molecules into group components was proposed, and previously unknown geometric parameters of the boron group and the energetic pair parameters of the boron group with the alkane group, ether group, amine-3d group, and alcohol group were determined within the framework of the Universal Functional Group Activity Coefficient (UNIFAC) model. The modeling of the extractive rectification process of an ethanol–water mixture with AEBA as extractant has been carried out. The dependences of the cost function on the extractant flow rate, the residual water content in it and the number of theoretical trays were obtained. A technological scheme for ethanol dehydration has been proposed, and its technological characteristics have been calculated.

Keywords: extraction; vapor–liquid equilibrium; aqueous solution; UNIFAC model

1. Introduction Anhydrous alcohols are widely used in the chemical industry as raw materials for the chemical synthesis of esters and ethers, as well as solvents in the production of paints, cosmetics, aerosols, perfumery, drugs, food and many other industrial areas. In addition, mixtures of anhydrous alcohols such as ethanol, isopropanol and gasoline are used as biofuels, increasing the octane number of gasoline and raising its combustion rate. Consequently, this reduces the level of carbon monoxide in the exhaust and reduces environmental pollution. In many cases of practical use of alcohols, their high purity is required, not less than 99 wt % [1]. It is known that many alcohols form azeotropic mixtures with water, for example, ethanol at a content of 95.5 wt %, and isopropanol at 87.4 wt %. Distillation (rectification) methods or their special types (extractive or azeotropic rectification) are traditionally used to separate water alcohol mixtures. Such separation methods are energy intensive, require the use of extractants − and additional equipment that are not always cheap, and the costs of their implementation largely determine the cost of the product (alcohol). Thus, this study is aimed at finding effective extractants capable of purposefully changing the relative volatility of the components of aqueous–alcoholic solutions, easily regenerating and having a low cost is relevant. Known extractants, solid salts and organic solvents used for the separation of water–alcohol mixtures have some disadvantages that limit their use for the production of anhydrous alcohols. The use of solid salts (solutions of salts of sodium nitrate, lithium chloride, sodium chloride, potassium chloride, calcium chloride, potassium iodide and potassium acetate) [2,3] has problems associated with their corrosiveness and regeneration. Organic solvents used for the dehydration of alcohols (ethylene glycol, diethyl ether, benzene, pentane and

Processes 2020, 8, 1466; doi:10.3390/pr8111466 www.mdpi.com/journal/processes Processes 2020, 8, 1466 2 of 21 glycerol) [4] are volatile and can contaminate the distillate. There are options for using salt solutions in organic solvents as extractants. Sometimes this approach gives a positive result during polar or non-polar systems separation [5–8]. Nevertheless, the low solubility of salts in organic solvents, with the remaining drawbacks listed above, limits the application of this approach. There are known examples of hyperbranched polymers usage for the separation of mixtures, ethanol–water and tetrahydrofuran–water, using extractive rectification and solvent extraction, respectively [9,10]. Among them, there are only a few hyperbranched polymers currently commercially available that meet the requirements for extractants (low viscosity, low melting point, thermal stability and selectivity). Recently, ionic liquids (IL) have also been used as separating agents in the process of extractive rectification [11–18]. Ionic liquids are a diverse group of salts that exist in liquid state at room temperature. Studies show that ILs are very effective extractants, but their industrial use is still limited by the high cost of their production. In this work, we investigated the possibility of using aqueous–alcoholic solutions of aminoethers of boric acid (AEBA) as effective extractants for extractive distillation. As shown in our previous work [19], AEBA exhibit some properties similar to the properties of ionic liquids: the melting point is below 100 ◦C, they are practically non-volatile, their aqueous solutions have high electrical conductivity, and the azeotrope is removable when added to aqueous–alcoholic solutions. In contrast to imidazole ionic liquids, AEBA have a much lower production cost, which makes them promising extractants for the separation of water–alcohol mixtures by rectification. To assess the efficiency of AEBA as extractive agents, the following problems were solved in this work: the conditions of vapor–liquid equilibrium in aqueous solutions of ethanol–water were identified, isopropanol–water in the presence of AEBA was studied, a method for calculating the activity coefficients of components in such mixtures was proposed, the calculations of the ethanol dehydration technological scheme were carried out, and its technological characteristics were determined.

2. Materials and Methods

2.1. Materials All glycols, i.e., diethylene glycol (DEG) and triethylene glycol (TEG) were purchased from PJSC Nizhnekamskneftekhim (Nizhnekamsk, Russia). Triethanolamine (TEA) was purchased from OJSC Kazanorgsintez. Boric acid (99.99 wt %) was purchased from Sigma-Aldrich. Glycols were additionally dehydrated at a vacuum depth of 1–3 mm Hg and a temperature of 90 ◦C to a moisture content less than 0.01 wt %. These conditions were chosen experimentally, and they ensure the constancy of the residual water content. For carrying out distillation experiments, ethanol (96.3 wt %) and isopropanol (99.8 wt %, JSC “EKOS-1”) were used. For the preparation of aqueous solutions and mixtures, deionized water was used, prepared on the Osmodemi 12 installation.

2.2. Synthesisprocess Aminoethers of boric acid based on di- and tri-ethylene glycol (AEBA-DEG/AEBA-TEG) were obtained in one step. The calculated amount of triethanolamine, boric acid and DEG/TEG was added to a three-necked round bottom flask at a molar ratio of [TEA]:[H3BO3]:[DEG/TEG] = 1:6:12. The use of such molar ratio was justified in previous work [19]. The mass of boric acid (6 mol) was 2.793 g, and that of triethanolamine (1 mol) was 1.124 g. The mass of DEG/TEG (12 mol) was 9.644 g/13.485 g, 1 respectively. The reaction mixture was heated to 90 ◦C with heating rate of 2 ◦C min− at a residual pressure of 10 mm Hg and was kept under these conditions for 2 h. This time has been chosen experimentally, and it ensures the constancy of the residual water content involved in the structural organization of AEBA. The vacuum was created with an oil pump connected to a U-shaped moisture trap filled with zeolite. Since the dissolution rate is greater than the reaction rate, the mixing was carried out by the natural bubbling of water released during the reaction. Processes 2020, 8, x FOR PEER REVIEW 3 of 22

The progress of the reaction was monitored by titration by determining the concentration of hydroxyl groups. When the number of hydroxyl groups reached the plateau, the synthesis was considered as completed. The synthesis product was placed into sealed jar. AEBA-TEG contained 3.96 wt % of water, and AEBA-DEG contained 4.97 wt % of water. The water content was measured using a volumetricProcesses titrator2020, 8, 1466 from Mettler Toledo V20 according to the Karl Fischer method.3 of 21

2.3. Preparation of AqueousThe progress Solutions of the reaction and wasMixtures monitored by titration by determining the concentration of hydroxyl groups. When the number of hydroxyl groups reached the plateau, the synthesis was For the preparationconsidered as completed. of AEBA The synthesisaqueous product soluti wasons placed and into mixtures, sealed jar. AEBA-TEG deionized contained water was used. Samples were 3.96prepared wt % of water, on a and ShincoADJ AEBA-DEG containedscales with 4.97 wt a %measurement of water. The water error content ofwas ±0.0001 measured g. using a volumetric titrator from Mettler Toledo V20 according to the Karl Fischer method.

2.4. Determination2.3. Preparationof the Boiling of Aqueous Point Solutions and Mixtures For the preparation of AEBA aqueous solutions and mixtures, deionized water was used. Samples To measurewere the prepared boiling on a point ShincoADJ of scalesbinary with mixtures a measurement, the error Swietoslavsky of 0.0001 g. ebulliometer were used [20] ± (JSC “Khimlaborpribor”).2.4. Determination The of the sche Boilingme Point of ebulliometer is shown in Scheme 1. The temperature was measured with anTo LT-300-N measure the boilingelectronic point of thermomete binary mixtures,r the (LLC Swietoslavsky Termeks) ebulliometer with an were error used [20of] ±0.05 °C. The thermometer was(JSC “Khimlaborpribor”).installed in pocket The scheme3 filled of with ebulliometer a water–ethylene is shown in Scheme glycol1. The mixture. temperature The was initial mixture measured with an LT-300-N electronic thermometer (LLC Termeks) with an error of 0.05 C. The was poured into cube 1 through refrigerator 5. The test mixture was heated by± a flexible◦ electric heater thermometer was installed in pocket 3 filled with a water–ethylene glycol mixture. The initial mixture mounted on thewas outer poured surface into cube of 1 throughcube 1. refrigerator The mixtur 5. Thee was test mixture brought was heatedto a boil, by a and flexible the electric boiling point was recorded as theheater value mounted that on theremained outer surface unchange of cube 1. Thed mixturefor 15 was min brought after to a boil,the andunit the boilingentered the mode. point was recorded as the value that remained unchanged for 15 min after the unit entered the mode. SimultaneouslySimultaneously with fixing with the fixing temperature, the temperature, the the liquid liquid phase phase was sampledwas sampled from the bottom from ofthe the bottom of the overflow tube overflow7 to clarify tube 7 the to clarify composition the composition of ofthe the boiling boiling mixture. mixture.

