Study of the Fusel Oil Distillation Process - Industrial & Engineering

Article pubs.acs.org/IECR

Study of the Fusel Oil Distillation Process Marcela C. Ferreira, Antonio J. A. Meirelles, and Eduardo A. C. Batista* ExTrAE, Laboratory of Extraction, Applied Thermodynamics and Equilibrium, Department of Food Engineering, Faculty of Food Engineering, University of Campinas-UNICAMP, 13083-862, Campinas, SP, Brazil ABSTRACT: Fusel oil is a byproduct obtained from bioethanol distilleries, composed of a mixture of higher alcohols such as isoamyl alcohol, isobutyl alcohol, and others. This study aimed to evaluate the industrial distillation process of fusel oil to obtain isoamyl alcohol using the Aspen Plus simulator, considering fusel oil as a mixture of nine components. Fusel oil samples collected in Brazilian industrial mills were analyzed by gas chromatography. An investigation of phase equilibrium (vapor−liquid equilibrium and liquid−liquid equilibrium) was carried out for the components involved in this mixture. Three configurations for the fusel oil separation process were proposed. The best design with minimum total annual cost (TAC) resulted in a recovery of 99.53% of isoamyl alcohol of product containing the isomers isoamyl alcohol (0.818 w/w) and active amyl alcohol (0.178 w/w). Dynamic control of this configuration was investigated, and the results show that reasonable control performance can be achieved.

1. INTRODUCTION The production of bioethanol generates byproducts, such as fusel oil, which consists of a mixture of higher alcohols. In countries where large quantities of ethanol are produced, like Brazil, alternatives for the use of these byproducts are of great importance to make ethanol production less polluting and more profitable. In Brazil, fusel oil is generally produced in the proportion of 2.5 L per 1000 L of ethanol. Brazilian ethanol production (anhydrous and hydrated) in 2011 exceeded 27 billion liters. At this scale a total of 67.5 million liters of fusel oil can be generated per year. According to Patil et al.,1 the term fusel oil is used to designate a mixture of higher alcohols obtained during distillation of bioethanol. Higher alcohols are those consisting of more than two carbon atoms such as isoamyl alcohol, isobutanol, propanol, butanol, and others. Values of the relative volatilities of higher alcohols as a function of ethanol concentration in the liquid phase were investigated by Batista and Meirelles2 via simulation of a liquid−vapor flash. The higher alcohols exhibit a decrease in volatility as the ethanol concentration in the liquid phase increases, acting as light components in the diluted ethanol range and as heavy components in the concentrated ethanol range. Because of this behavior they should be classified as components with intermediate volatility. In a bioethanol rectifying column, the region near its bottom presents low ethanol contents; therefore, the volatility of higher alcohols is increased and they concentrate in the vapor phase, rising up the column. In regions near the top of this column, the ethanol content increases; the volatility of higher alcohols therefore decreases and they tend to concentrate in the liquid phase, going back down the column. Thus, the higher alcohols form a cycle inside the column, reaching their maximum concentration in regions near the bottom of the rectifying column, from where they must be withdrawn so not to interfere in operation of the column. Distilleries in Brazil do not use fusel oil but instead sell it to other industries. The distilleries efficiently remove ethanol © 2013 American Chemical Society

present in the fusel oil sidestream in order to maximize its economic value. The higher alcohol stream removed from the bioethanol rectifying column is collected and goes to a decanter, where the fusel oil is washed in countercurrent with water. In this process two phases are formed. At the top of the decanter fusel oil is collected, and at the bottom there is a mixture rich in water and ethanol, which is returned to the columns of the bioethanol process. As mentioned before, fusel oil is a mixture of several alcohols, which limits its direct use as a solvent. The higher alcohols present in fusel oil are considered natural products, which gives them high commercial values.3 Isoamyl alcohol, the main component of fusel oil, can be used in the production of organic esters, which are used as industrial solvents, flavoring agents, and plasticizers. Due to the presence of water in fusel oil, heterogeneous azeotropes form between water and most of the higher alcohols. The separation of a heterogeneous azeotrope mixture is much easier than that of a homogeneous azeotrope mixture, because the liquid−liquid equilibrium that occurs in the decanter can be used to facilitate this separation.4 Isoamyl alcohol and water form a heterogeneous azeotrope at atmospheric pressure and 95.1 °C, with a composition of 50.4% w/w of isoamyl alcohol. Fusel oil cannot be discarded directly into the environment, since it would cause undesirable environmental impacts. Some authors have indicated direct applications of fusel oil: it can be burned to supply energy in the distilleries5 or it can be added to diesel fuel to improve the cetane index.1 Considering the amount of fusel oil in Brazil, a plant capable of producing isoamyl alcohol from fusel oil obtained from different distilleries is an interesting industrial application, since this alcohol is highly valued; its market price is approximately 3 times the price of ethanol fuel. Received: Revised: Accepted: Published: 2336

March 12, 2012 December 18, 2012 January 10, 2013 January 10, 2013 dx.doi.org/10.1021/ie300665z | Ind. Eng. Chem. Res. 2013, 52, 2336−2351

Industrial & Engineering Chemistry Research

Article

There are many studies in literature that report the application of fusel oil for obtaining esters by enzymatic or biotechnological processes,5−8 but most studies use the pure isoamyl alcohol in their experiments without considering the fusel oil purification process. Simulation studies (both static and dynamic) for fusel oil purification processes were not studied. The present work investigated fusel oil purification using computational simulation. For this study, nine compounds were selected to represent the fusel oil: water, ethanol, 3-methyl-1butanol (isoamyl alcohol), 2-methyl-1-butanol (active amyl alcohol), isobutanol, pentanol, butanol, propanol, and methanol.

Table 1. Composition of Fusel Oil mass fraction components isoamyl alcohol active amyl alcohol isobutanol butanol pentanol propanol ethanol methanol water total

2. CHROMATOGRAPHIC ANALYSIS Fusel oil samples from two bioethanol production plants in the state of São Paulo, Brazil, were analyzed for detailed characterization of their compositions. The samples were coded as samples 1 and 2. For analysis of the alcohols, a gas chromatograph (Shimadzu GC 17 A) with automatic injector (Shimadzu AOC 20i), flame ionization detector (FID), and an Agillent DB-624 column (6% phenylcyanopropyl and 94% dimethylpolysiloxane, measuring 60 m in length, 0.25 mm in internal diameter, and 1.4 μm in film thickness) was used. Analyses were conducted using methanol as the solvent. Operational conditions of the chromatograph were as follows. The injection and detection temperatures were fixed at 230 °C, and column pressure was 178 kPa. The temperature program used for the column was held at 40 °C for 4 min, 1 °C/min ramp to 60 °C, hold at 60 °C for 1 min, ramp from 60 to 83 °C at a rate of 5 °C/min, from 83 to 88 °C at a rate of 0.5 °C/min, and from 88 to 210 °C at a rate of 15 °C/min. The sample injection volume was 1.0 μL with split ratio of 1:30. Hydrogen was used as ignition gas and helium as carrier gas, with a flow of 1.50029 mL/min and linear velocity of 26.2437 cm/s. The total run time was 50 min. For analysis of methanol, water was used as the solvent, utilizing the following temperature program: hold at 40 °C for 4 min, 1 °C/min ramp to 60 °C, hold at 60 °C for 1 min, then a ramp 7 °C/min to 210 °C, and hold at 210 °C for 1 min. The injection volume was 1.5 μL and the other conditions were identical to the previous procedure. The components were quantified by the external standard technique based on the construction of calibration curves for the eight components, except water. All standard components were chromatographic grade and purchased from Sigma Aldrich. Quantification of water was performed in triplicate using a Karl Fischer titrator (model Metrohm 710 KF Tritino). Table 1 summarizes the results of fusel oil samples from two bioethanol plants, analyzed by GC and Karl Fischer. It can be observed that isoamyl alcohol, active amyl alcohol, water, isobutanol, and ethanol are the major compounds. The presence of heavier compounds, which were not quantified, was observed in all chromatograms. These components are possibly esters and alcohols with more than six carbons.9 However, because they were found in minimal concentrations, they were not taken into account in this study. In order to obtain a single fusel oil composition to be used in the simulations, compositions of samples 1 and 2 were averaged, since they are very similar. A normalization of this composition was carried out, resulting in the normalized fusel oil composition also shown in Table 1.