Scheme 1. SwietoslavskyScheme 1. Swietoslavsky ebulliometer: ebulliometer: 1—cube; 1—cube; 2—Cottrell 2—Cottrell pump; pump; 3—pocket 3—pocket for a thermometer; for a thermometer; 4— 4—separation space; 5—refrigerator; 6—drop counter; 7—overﬂow tube. separation space; 5—refrigerator; 6—drop counter; 7—overflow tube.

2.5. Phase Equilibrium Experiments To conduct experimental studies on the phase equilibrium of vapor–liquid in the alcohol–water– AEBA system, an approach based on the method of open evaporation was used, described in the work [19]. This approach provides a qualitative and quantitative assessment of the effect of the addition of an extractant on the conditions of phase equilibrium in an azeotropic mixture in a certain concentration range. To conduct experimental studies on the phase equilibrium of vapor–liquid, an IKA-RV 10 digital rotary evaporator was used. The experimental technique and the evaluation of the results obtained are described in detail in the work [19].

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2.5. Phase Equilibrium Experiments To conduct experimental studies on the phase equilibrium of vapor–liquid in the alcohol–water–AEBA system, an approach based on the method of open evaporation was used, described in the work [19]. This approach provides a qualitative and quantitative assessment of the eﬀect of the addition of an extractant on the conditions of phase equilibrium in an azeotropic mixture in a certain concentration range. To conduct experimental studies on the phase equilibrium of vapor–liquid, an IKA-RV 10 digital rotary evaporator was used. The experimental technique and the evaluation of the results obtained are described in detail in the work [19].

2.6. Phase Equilibrium Calculation To calculate the activity coeﬃcients of the components of the mixture, the UNIFAC model was used [21]. The UNIFAC model is a universal pseudo-chemical equation for calculating the activity coeﬃcients of the components of a solution based on data on the functional groups that make up the molecules of substances. This makes it possible to predict the conditions of vapor–liquid and liquid–liquid phase equilibrium for systems that have not been studied experimentally. Within the UNIFAC model, the activity rate is determined as follows:

C R lnγi = lnγi + lnγi (1)

The ﬁrst term is called the combination contribution, and the second is called the residual. The value of the combination contribution depends only on the solution composition, as well as on the volume and surface of its constituent groups:

Φ z θ Φ X lnγC = ln i + q ln i + l i x l , (2) i x 2 i Φ i − x j j i i i j where xi is the component molar fraction of the in the solution; Φi and θi are the volume and surface fractions of a component in a solution:

qixi rixi θi = P ; Φi = P , (3) j qjxj j rjxj

The parameters of the pure components ri and qi characterize the molecular Van der Waals volumes and surface areas of the molecules, respectively. These parameters are determined by summing the group parameters of the volume (R) and surface (Q): X X = (i) = (i) rj νk Rk; qi νk Qk, (4) k k

(i) where vk is an integer and determines the number of groups of type k in molecule i. The residual (energy) part of the activity coeﬃcient in group models is represented by the sum of the contributions of the groups included in the molecule: X (i) (i) lnγR = v lnΓ lnΓ , (5) i k k − k

(i) where Γk—residual activity coeﬃcients of the group (k) in solution; Γk is the residual activity coeﬃcients of the group (k) in some comparative solution containing only type i molecules. Processes 2020, 8, 1466 5 of 21

(i) The activity coeﬃcients of the group in the solution Γk and Γk are determined by the equation:    X X θmψkm  = Q  ( )  ln Γk k1 ln θmψmk P , (6) − m − m n θnψnm where θm is the surface fraction of group m, and the sums include all the diﬀerent groups:

QmXm θm = P , (7) n QnXn where Xm is the molar fraction of group m in the mixture. The group interaction parameter ψmn is determined by the equation: ! Umn Unn amn ψmn = exp − = exp , (8) − RT − T where amn–characterizes the interaction between groups n and m; for each group–group interaction, two parameters amn , amn are used. The UNIFAC group constituent method allows one to calculate the properties of various substances using a limited number of parameters characterizing the contributions of individual atoms or groups. Therefore,Processes this2020, method8, x FOR PEER requires REVIEW dividing the molecule into some standard set of groups. 5 of 22

2.7.2.7. Simulation Simulation of the of the Extractive Extractive Distillation Distillation Process Process ModelingModeling of theof the extractive extractive distillation distillation process of of water–alcohol water–alcohol mixtures mixtures was was carried carried out in out the in theUniSim package package (UniSim (UniSim Design Design Academic Academic Program). Program). The The calculations calculations were were carried carried out by out the by thetheoretical tray tray method. method. AEBA-TEG AEBA-TEG was was investig investigatedated as a as hypothetical a hypothetical component component through through the the UNIFAC framework. The The UNIFAC UNIFAC model model was was used used to to calculate calculate the the activity activity coefficients coeﬃcients of of the the mixturemixture components. components.

3. Results3. Results

3.1.3.1. Vapor VaporLiquid−Liquid Equilibrium Equilibrium in in Aqueous Aqueous Solutions Solutions of of Alcoholsin Alcoholsin thethe PresencePresence of AEBA − TheThe synthesis synthesis of AEBAof AEBA is carriedis carried out out according according to to the the following following scheme scheme (Figure(Figure1 1):):

Figure 1. Scheme of synthesis of aminoethers of boric acid based on triethylene glycol AEBA-TEG.

Glycols of various structures can be used for synthesis. In this work, variants with di- and triethylene glycol (DEG and TEG) are considered, which makes it possible to obtain AEBA-DEG and AEBA-TEG, respectively. The UNIFAC model was used to describe the conditions of phase equilibrium in the alcohol–water–AEBA systems [21]. In this case, the activity coefficients were calculated from the group constituents parameters of the mixture molecules. AEBA molecules synthesized with various glycols differ only in the number of groups of the same type, the parameters of which can be determined based on specific AEBA, for example, AEBA-DEG and AEBA-TEG. In this case, the UNIFAC model has the potential to predict the conditions of phase equilibrium in aqueous alcohol solutions with AEBA of different molecular composition. Within the UNIFAC model framework, molecules of substances are separated into group components. For water molecules and various alcohols, such a division has already been proposed [21]. Here it is proposed to divide the AEBA molecules into the following groups: B–boron group, СН2–alkane group, OCH2–ether group, NCH2–amine-3d group, and OH–alcohol group. The details of the split are shown in Figure 1. Thus, AEBA-TEG consists of 35–CH2, 42–OCH2, 6–B, 9–OH and 1– NCH2 groups, and AEBA-DEG consists of 23–CH2, 30–OCH2, 6–B, 9–OH and 1–NCH2 groups. The molecular weights of AEBA-DEG and AEBA-TEG are 1470 and 1998, respectively. Analysis of the literature [21,22] and the most complete database on the group interaction parameters for the UNIFAC model in the form of UNIFAC Matrix 2020 presented in the Dortmund

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Glycols of various structures can be used for synthesis. In this work, variants with di- and tri-ethylene glycol (DEG and TEG) are considered, which makes it possible to obtain AEBA-DEG and AEBA-TEG, respectively. The UNIFAC model was used to describe the conditions of phase equilibrium in the alcohol–water–AEBA systems [21]. In this case, the activity coefficients were calculated from the group constituents parameters of the mixture molecules. AEBA molecules synthesized with various glycols differ only in the number of groups of the same type, the parameters of which can be determined based on specific AEBA, for example, AEBA-DEG and AEBA-TEG. In this case, the UNIFAC model has the potential to predict the conditions of phase equilibrium in aqueous alcohol solutions with AEBA of different molecular composition. Within the UNIFAC model framework, molecules of substances are separated into group components. For water molecules and various alcohols, such a division has already been proposed [21]. Here it is proposed to divide the AEBA molecules into the following groups: B–boron group, CH2–alkane group, OCH2–ether group, NCH2–amine-3d group, and OH–alcohol group. The details of the split are shown in Figure1. Thus, AEBA-TEG consists of 35–CH 2, 42–OCH2, 6–B, 9–OH and 1–NCH2 groups, and AEBA-DEG consists of 23–CH2, 30–OCH2, 6–B, 9–OH and 1–NCH2 groups. The molecular weights of AEBA-DEG and AEBA-TEG are 1470 and 1998, respectively. Analysis of the literature [21,22] and the most complete database on the group interaction parameters for the UNIFAC model in the form of UNIFAC Matrix 2020 presented in the Dortmund Data Bank (DDB) [23] showed the absence of data on the boron group. Therefore, in this work, based on the available experimental data, the geometric parameters of the boron group and the interaction energy parameters of this group with other groups of the considered solutions were determined. These are B-CH2, B-OCH2, B-NCH2, B-OH, and B-H2O. To determine the configurational contribution (2) to the activity coefficient (1), the group parameters R and Q data are required, which are related to the values of the Van der Waals group volume Vk and the surface area Ak (Van der Waals group volume and surface areas) [21,24]:

Vωk Aωk Rk = ;Qk = (9) 15.17 2.5 109 · The value of the Van der Waals radii rW for the boron atom was determined in diﬀerent 0 0 sources [25–27] in the range 1.65–1.92 A . Here, the value rW = 1.8 A was taken for boron. Hence, the values RK = 0.9371 and QK = 0.9809. The group interaction parameters amn in the group interaction parameter ψmn (8) were determined from the experimental VLE data using a regression analysis algorithm based on the least squares method. For binary systems, the objective function is:

N Xh i2 F = yE yC min (10) K − K → k

x ln(γ)Ps(T) yc = (11) P here y and x are the concentrations of component in vapor and liquid, respectively; PS(T) is the saturated vapor pressure of the pure component; superscripts E and C designate experimental and calculated data; k and N are the number and quantity of experimental points, respectively. The available data on vapor–liquid equilibrium (VLE data) for binary solutions of borates were analyzed. DDB analysis [28] showed that VLE data with borate solutions are rather limited. VLE data turned out to be useful for binary systems trimethyl borate–cyclohexane [29] and trimethyl borate–n-heptane [30]. These data were used to determine the parameters B-CH2, B-OCH2 (Table1). The data on the saturated vapor pressure of trimethyl borate (TMB), cyclohexane, and n-heptane were taken from the NIST database [31]. To demonstrate the description reliability of the phase equilibrium condition, Figure2 shows a comparison of the calculated data according to the UNIFAC model, Processes 2020, 8, 1466 7 of 21 obtained with new parameters (Table1), with known experimental data. The average deviation of the calculated concentrations from the experimental ones was 3%, and for temperatures, it was 1%. − Table 1. Boron group interaction parameters (Bn).

m amn anm CH 170.60 384.58 2 − OCH2 405.99 1825.92 OH 281.82 722.30 − − NCH 113.96 13.53 2 − H2O 237.83 1136.35 Processes 2020, 8, x FOR PEER REVIEW − − 7 of 22

y, mol frac TMB

0.8

0.6

0.4

0.2

0 0.2 0.4 0.6 0.8 1 x, mol frac TMB Figure 2. EquilibriumEquilibrium compositions of vapor and liquid for a mixture mixture of trimethyl trimethyl borate–cyclohexane (circles,(circles, solidsolid line)line) andand trimethyltrimethyl borate–n-hexaneborate–n-hexane (rhombuses,(rhombuses, dasheddashed lines).lines). Geometric ﬁguresfigures are experimental data, and lines areare calculatedcalculated data.data.

The remaining remaining unknown unknown parameters parameters were were determined determined from the from experimental the experimental data on the data boiling on thepoints boiling at atmospheric points at atmospheric pressure of pressure binary ofsolutions binary solutions AEBA-DEG/AEBA-TEG–ethanol, AEBA-DEG/AEBA-TEG–ethanol, AEBA- AEBA-DEGDEG/AEBA-TEG–isopropanol,/AEBA-TEG–isopropanol, and AEBA-DEG/AEBA and AEBA-DEG/-TEG–water.AEBA-TEG–water. The experimental The experimental data were data wereobtained obtained using usingthe Swietoslawski the Swietoslawski setup [20]. setup The [20 accuracy]. The accuracy of determining of determining the boiling the boilingpoints of points pure ofwater pure and water alcohols and alcoholsusing this using setup this was setup 0.5%. was Figure 0.5%. 3 shows Figure 3the shows boiling the points boiling of pointsaqueous of aqueoussolutions solutionsof AEBA. ofAqueous AEBA. solutions Aqueous of solutions ionic liquids of ionic have liquids similar have behavior, similar but behavior, with a butreduced with role a reduced of the rolenon-volatile of the non-volatilecharacter of the character compound. of the As compound.previously revealed, As previously the volatility revealed, of AEBA-DEG/AEBA- the volatility of AEBA-DEGTEG in comparison/AEBA-TEG with water in comparison and alcohol with can water be negl andected alcohol [19]; therefore, can be neglected the activity [19 coefficient]; therefore, of thewater activity or alcohol coeﬃ cientin a mixture of water with or alcohol AEBA incan a mixturebe determined with AEBA from can the be known determined boiling from point the TS known of the boilingmixture point as TS of the mixture as P ln( ) = (12) ln(𝛾)γ = s xP(()Ts) (12)

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First, the parameters B-NCH2 and B-OH were determined by the boiling points of AEBA-DEG/AEBA-TEG–ethanol and AEBA-DEG/AEBA-TEG–isopropanol. Then, the parameter B–H2O was determined from the boiling point of AEBA-DEG/AEBA-TEG–water. The results of comparing the calculated and experimental data are shown in Figures3 and4. The average discrepancy was 1%. The deﬁned parameters of the boron group interaction are presented in Table1. The adequacy veriﬁcation of the description of the vapor–liquid equilibrium of aqueous–alcoholic solutions in the presence of AEBA was carried out by comparing the calculated and experimental data on open evaporation. Open evaporation equation: dx y(x)∗ = (1 e) + x (13) −d(e) −

where e = P/L0; P is the amount of distillate withdrawn; L0 is the initial mass of the mixture in the still. Processes 2020, 8, xThe FOR experimental PEER REVIEW setup and the experimental procedure are described in detail in [19]. 8 of 22

120

T, 0C

115

110

105

100

0 0.2 0.4 0.6 0.8 x, mass frac AEBA Figure 3. BoilingFigure 3. pointsBoiling pointsof aqueous of aqueous AEBA AEBA solutions solutions atat atmospheric atmospheric pressure. pressure. Geometric Geometric figures are figures are experimentalexperimental values (circles—AEBA-diethylene values (circles—AEBA-diethylene glycol glycol (DEG); (DEG); triangles—AEBA-TEG), triangles—AEBA-TEG), and lines and are lines are calculated values. calculated values. Figure5 shows the results of comparison of the curves of residual compositions obtained for the ethanol–water mixture at different concentrations of AEBA-TEG. A good reproduction of the experimental data calculations using the UNIFAC model with the proposed parameters can be seen. 120 Figure5 also shows the curves of residual compositions for the ethanol–water mixture in the absence of AEBA forT, comparison, 0C which shows the effect of AEBA on the relative volatility of ethanol. Experiments

110

100

00.20.40.60.8 x, mass frac AEBA Figure 4. Boiling points of alcoholic solutions of AEBA at atmospheric pressure. Geometric figures are experimental values (circles—AEBA-DEG–ethanol; triangles—AEBA-TEG–ethanol; rhombuses— AEBA-TEG–isopropanol), and lines are calculated values.

The defined parameters of the boron group interaction are presented in Table 1.

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120

T, 0C

115

110 Processes 2020, 8, 1466 9 of 21

and calculation results have shown that the residual composition of the ethanol–water mixture depends on the type of added AEBA (based on TEG or DEG). The effect of AEBA-DEG on the relative volatility 105 of ethanol is, on average, 10% stronger. To assess the efficiency of using AEBA, we compared the coefficients of the relative volatility of ethanol in an aqueous solution in the presence of AEBA, glycerol and propylene glycol at their content of 30 wt % (Figure6). The relative volatility coefficient was determined as: 100 , y1 y2 α12 = (14) x1 x2

where the y1 and y2 equilibrium concentration of ethanol and water in steam above the solution was 95 determined by the UNIFAC model; x1 and x2 are concentrations of ethanol and water in solution. Figures6 and7 show0 a comparison 0.2 of the AEBA separation 0.4 ability with 0.6 known entrainers 0.8 currently used to separate azeotropic alcohol mixtures (glycerol and propylene glycol).x, mass The frac results AEBA in Figures6 and7 show that in the entire range of the considered compositions, with the addition of AEBA, Figurethe 3. relativeBoiling volatility points coeof ﬃaqueouscient of alcoholAEBA issolutions 1.5–2 times at higher, atmospheric especially pressure. in the region Geometric of high alcohol figures are experimentalconcentrations values near (circles—AEBA-diethylene the azeotrope point. The dependence glycol (DEG); nature oftriangles—AEBA-TEG), the relative volatility coeﬃ andcient lines in are calculatedthe presence values. of AEBA is similar to some imidazole ILs ([mim][DMP] and [mim][Cl]).

120

T, 0C

110

100

00.20.40.60.8 x, mass frac AEBA Figure 4. BoilingFigure 4. pointsBoiling of points alcoholic of alcoholic solutions solutions of AEBA of AEBA at atmospheric at atmospheric pressure. pressure. GeometricGeometric figures ﬁgures are experimental values (circles—AEBA-DEG–ethanol; triangles—AEBA-TEG–ethanol; are experimental values (circles—AEBA-DEG–ethanol; triangles—AEBA-TEG–ethanol; rhombuses— rhombuses—AEBA-TEG–isopropanol), and lines are calculated values. AEBA-TEG–isopropanol), and lines are calculated values.

The defined parameters of the boron group interaction are presented in Table 1.

Processes 2020, 8, 1466 10 of 21 ProcessesProcesses 20202020,, 88,, xx FORFOR PEERPEER REVIEWREVIEW 1010 of 2222

0.9 X ethanol mass frac

0.8

0.7

0.6

0.5

0 0.2 0.4 0.6 0.8 1 P/L 0 FigureFigureFigure 5. 5. 5.Residual ResidualResidual curves curves curves for the for for ethanol–water–AEBA-TEG thethe ethanol–water–AEethanol–water–AEBABA-TEG mixture.-TEG Geometricmixture. mixture. figuresGeometric Geometric are experimental figures figures areare valuesexperimentalexperimental (circles—AEBA-TEG values values (circles—AEBA-TEG (circles 20— wtAEBA %, rhombuses—AEBA-TEG-TEG 20 20 wtwt %, %, rhombuses—AEBA-TEG rhombuses 40 wt— %,AEBA triangles—AEBA-TEG-TEG 40 40 wtwt %, %, triangles— triangles 60 wt %),— andAEBA-TEGAEBA continuous-TEG 6060 lineswtwt %),% are), andand calculated continuouscontinuous values. lineslines Dotted areare calculatedcalculated lines—calculated values.values. values DottedDotted for lines—calculatedlines the— ethanol–watercalculated valuesvalues mixture forfor withoutthethe ethanol–waterethanol AEBA.–water mixturemixture withoutwithout AEBA.AEBA.

3.2

2.8

AEBA-TEG 2.4 AEBA-DEG Glycerol ? 1,2-propylene glycol

1.6

1.2

0.50.60.70.80.91 x mas. frac. ethanol FigureFigureFigure 6. 6.6. TheThe coecoefficientcoefficientﬃcient ofofof thethethe componentscomponentscomponents relativerelativerelative volatilityvolatilityvolatility ofofof thethethe ethanol–waterethanol–waterethanol–water mixturemixturemixture ininin thethethe presencepresencepresence of ofof various variousvarious extractants. extractants.extractants.

FiguresFigures 66 andand 77 show show a a comparison comparison of of the the AEBA AEBA separationseparation ability ability with with known known entrainers entrainers currentlycurrently usedused toto separateseparate azeotropicazeotropic alcoholalcohol mixturesmixtures (glycerol(glycerol andand propylenepropylene glycol).glycol). TheThe resultsresults inin FiguresFigures 66 andand 77 showshow thatthat inin thethe entireentire rangerange ofof thethe consideredconsidered compositions,compositions, withwith thethe additionaddition ofof

Processes 2020, 8, x FOR PEER REVIEW 11 of 22

AEBA, the relative volatility coefficient of alcohol is 1.5–2 times higher, especially in the region of high Processesalcohol2020 concentrations, 8, 1466 near the azeotrope point. The dependence nature of the relative11 volatility of 21 coefficient in the presence of AEBA is similar to some imidazole ILs ([mim][DMP] and [mim][Cl]).

FigureFigure 7. The 7. Thecoefficient coefficient of the of the components components relative relative volatility of of the the isopropanol–water isopropanol–water mixture mixture in the in the presencepresence of various of various extractants. extractants. 3.2. The Process of Extractive Distillation of Ethanol–Water Mixtures Using AEBA-TEG as an Extractant 3.2. The Process of Extractive Distillation of Ethanol–Water Mixtures Using AEBA-TEG as an Extractant To analyze the process of extractive rectification of an ethanol–water mixture, AEBA-TEG was chosenTo analyze as an extractivethe process agent. of extractive The choice rectification was determined of byan theethanol fact that–water AEBA-DEG mixture, and AEBA AEBA-TEG-TEG was chosenhave as similaran extractive indicators agent. of influence The choice on thewas relative determined volatility by ofthe water fact that and alcoholAEBA-DE (FiguresG and6– 9AEB), A- TEG buthave AEBA-TEG similar indicators has a higher of influence temperature on ofthe thermal relative decomposition volatility of water [19]. Since and thealcohol purpose (Figure of thes 6–9), but AEBAresearch-TEG was has to assess a higher the energy temperature and material of thermal costs of decomposition carrying out the [19]. process Since without the purpose the need toof the researchperform was design to assess calculations, the energy the calculationsand material of thecosts rectification of carrying process out werethe process carried outwithout on the the basis need of to performthe theoretical design calculations, tray model [ 32the]. Thecalculations theoretical of tray the method rectification takes into process account were only carried the thermodynamic out on the basis of thefactor theoretical and makes tray it possible model to [3 determine2]. The the theoretical fundamental tray possibility method and takes estimate into theaccount complexity only the of the mixture separation. The actual dimensions of the column depend on the rate of mass transfer thermodynamic factor and makes it possible to determine the fundamental possibility and estimate between the boundary and the core of the phases. The mass transfer coefficient is usually characterized the complexity of the mixture separation. The actual dimensions of the column depend on the rate of by the mass transfer coefficient or Murphree efficiency. These values largely depend on the physical massproperties transfer (viscosity, between di theffusion boundary and density) and the of the core vapor of andthe liquidphases. phases. The massDuring transfer distillation, coefficient when is usuallythe viscositycharacterized of the liquidby the phase mass is transfer low (1–10 coefficient mPa s), the or mass Murphree transfer efficiency. in the vapor These phase values is limiting. largely · dependThe on presence the physical of AEBA properties can significantly (viscosity, increase diffusion the viscosity and of density) the liquid of mixture. the vapor However, and liquid as shown phases. Duringby measurementsdistillation, wh [19en] for the conditions viscosity in of a distillationthe liquid column,phase is where low the(1–10 concentration mPa∙s), the of mass AEBA transfer in the in the vaporliquid phase mixture is limiting. is about 30–50 The presence wt % and of at AEBA a temperature can significantly of about 80 increase◦C, the the viscosity viscosity of the of liquidthe liquid mixture.phase However, will not exceed as shown 2 mPa bys. measurements It is hoped that [19] under for these conditions conditions, in a distillation the main resistance column, to where mass the · concentrationtransfer will of remain AEBA in in the the vapor liquid phase. mixture Thus, is the about design 30 diagram–50 wt of% the and distillation at a temperature column (Figure of about 10) 80 °C, thewas viscosity built on theof the basis liquid of a systemphase will of material not exceed and heat2 mPa balance∙s. It is equations, hoped that as wellunder as these equations conditions for , vapor–liquid phase equilibrium. Activity coefficients were calculated using the UNIFAC model with the main resistance to mass transfer will remain in the vapor phase. Thus, the design diagram of the the parameters given in Table1. distillation column (Figure 10) was built on the basis of a system of material and heat balance equations, as well as equations for vapor–liquid phase equilibrium. Activity coefficients were calculated using the UNIFAC model with the parameters given in Table 1.

Processes 2020, 8, 1466 12 of 21 Processes 2020, 8, x FOR PEER REVIEW 12 of 22

Processes 2020, 8, x FOR PEER REVIEW 13 of 22 FigureFigure 8. The 8. coefficientThe coeﬃcient of relative of relative volatility volatility depending depending on on the concentrationconcentration of of the the extractant extractant at the at the azeotropic azeotropic point point of the of theethanol ethanol–water–water mixture. mixture.

FigureFigure 9. The 9. Thecoefficients coeﬃcients of the of the relative relative volatility volatility depending depending onon the the concentration concentration of theof the extractant extractant at at the azeotropicthe azeotropic point point of the of the isopropanol isopropanol–water–water mixture. mixture.

Figure 10. Calculation scheme of the extractive distillation column: 1—distillation column, 2— separator, 3—boiler, 4—reflux condenser.