average (samples sample 1 sample 2 1 and 2)

normalized composition

0.5305 0.1477

0.5438 0.0850

0.5372 0.1164

0.5570 0.1207

0.0561 0.0074 0.0002 0.0032 0.0681 0.0004 0.1478 0.9613

0.0733 0.0036 0.0005 0.0149 0.0989 0.0000 0.1475 0.9674

0.0647 0.0055 0.0003 0.0090 0.0835 0.0002 0.1476 0.9644

0.0671 0.0057 0.0003 0.0094 0.0866 0.0002 0.1531 1.0000

3. EVALUATION AND THERMODYNAMIC MODELING OF PHASE EQUILIBRIUM For accurate design and optimization of the fusel oil distillation processes, reliable knowledge of the phase equilibrium behavior is required. In this work, the description of phase equilibrium [vapor−liquid equilibrium (VLE) or liquid−liquid equilibrium (LLE)] of the components found in fusel oil was investigated using the parameters found in the Aspen Plus databank (commercial simulator from Aspen Technology). The phase equilibrium was calculated using the parameters for each binary subsystem of the simulator and compared to the experimental data found in the literature. 3.1. Vapor−Liquid Equilibrium (VLE). Vapor−liquid equilibrium data for 36 binary parameters were investigated using experimental data collected from several published studies.10−27 There are no experimental measurements of vapor−liquid equilibrium using fusel oil, and even equilibrium data for the binary systems considered in this work are quite scarce. The aim of this part was just to check the parameters, since the efficiency of the simulations is related to the quality of these parameters in describing VLE. The NRTL model was chosen for calculating the activity coefficient, and the fugacity coefficient of the vapor phase was calculated by the Virial equation in association with the Hayden−O’Connell equation. The NRTL and Hayden−O’Connell models were used by Bessa et al.28 to reproduce experimental process data from Brazilian bioethanol distilleries, leading to lower deviations. Due to the similarity of the systems, these models were chosen in this work. The equilibrium was calculated using the NRTL interaction parameters available in the Aspen Plus databank and compared to the experimental data. Also it was calculated the average absolute deviation of the vapor phase molar fraction. The criterion defined for accepting the parameters present in the database of the simulator was that the average deviations of the experimental and calculated mole fraction must be less than 0.03. On the other hand, if the deviation was higher than this criterion, the NRTL parameters for this binary were readjusted. For the binaries that had no parameters in Aspen Plus, thermodynamic modeling of the experimental data was performed with the commercial simulator. Isoamyl alcohol + water13 and pentanol + water14 presented absolute deviations in vapor phase composition greater than 0.03. Consequently, regression of the Aspen Plus parameters from experimental data was carried out for these two binaries. Before the readjustment, the average deviations in the vapor phase were 0.037 for isoamyl alcohol + water and 0.039 for 2337

dx.doi.org/10.1021/ie300665z | Ind. Eng. Chem. Res. 2013, 52, 2336−2351


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Figure 1. T−x−y diagram for the water (1) + isoamyl alcohol (2) system, at 1 atm (experimental data from Cho et al.13 and calculated results are shown).

columns were simulated at atmospheric pressure and using a Murphree efficiency equal to 0.7. In the simulations conducted the following responses were evaluated: mass fraction of isoamyl alcohol in final product (IA), the percentage of isoamyl alcohol recovery (RE), and specific steam consumption (SSC). The variable IA is obtained directly in the simulations, while the variable recovery (RE) was obtained by eq 1.