Initially, the influence of the extractant amount supplied, reflux ratio and number of theoretical trays on the distillate composition was investigated. For the calculations, the following conditions were used:

- Feed composition XF (0.9 mass fraction of ethanol; 0.1 mass fraction of water), which is close to the azeotropic point. The required alcohol content in the distillate XdE => 99.5 wt %; alcohol in the still XWe < 10−7 wt %. - The pressure in the column was equal to atmospheric pressure; the pressure drop across the column was neglected. - NT is the total number of trays; NF is feed tray. Due to the fact that the volatility of AEBA is negligible, the extractant flow was fed to the first tray NE = 1.

Processes 2020, 8, 1466 13 of 21 Processes 2020, 8, x FOR PEER REVIEW 13 of 22

Figure 10. Calculation scheme of the extractive distillation column: 1—distillation column, 2—separator, Figure 10.3—boiler, Calculation 4—reflux condenser.scheme of the extractive distillation column: 1—distillation column, 2— separator, 3—boiler, 4—reflux condenser. Initially, the influence of the extractant amount supplied, reflux ratio and number of theoretical trays on the distillate composition was investigated. For the calculations, the following conditions Initially,were used: the influence of the extractant amount supplied, reflux ratio and number of theoretical trays on the distillate composition was investigated. For the calculations, the following conditions - Feed composition XF (0.9 mass fraction of ethanol; 0.1 mass fraction of water), which is close to were used: the azeotropic point. The required alcohol content in the distillate XdE => 99.5 wt %; alcohol in 7 the still XWe < 10− wt %. - Feed composition XF (0.9 mass fraction of ethanol; 0.1 mass fraction of water), which is close to - The pressure in the column was equal to atmospheric pressure; the pressure drop across the the azeotropiccolumn was point. neglected. The required alcohol content in the distillate XdE => 99.5 wt %; alcohol in −7 the -still NXTWeis < the 10 total wt number %. of trays; NF is feed tray. Due to the fact that the volatility of AEBA is - The pressurenegligible, in thethe extractant column flow was was equal fed to to the atmo first trayspheric NE = 1. pressure; the pressure drop across the column- Complete was neglected. condensation of vapor occurred in the reflux condenser; complete evaporation of the liquid mixture occurred in the cube. - NT is the total number of trays; NF is feed tray. Due to the fact that the volatility of AEBA is - R is the reflux ratio. negligible, the extractant flow was fed to the first tray NE = 1. - CompleteThe condensation flow rate of the of initial vapor mixture occurred F was conventionallyin the reflux assumedcondenser; to be complete 100 kg/h, sinceevaporation the of the calculation results are influenced not by the flow rate itself, but by its ratio to the extractant flow rate liquidef = mixtureE/F. The feedoccurred tray was in selected the cube. from the condition of the maximum concentration of ethanol in - R isthe the distillate reflux underratio. other specified conditions. Figure 11 shows the effect of reflux ratio and extractant flow rate on distillate composition in Thea flow column rate with of 18 the theoretical initial trays. mixture It can F be was seen thatconv theentionally required alcohol assumed content to in be the 100 distillate kg/h, since the calculation(99.5 results wt %) is are achieved influenced for the extractantnot by the consumption flow rate starting itself, frombut efby= its0.3. ratio With to an the increase extractant in the flow rate ef = E/F. consumptionThe feed tray of the was extractant, selected a smaller from amount the condition of reflux is requiredof the maximum to achieve a given concentration composition ofof ethanol in the distillatethe distillate. under Forother all consideredspecified cases,conditions. the value of the reflux ratio was less than 1. The dependences from Figure9 have a characteristic maximum associated with the fact that with an increase in the reflux Figureratio, 11 the shows amount the of reflux effect supplied of reflux increases, ratio and and when extractant mixed with flow the extractant,rate on distillate the concentration composition in a column withof the 18 latter theoretical decreases. trays. It can be seen that the required alcohol content in the distillate (99.5 wt %) is achievedFigure 12 for shows the the extractant calculations consumption for three different starting values of from the number ef = of0.3. theoretical With an trays increase in in the consumptionthe column: of the N extractant,T = 18; NT = a25 smaller and NT = amount30. It is seenof reflux that with is required an increase to in achieve the number a given of trays, composition the required composition of the distillate can be obtained at a lower consumption of the extractant. of the distillate.At the same For time, all considered increasing the cases, number the of trays value by moreof the than reflux 25 is impracticalratio was dueless to than a slight 1. The change dependences from Figurein the 9 reflux have ratio, a characteristic which makes it possiblemaximum to reduce associated capital costs. with the fact that with an increase in the reflux ratio, the amount of reflux supplied increases, and when mixed with the extractant, the concentration of the latter decreases.

Processes 2020, 8, 1466 14 of 21 Processes 2020, 8, x FOR PEER REVIEW 14 of 22

0.99

0.98

ef=0.2 ef=0.3 ef=0.4 ef=0.6

0.97

0.5 1 1.5 2 2.5 3 R FigureFigure 11. 11. InfluenceInfluence of of reflux reflux ratio ratio on on ethanol ethanol concentr concentrationation in indistillate distillate with with different different amounts amounts of extractant.of extractant. Processes 2020, 8, x FOR PEER REVIEW 15 of 22 Figure 12 shows the calculations for three different values of the number of theoretical trays in 1.6 the column: NT = 18; NT = 25 and NT = 30. It is seen that with an increase in the number of trays, the required composition of the distillate can be obtained at a lower consumption of the extractant. At N =25 the same time, increasing the number of trays by more than 25 isT impractical due to a slight change NT=18 in the reflux ratio, which makes it possible to reduce capital costs.NT=30

1.2

R 0.8

0.4

0.2 0.3 0.4 0.5 0.6 ef

Figure 12. ChangeChange in in reflux reﬂux ratio ratio from from the the amount amount of of ex extractanttractant supplied supplied with with a different a diﬀerent number number of traysof trays (NT). (NT).

The technology for the dehydration of ethanol, together with an extractive rectification unit, should include a unit for recovering the extractant from its aqueous solution. To extract AEBA from an aqueous solution, various processes can be used, as well as their combination [33]. Often, various types of distillation are used for these purposes: distillation under vacuum, stripping with inert gas, and molecular distillation. The efficiency of distillation processes can be increased by combining them with membrane separation or adsorption. The choice of a specific scheme for AEBA regeneration requires a comparative technical and economic analysis. Since the key task in this work was to study the process of extractive rectification, we did not perform such an analysis. Distillation under vacuum was chosen for the AEBA regeneration, as it is simple enough for modeling and traditionally used in many studies of extractive schemes for the dehydration of alcohols. The latter allows the comparison of costs for the entire ethanol recovery scheme. Considering that the volatility of AEBA-TEG compared to water is negligible, the regeneration process can be carried out by simple distillation. However, it is necessary to bear in mind a significant increase in the boiling point of the AEBA solution at a low water concentration and the onset temperature of the AEBA-TEG, which is about 190 °C [19]. Figure 13 shows saturation pressure isotherms versus water concentrations in a solution with AEBA-TEG. The most acceptable conditions for the industrial version of the AEBA- TEG aqueous solution distillation process are temperatures in the range of 110–150 °C and a pressure of 0.05–0.3 atm.

Processes 2020, 8, 1466 15 of 21

The technology for the dehydration of ethanol, together with an extractive rectification unit, should include a unit for recovering the extractant from its aqueous solution. To extract AEBA from an aqueous solution, various processes can be used, as well as their combination [33]. Often, various types of distillation are used for these purposes: distillation under vacuum, stripping with inert gas, and molecular distillation. The efficiency of distillation processes can be increased by combining them with membrane separation or adsorption. The choice of a specific scheme for AEBA regeneration requires a comparative technical and economic analysis. Since the key task in this work was to study the process of extractive rectification, we did not perform such an analysis. Distillation under vacuum was chosen for the AEBA regeneration, as it is simple enough for modeling and traditionally used in many studies of extractive schemes for the dehydration of alcohols. The latter allows the comparison of costs for the entire ethanol recovery scheme. Considering that the volatility of AEBA-TEG compared to water is negligible, the regeneration process can be carried out by simple distillation. However, it is necessary to bear in mind a significant increase in the boiling point of the AEBA solution at a low water concentration and the onset temperature of the AEBA-TEG, which is about 190 ◦C[19]. Figure 13 shows saturation pressure isotherms versus water concentrations in a solution with AEBA-TEG. The most acceptable conditions for the industrial version of the AEBA-TEG aqueous solution distillation Processesprocess 2020, 8, arex FOR temperatures PEER REVIEW in the range of 110–150 ◦C and a pressure of 0.05–0.3 atm. 16 of 22

0.5

130оС 110оС 150оС 0.4

0.3

0.2

0.1

0.005 0.01 0.015 0.02 0.025 0.03 x , mas. frac. water w FigureFigure 13. Isotherms 13. Isotherms of saturated of saturated water water vapor vapor pressure pressure overover a binarya binary water–AEBA-TEG water–AEBA-TEG solution, solution, dependingdepending on the on composition. the composition.

The residual water content in the regenerated AEBA-TEG will be 1–3 wt %. The water content in Thethe residual extractant water will negatively content ainﬀ ectthe the regenerated separation e ﬃAEBA-TEGciency of the will water–alcohol be 1–3 wt mixture; %. The however,water content in the withextractant an increase will in thenegatively water content affect in the the extractant, separation the costs efficiency of its regeneration of the willwater–alcohol decrease. These mixture; however,costs with are associated an increase with thein depththe water of the content vacuum andin the temperatureextractant, of the the heatingcosts of medium. its regeneration The total will decrease. These costs are associated with the depth of the vacuum and the temperature of the heating medium. The total cost of obtaining dehydrated alcohol by the extractive rectification technology can be represented by the following function: =α +β ++η++λ Q E (R11 ) N(R ) PV (15) The following are the cost items: - E is the amount of fresh extractant supplied, which is equal to the possible losses for the regeneration scheme. We will assume that these costs are negligible compared to others. - (R + 1) is the costs of heat in the evaporator cube of the extractive rectification column. - N(R + 1) capital costs associated with the size of the extractive rectification column; the number of trays N is proportional to the height of the column; (R + 1) is proportional to the steam flow rate, which determines the diameter of the column. - Pv is the costs of extractant regeneration associated with the depth of the vacuum and the corresponding material and operating costs. - α, β, η and λ are normalization coefficients. If the composition of the distillate and the cube is specified by the alcohol–water content, then the cost function (15) will depend on three variables: the specific consumption of the extractant, the number of trays or the reflux ratio in the extractive rectification column, and the water content in the extractant after regeneration. It is easy to see that the cost items for heat in the cube of the evaporator of the extractive rectification column and for regeneration of the extractant are proportional to the reflux ratio and water content in the extractant, respectively. Therefore, here, we provided studies of the independent variables’ influence on capital costs, defined as N(R + 1). Figures 14 and 15 show the isolines of costs at constant flow rates of the extractant and the number of trays, respectively, in the absence of water in the extractant. In this case, the cost function of two variables does not have a

Processes 2020, 8, 1466 16 of 21 cost of obtaining dehydrated alcohol by the extractive rectiﬁcation technology can be represented by the following function: Q = αE + β(R + 1) + ηN(R + 1) + λPV (15) The following are the cost items:

- E is the amount of fresh extractant supplied, which is equal to the possible losses for the regeneration scheme. We will assume that these costs are negligible compared to others. -( R + 1) is the costs of heat in the evaporator cube of the extractive rectification column. - N(R + 1) capital costs associated with the size of the extractive rectification column; the number of trays N is proportional to the height of the column; (R + 1) is proportional to the steam flow rate, which determines the diameter of the column.

- PV is the costs of extractant regeneration associated with the depth of the vacuum and the corresponding material and operating costs. - α, β, η and λ are normalization coeﬃcients.

If the composition of the distillate and the cube is specified by the alcohol–water content, then the cost Function (15) will depend on three variables: the specific consumption of the extractant, the number of trays or the reflux ratio in the extractive rectification column, and the water content in the extractant after regeneration. It is easy to see that the cost items for heat in the cube of the evaporator of the extractive rectification column and for regeneration of the extractant are proportional to the reflux ratio and water content in the extractant, respectively. Therefore, here, we provided studies of the independent variables’ influence on capital costs, defined as N(R + 1). Figures 14 and 15 show the isolines of costs at constant flow rates of the extractant and the number of trays, respectively, in Processes 2020, 8, x FOR PEER REVIEW 17 of 22 the absence of water in the extractant. In this case, the cost function of two variables does not have a minimum; the cost value decreases monotonically with an increase in the extractant flow flow rate and a decrease in the number of plates. In this case, it is is necessary necessary to to be be guided guided by by a a reasonable reasonable consumption consumption of the extractant. The situation changes significan significantlytly in the presence of water in thethe extractant.extractant.

ef=0.3 45 ef=0.4 ef=0.6

N(R+1)

16 20 24 28 N T Figure 14. Capital costs for extractive extractive rectification rectification from th thee number of trays in the column at different different flowflow rates of fresh extractant.

NT=28

NT=24

NT=18 44

N(R+1)

0.30.40.50.6 ef Figure 15. Capital costs for extractive rectification from the fresh extractant consumption with a different number of trays in the column.

Figures 16 and 17 show similar calculation results for the case of water content in the extractant of −3 wt %. It can be seen that minima appear on the cost function. The minimum occurs when the water content in the extractant is more than 1%. Its size and position depend on the water content in

Processes 2020, 8, x FOR PEER REVIEW 17 of 22

minimum; the cost value decreases monotonically with an increase in the extractant flow rate and a decrease in the number of plates. In this case, it is necessary to be guided by a reasonable consumption of the extractant. The situation changes significantly in the presence of water in the extractant.

ef=0.3 45 ef=0.4 ef=0.6

N(R+1)

16 20 24 28 N T Processes 2020Figure, 8, 1466 14. Capital costs for extractive rectification from the number of trays in the column at different 17 of 21 flow rates of fresh extractant.

NT=28

NT=24

NT=18 44

N(R+1)

0.30.40.50.6 ef FigureFigure 15. 15.Capital Capital costs costs for for extractive extractive rectiﬁcationrectification fr fromom the the fresh fresh extractant extractant consumption consumption with witha a diﬀerentdifferent number number of trays of trays in thein the column. column.

FiguresFigures 16 and 16 and 17 show 17 show similar similar calculation calculation results results for for thethe casecase of water water content content in in the the extractant extractant of Processes3 wtof %. − 20203 Itwt can, 8%., x beFORIt can seen PEER be that seenREVIEW minima that minima appear appear on the on costthe cost function. function. The The minimum minimum occurs occurs when when the the18water of 22 − water content in the extractant is more than 1%. Its size and position depend on the water content in content in the extractant is more than 1%. Its size and position depend on the water content in the the extractant. For example, for 2 wt % of water in the extractant minimum, capital costs correspond extractant. For example, for 2 wt % of water in the extractant minimum, capital costs correspond to the to the relative consumption of the extractant 0.45 and 20 theoretical trays, and for 3 wt % of water, relative consumption of the extractant 0.45 and 20 theoretical trays, and for 3 wt % of water, relative relative consumption of the extractant 0.4 and 24 theoretical trays. Thus, with an increase in water consumption of the extractant 0.4 and 24 theoretical trays. Thus, with an increase in water content content from 0 to 3 wt %, costs are doubled, and with an increase in the concentration of water in the fromextractant 0 to 3 wtfrom %, costs2 to 3 are wt doubled, %, costs and increase with anby increase 50%. Thus, in the it concentrationis necessary ofto wateranalyze in thethe extractantpossible fromcompensation 2 to 3 wt %, for costs the increaseincrease byin 50%.capital Thus, costs it isfor necessary extractive to analyzerectification the possibleby reducing compensation the cost of for thesolvent increase regeneration. in capital costs for extractive rectiﬁcation by reducing the cost of solvent regeneration.

ef=0.3 ef=0.4 ef=0.5 ef=0.6 80

N(R+1)

20 22 24 26 28 30 N T FigureFigure 16.16. CapitalCapital costs for extractive rectification rectification of of the the trays trays number number in in the the column column at at different different flow flow ratesrates of of thethe extractantextractant withwith aa waterwater contentcontent of 3 wt %.

NT=22

NT=25

NT=30

N(R+1)

0.2 0.3 0.4 0.5 0.6 ef Figure 17. Capital costs for extractive rectification from the flow rate of the extractant with a water content of 3 wt % with a different number of trays in the column.

Figure 18 shows the process flow diagram of ethanol dehydration. The process of obtaining ethanol with a water content of 0.005 wt % takes place in the distillation column C-1, where the extractant AEBA-TEG is fed to the first tray. An aqueous solution of the extractant with alcohol content of 10-5 wt % is removed from the bottom of the column C-1. Further, the extractant is fed to the simple distillation apparatus C-2, where the AEBA is dehydrated. At the outlet of the cube, the

Processes 2020, 8, x FOR PEER REVIEW 18 of 22 the extractant. For example, for 2 wt % of water in the extractant minimum, capital costs correspond to the relative consumption of the extractant 0.45 and 20 theoretical trays, and for 3 wt % of water, relative consumption of the extractant 0.4 and 24 theoretical trays. Thus, with an increase in water content from 0 to 3 wt %, costs are doubled, and with an increase in the concentration of water in the extractant from 2 to 3 wt %, costs increase by 50%. Thus, it is necessary to analyze the possible compensation for the increase in capital costs for extractive rectification by reducing the cost of solvent regeneration.

ef=0.3 ef=0.4 ef=0.5 ef=0.6 80

N(R+1)

20 22 24 26 28 30 N T

ProcessesFigure2020 ,16.8, 1466Capital costs for extractive rectification of the trays number in the column at different flow18 of 21 rates of the extractant with a water content of 3 wt %.