pentanol + water, and after readjustment they were reduced to 0.007 and 0.016, respectively. This new parameter databank enables more accurate computational simulations of fusel oil in distillation columns. Figure 1 shows the T−x−y diagram for the isoamyl alcohol + water13 system at 1 atm and the VLE results obtained by the NRTL parameters, before and after the readjustment, where the presence of a heterogeneous azeotrope can be verified. 3.2. Liquid−Liquid Equilibrium (LLE). From the nine components considered for fusel oil in this study, it was found that five binaries present restricted miscibility at room temperature: water + 3-methyl-1-butanol, water + 2-methyl-1-butanol, water + isobutanol, water + butanol, and water + pentanol. Therefore, it was necessary to search for ternary LLE data whose components were among those studied in the present work, and 12 data sets were found in the literature.29−32 Several tests were performed to select the thermodynamic model that best represents the experimental data, also evaluating the libraries of the parameters present in the simulator. For the thermodynamic parameters of the LLE, the phase composition of the system was better described by the UNIQUAC model (using the LLE-LIT parameters of the Aspen databank) when compared to the NRTL model. Because it is possible to use different parameter banks in the same simulation in Aspen Plus, it was decided to use the UNIQUAC model to describe the LLE in the decanters of the following simulations.

RE (%)

⎛ isoamyl alcohol mass flow in the bottom stream ⎞ = 100⎜ ⎟ ⎝ isoamyl alcohol mass flow in the feed stream ⎠ (1)

The energy requirements of the reboiler are industrially expressed through the specific steam consumption, which is the mass of steam consumed in the heat exchanger per kg of final product produced. However, the simulator presents the energy consumed in the reboiler in kcal/h. Hence, SSC is calculated using eq 2, where the heat of vaporization (λ) assumes the value of 503.8 kcal/kg33 (considering saturated steam at a pressure of 5 kgf/cm2). ⎛ kg steam ⎞ SSC⎜ ⎟ ⎝ kg product ⎠ =

energy (kcal/h) λ (kcal/kg vapor) × product mass flow (kg product/h) (2)

4.1. Conceptual Design of the Ternary System: Water + Ethanol + Isoamyl Alcohol. Distillation of the ternary system water + ethanol + isoamyl alcohol was studied to better understand the process of heterogeneous azeotropic distillation of fusel oil, considering the following fusel oil composition (in mass fraction): 0.16 of water, 0.24 of ethanol, and 0.60 of isoamyl alcohol. A column with 10 trays was modeled with feed at its middle, plus two stages, where the first stage corresponds to the condenser and the last one represents the reboiler, totalizing 12 stages. The reflux ratio was set at 1.5 and the bottom mass flow was defined as the maximum mass flow which allows for

4. COMPUTATIONAL SIMULATIONS OF FUSEL OIL DISTILLATION PROCESS Simulations were conducted using the commercial simulator Aspen Plus from Aspen Technology. In all simulations, the distillation column was represented using the RadFrac model, where calculations are performed based on the MESH equations. Numbering of the stages in the simulations starts at the top of the column, where the condenser is represented by the first stage and the reboiler as the last stage. The internal algorithm used by the simulator to solve the equations was Newton. All distillation 2338



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Figure 2. Ternary diagram for water + ethanol + isoamyl alcohol at 1 atm and mass balance lines of the one column design.

obtaining a bottom product with isoamyl alcohol purity greater than 99%. The vapor−liquid−liquid equilibrium (VLLE) diagram for this ternary system can be observed in Figure 2. A distillation boundary divides the system into two distillation regions. The location of the feed composition and the results for the distillate and bottom compositions for simulation are also marked in Figure 2, as well as the liquid composition profile of the column. From the analysis of this figure, it is observed that it is a possible distillation process, since the points of composition of distillate and bottom products are close to a residual curve and these points are aligned with the point of feed composition. Using a bottom product mass flow rate of 51.77 kg/h, a purity of 99% isoamyl alcohol in this stream was acquired. It was noted that trays 2−10 exhibited separation of liquid phases in the simulated process. This is in accordance with the diagram shown in Figure 2, since the points of the liquid composition of these trays are located in the VLLE region. It was observed that the distillate of this column still has a high concentration of isoamyl alcohol (0.18 in mass fraction). The point of the distillate composition is located near the distillation boundary in the diagram. This boundary cannot be crossed with a single column distillation; consequently, in this case it is not possible to obtain pure components. But a decanter can be used to cross the distillation boundary and consequently obtain purer products with greater recovery. This observation allowed for developing the configuration of the ternary system as shown in the flow diagram in Figure 3, where a liquid sidestream (sidestream) is removed from the column and feeds the decanter. The organic phase (organic) rich in higher alcohols returns to the column, while the aqueous phase (aqueous) is removed from the process. The organic stream returns to the distillation column at one stage below the stage where the sidestream is removed, in order to maintain the liquid mass flow in the column. Still analyzing the ternary system, a simulation was performed with the column operating under the same conditions as in the