NT=22

NT=25

NT=30

N(R+1)

0.2 0.3 0.4 0.5 0.6 ef ProcessesFigure 2020 ,17. 8, xCapitalCapital FOR PEER costs costs REVIEW for for extractive extractive rectification rectification from from the the flow flow rate of the extractant with a water 19 of 22 content of 3 wt % with a different different number of trays inin thethe column.column. pressure increases and the temperature drops to the level of these indicators on the first tray of columnFigure C-1. 18 Since shows at low the concentrations process flowflow diagramdiagramof water the ofof ethanolethanolAEBA solution dehydration.dehydration. has a high The viscosity, process ofit is obtaining possible ethanolto reduce with it by a mixingwater contentwith the of reflux 0.005 stream wt % fromtakes theplace C-1 in column. the distillation column C-1, where the extractantThe results AEBA-TEGAEBA-TEG of calculating isis fed fed to to the the first mainfirst tray. tray. parameters An aqueousAn aqueous of solution the solution technological of the of extractant the extractantscheme with (Figure alcohol with alcohol content18) for -5 contentofdifferent 10 wt of variants %10 is-5 removedwt %of is the removed from water the content from bottom the in of bottom the the extrac column of thetant C-1. column are Further, presented C-1. the Further, extractantin Table the 2. extractant is The fed values to the is simple fedof the to thedistillationtrays simple number distillation apparatus and the apparatusC-2,extractant where C-2, flow the where AEBA rate were the is dehydrated. AE selectedBA is dehydrated.from At the outletcondition At the of theoutlet of cube,the of minimum the the cube, pressure costthe increasesfunction shown and the in temperature Figures 15–17. drops For to thecomparison level of these, the indicatorssame table on shows the first the tray results of column of energy C-1.

Sinceconsumption at low concentrations per 1 kg of dried of waterethanol the when AEBA using solution glycerol has [34] a high and viscosity, glycols [35] it is as possible an extractant, to reduce and it byfor mixingazeotropic with rectification the reﬂux stream [36]. from the C-1 column.

Figure 18.18. Technology system of ethyl alcohol separation from anan aqueousaqueous solutionsolution processprocess withwith AEBA-TEGAEBA-TEG asas extractant.extractant. The results of calculating the main parameters of the technological scheme (Figure 18) for different Table 2. Results of modeling the process of ethyl alcohol separation from an aqueous solution. variants of the water content in the extractant are presented in Table2. The values of the trays number and the extractant flow rate were selected from the conditionXdE, ofXwE the, minimum cost function shown in Energy Consum., Figures 15Extractant–17. For comparison, ef the N sameт Np table Nе showsR Mass the results Mass of energy P(vac.), consumption Bar per 1 kg of dried mJ/kg ethanol when using glycerol [34] and glycols [35] asFrac. an extractant, Frac. and for azeotropic rectification [36]. AEBA-TEG 0.4 18 8 1 0.81 0.995 10−7 0.1 1.74 (99.99 wt %) AEBA-TEG (98 wt %) 0.45 20 10 1 1.06 0.995 10−7 0.1 1.98 AEBA-TEG (97 wt %) 0.4 24 15 1 1.56 0.995 10−7 0.1 2.56 Glycerol 1 18 10 2 0.35 0.998 0.005 0.1 1.63 (99.99 wt %) [34] Ethylene glycol 0.71 24 12 2 0.5 0.998 0.004 0.1 1.76 (99.99 wt %) [35] Pentane (azeotrop. Dist.) 6.3 [36]

As can be seen from the table, the highest energy consumption corresponds to the process of azeotropic rectification. For the scheme with extractive rectification, the energy consumption per 1 kg of alcohol turns out to be similar for different extractants. In this case, for AEBA, the required amount of extractant per 1 kg of alcohol is 1.5 times less than for other extractants.

4. Conclusions In this work, AEBA were considered as promising extractants for the extraction of alcohols from aqueous solutions using rectification. The conditions of vapor–liquid equilibrium in the AEBA- DEG/AEBA-TEG–aqueous solution of ethanol/isopropanol systems were modeled on the basis of the model for the UNIFAC activity coefficients. For this, a variant of the division of AEBA molecules into

Processes 2020, 8, 1466 19 of 21

Table 2. Results of modeling the process of ethyl alcohol separation from an aqueous solution.

Energy X , X , Extractant ef N Np Ne R dE wE P (vac.), Bar Consum., T Mass Frac. Mass Frac. mJ/kg AEBA-TEG 0.4 18 8 1 0.81 0.995 10 7 0.1 1.74 (99.99 wt %) − 7 AEBA-TEG (98 wt %) 0.45 20 10 1 1.06 0.995 10− 0.1 1.98 7 AEBA-TEG (97 wt %) 0.4 24 15 1 1.56 0.995 10− 0.1 2.56 Glycerol 1 18 10 2 0.35 0.998 0.005 0.1 1.63 (99.99 wt %) [34] Ethylene glycol 0.71 24 12 2 0.5 0.998 0.004 0.1 1.76 (99.99 wt %) [35] Pentane (azeotrop. Dist.) [36] 6.3

As can be seen from the table, the highest energy consumption corresponds to the process of azeotropic rectification. For the scheme with extractive rectification, the energy consumption per 1 kg of alcohol turns out to be similar for different extractants. In this case, for AEBA, the required amount of extractant per 1 kg of alcohol is 1.5 times less than for other extractants.

4. Conclusions In this work, AEBA were considered as promising extractants for the extraction of alcohols from aqueous solutions using rectification. The conditions of vapor–liquid equilibrium in the AEBA-DEG/AEBA-TEG–aqueous solution of ethanol/isopropanol systems were modeled on the basis of the model for the UNIFAC activity coefficients. For this, a variant of the division of AEBA molecules into group components has been proposed, and the previously unknown geometric and energy parameters of the boron group have been identified. The identification of the parameters was carried out according to the known data of vapor–liquid equilibrium in binary solutions of borates, as well as from the boiling points of AEBA solutions in water, ethanol, and isopropanol measured in this work. The adequacy of the calculation of the activity coefficients in the three-component mixture AEBA-DEG/AEBA-TEG–aqueous solution of ethanol/isopropanol was confirmed by comparing the solution of the open evaporation equation with the experimental data of the residue curve. Thus, the obtained matrix of UNIFAC parameters (UNIFAC Matrix) makes it possible to calculate and predict the activity coefficients of the components of aqueous–alcoholic solutions in the presence of AEBA of various molecular structures with satisfactory accuracy. Modeling of the extractive rectification process of an ethanol–water mixture with an extractant AEBA-TEG for the production of ethanol with a purity of 99.5 wt % was conducted. Additionally, the behavior of the cost function on the reflux ratio (the number of theoretical trays) and the relative consumption of the extractant and the water content in it were investigated. It has been shown that the water content in the extract increases the costs of rectification and leads to the appearance of a minimum on the cost function, which must be taken into account when designing industrial technologies. Modeling of the process flow diagram for the production of ethanol with a water content of 0.5 wt % was conducted, and regeneration of the extractant AEBA-TEG was carried out. It is shown that the energy consumption per 1 kg of alcohol is similar for different extractants (AEBA, glycerol, glycols); however, when using AEBA, the amount of extractant per 1 kg of alcohol is required 1.5 times less.

Author Contributions: A.V.K. conceived of the study, coordinated the study, carried out data and results analysis, carried out sequence alignments and drafted the manuscript; A.V.M., designed the study, coordinated and participated in the experimental part of the study and carried out the results analysis; A.R.K. carried out the separation experiment, carried out phase equilibrium experiments, carried out sequences of the study and provided necessary calculations; S.E.D. carried out synthesis lab work and participated in the manuscript drafting; I.M.D. coordinated lab work, carried out sequences, analyzed the obtained results and critically revised the manuscript. All authors have read and agreed to the published version of the manuscript. Processes 2020, 8, 1466 20 of 21

Funding: This research was funded by the Ministry of Science and Higher Education of the Russian Federation grant number 075-00315-20-01 “Energy saving processes of liquid mixtures separation for the recovery of industrial solvents”. Conﬂicts of Interest: There are no conﬂict to declare.

Abbreviations

AEBA Aminoethers of boric acid IL Ionic liquid DEG diethylene glycol TEG triethylene glycol AEBA-DEG aminoethers of boric acid based on diethylene glycol AEBA-TEG aminoethers of boric acid based on triethylene glycol VLE Vapor–liquid equilibrium TMB Trimethyl borate

References

1. State Standard: GOST P 56146-2014 Denatured Ethanol Used as a Fuel Component for Spark igniTion Engines–M; National Standard of the Russian Federation: Moscow, Russia, 2015; p. 27. 2. Stabnikov, V.N. Distillation and Rectification of Ethyl Alcohol; Food industry: Moscow, Russia, 1969; 456p. 3. Furter, W.F. Extractive distillation by salt effect. Chem. Eng. Commun. 1992, 116, 35–40. [CrossRef] 4. Lei, Z.; Zhou, R.; Duan, Z. Process improvement on separating C4 by extractive distillation. Chem. Eng. J. 2002, 85, 379–386. [CrossRef] 5. Rongqi, Z.; Zhanting, D. Extractive Distillation with Salt in Solvent. Sci. Technol. 1999, 4, 1477–1479. 6. Lei, Z.; Wang, H.; Zhou, R.; Duan, Z. Influence of salt added to solvent on extractive distillation. Chem. Eng. J. 2002, 87, 149–156. [CrossRef] 7. Fu, J. Simulation of Salt-Containing Extractive Distillation for the System of Ethanol/Water/ Ethanediol/KAc. 1. Calculation of the VaporLiquid Equilibrium for the Salt-Containing System. Ind. Eng. Chem. Res. 2004, 43, 1274–1278. [CrossRef] 8. Fu, J. Simulation of Salt-Containing Extractive Distillation for the System of Ethanol/Water/Ethanediol/KAc. 2. Simulation of Salt-Containing Extractive Distillation. Ind. Eng. Chem. Res. 2004, 43, 1279–1283. [CrossRef] 9. Tiverios, P.G.; Van Brunt, V. Extractive Distillation Solvent Characterization and Shortcut Design Procedure for Methylcyclohexane Toluene Mixtures. Ind. Eng. Chem. Res. 2000, 39, 1614–1623. [CrossRef] − 10. Seiler, M.; Jork, C.; Kavarnou, A.; Arlt, W.; Hirsch, R. Separation of azeotropic mixtures using hyperbranched polymers or ionic liquids. AICHE J. 2004, 50, 2439–2454. [CrossRef] 11. Jin, Z.; Cong, D.; Chun, L.X. Isobaric Vapor–Liquid Equilibria for Ethanol–Water System Containing Different Ionic Liquids at Atmospheric Pressure. Fluid Phase Equilib. 2006, 242, 147–153. 12. Jiang, X.C.; Wang, J.F.; Li, C.X. Vapor Pressure Measurement for Binary and Ternary Systems Containing Methanol, Ethanol and Ionic Liquid Ethy-3-Ethylimidazolium Diethylphosphate. J. Chem. Thermodyn. 2007, 39, 841–846. [CrossRef] 13. Calvar, N.; González, B.; Macedo, E.A.; Domínguez, Á. Study of the behaviour of the azeotropic mixture ethanol–water with imidazolium-based ionic liquids. Fluid Phase Equilibria 2007, 259, 51–56. [CrossRef] 14. Wang, J.; Li, C.-X.; Wang, Z.-H.; Li, Z.-J.; Jiang, Y.-B. Vapor pressure measurement for water, methanol, ethanol, and their binary mixtures in the presence of an ionic liquid 1-ethyl-3-methylimidazolium dimethylphosphate. Fluid Phase Equilibria 2007, 255, 186–192. [CrossRef] 15. Zhao, J.; Jiang, X.-C.; Li, C.; Wang, Z.-H. Vapor pressure measurement for binary and ternary systems containing a phosphoric ionic liquid. Fluid Phase Equilibria 2006, 247, 190–198. [CrossRef] 16. Ge, Y.; Zhang, L.; Yuan, X.; Geng, W.; Ji, J. Selection of ionic liquids as entrainers for separation of (water + ethanol). J. Chem. Thermodyn. 2008, 40, 1248–1252. [CrossRef] 17. Wang, J.; Li, C.X.; Wang, Z.H. Measurement and Prediction of Vapor Pressure of Binary and Ternary Systems Containing 1-Ethyl-3-methylimidazolium Ethyl Sulfate. J. Chem. Eng. Data 2007, 52, 1307–1312. [CrossRef] Processes 2020, 8, 1466 21 of 21

18. Calvar, N.; Gonzalez, B.; Gomez, E.; Dominguez, A. Vapor-Liquid Equilibria for Ternary System Ethanol + Water + 1-Ethyl-3-Methylimidazolium Ethyl Sulfate and Corresponding Binary Systems Containing the Ionic Liquids at 101.3 kPa. J. Chem. Eng. Data 2008, 53, 820–825. [CrossRef] 19. Davletbaeva, I.M.; Klinov, A.V.; Khairullina, A.R.; Malygin, A.V.; Dulmaev, S.E.; Davletbaeva, A.R.; Mukhametzyanov, T.A. Organoboron Ionic Liquids as Extractants for Distillation Process of Binary Ethanol + Water Mixtures. Processes 2020, 8, 628. [CrossRef] 20. Swietoslawski, W. Azeotropy and Polyazeotropy; Macmillan Company: New York, NY, USA, 1963. 21. Reid, R.C.; Prausnitz, J.M.; Poling, B.E. The Properties of Gases and Liquids 1987, 4th ed.; McGraw-Hill: New York, NY, USA, 1987; p. 753, ISBN 10:0070517991. 22. Wittig, R.; Lohmann, J.; Gmehling, J. Vapor-Liquid Equilibria by UNIFAC Group Contribution. Revision and Extension. Ind. Eng. Chem. Res. 2003, 42, 183–188. [CrossRef] 23. Online UNIFAC Source. Available online: http://unifac.ddbst.de/unifac-matrix.html (accessed on 24 June 2020). 24. Abrams, D.S.; Prausnitz, J.M. Statistical thermodynamics of liquid mixtures: A new expression for the excess Gibbs energy of partly or completely miscible systems. AIChE J. 1975, 21, 116–128. [CrossRef] 25. Have you looked at the Journal of Physical Chemistry recently. Chem. Eng. News Arch. 1964, 42, 441. [CrossRef] 26. Badenhoop, J.K.; Weinhold, F. Natural steric analysis:Ab initiovan der Waals radii of atoms and ions. J. Chem. Phys. 1997, 107, 5422–5432. [CrossRef] 27. Mantina, M.; Chamberlin, A.C.; Valero, R.; Cramer, C.J.; Truhlar, D.G. Consistent van der Waals Radii for the Whole Main Group. J. Phys. Chem. A 2009, 113, 5806–5812. [CrossRef][PubMed] 28. +Online Source: Dortmund Data Bank. Thermophysical Data for Process Design. Available online: http://ddbonline.ddbst.com/DDBSearch/onlineddboverview.exe (accessed on 24 June 2020). 29. Kraher, M.V.; Plank, C.A.; Laukhuf, W.L.S. Vapor-Liquid Equilibria of the System Trimethyl Borate (1)-n-Cyclohexane (2). J. Chem. Eng. Data 1986, 31, 387. [CrossRef] 30. Niswonger, D.S.; Plank, C.A.; Laukhuf, W.L.S. Vapor-liquid equilibria of the system trimethyl borate (1)-n-heptane (2). J. Chem. Eng. Data 1985, 30, 209–211. [CrossRef] 31. Online Source: NIST Standard Reference Database. Available online: https://webbook.nist.gov/chemistry/ (accessed on 24 June 2020). 32. Richardson, J.F.; Harker, J.H.; Backhurst, J.R. Particle Technology and Separation Processes. In Chemical Engineering, 5th ed.; Oxford Butterworth-Heinemann: Oxford, UK, 2003; Volume 2, p. 1231, ISBN 978-0-7506-4445-7. 33. Mai, N.L.; Ahn, K.; Koo, Y.-M. Methods for recovery of ionic liquids—A review. Process Biochem. 2014, 49, 872–881. [CrossRef] 34. Gil, I.D.; Gómez, J.M.; Rodríguez, G. Control of an extractive distillation process to dehydrate ethanol using glycerol as entrainer. Comput. Chem. Eng. 2012, 39, 129–142. [CrossRef] 35. Meirelles, A.; Weiss, S.; Herfurth, H. Ethanol dehydration by extractive distillation. J. Chem. Technol. Biotechnol. 2007, 53, 181–188. [CrossRef] 36. Kiss, A.A.; David, J.; Suszwalak, P.C. Enhanced bioethanol dehydration by extractive and azeotropic distillation in dividing-wall columns. Sep. Purif. Technol. 2012, 86, 70–78. [CrossRef]

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