Figure 3. Configuration for the ternary system.

previous column simulation, but with a bottom mass flow of 60 kg/h (total input mass flow of isoamyl alcohol in the process), and the minimum reflux ratio required was determined to generate a bottom product with a purity of at least 0.99 in isoamyl alcohol mass fraction. The reflux ratio acquired was 18. The sidestream was withdrawn from the column at a mass flow of 43 kg/h at stage 9 (corresponding to tray 8) and the organic phase returned to the column in stage 10. Operating conditions of the decanter were atmospheric pressure and temperature of 86.3 °C, the same temperature of the stage where the sidestream was withdrawn. The results for the compositions of the main streams are shown in Figure 4. With this configuration a mass fraction of 0.99 isoamyl alcohol was obtained along with 99% recovery of this alcohol. It was observed that inclusion of the 2339



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Figure 4. Ternary diagram for water + ethanol + isoamyl alcohol at 1 atm and mass balance lines for configuration of Figure 3.

Figure 5. Liquid composition profile A: one column, mass flow bottom of 60 kg/h.

decanter led to a decrease in the number of trays presenting the formation of two liquid phases, since part of the water was removed from the process via the aqueous stream. 4.2. One-Column Process: Multicomponent System. The product of interest in this study is the isoamyl alcohol, and this is one of the heaviest components present in fusel oil. Thus, in the distillation process this component is obtained at high purity in the bottom stream. In order to evaluate the impact of the mass flow of the bottom stream in the isoamyl alcohol purification process, preliminary simulations were carried out using a configuration with only one distillation column with the following specifications: feed mass flow of 100 kg/h, column with 30 stages, feed at stage 15 (normalized composition shown in Table 1), and reflux ratio of 5.75. In these simulations, the bottom product mass flow varied from 40 to 70 kg/h. It was possible to note two main composition profiles of the liquid phase in the column, referred to as profile A (Figure 5) and profile B (Figure 6). Profile A is found when the bottom mass flow is equal to or greater than 60 kg/h. In this case, two liquid phases emerge inside

the column, with water exiting in the bottom stream, making it impossible to obtain isoamyl alcohol with high purity. It should be noted that the composition shown in Figure 5 refers to the overall composition of the liquid phase. A configuration was proposed (named configuration A) in which this column operates with increased bottom stream mass flow rate. This configuration will be studied later. Profile B is obtained in situations where the bottom mass flow is less than 60 kg/h. In these cases, it is possible to obtain a stream of high-purity isoamyl alcohol, and liquid phase separation is not observed in the column. Nevertheless, the low mass flow results in low recoveries, and therefore, it is necessary to include other unit operations in the process to increase recovery. The configuration in which this column operates with a bottom mass flow lower than 60 kg/h was defined as configuration B and it will be discussed later. Water is a limiting factor in this distillation process because it generates heterogeneous azeotropes with various compounds present in the fusel oil, which then leads to the emerging of two liquid phases. Hence, the reduction of water content in the feed 2340



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Figure 6. Liquid composition profile B: one column, mass flow bottom of 40 kg/h.

stream facilitates fractionation of the solution, this observation aided in the development of a new configuration, called configuration C, described later. It was observed in all cases that it is not possible to obtain only isoamyl alcohol (3-methyl-1-butanol) in the bottom stream due to the presence of active amyl alcohol (2-methyl-1-butanol). These alcohols are isomers and have relative volatilities close to unity, which complicates the separation process. Although the product of interest in this study is isoamyl alcohol, results were expressed as a sum of the mass fraction of isoamyl alcohol and active amyl alcohol. This is a reasonable approximation, since commercial isoamyl alcohol consists of a mixture of these alcohols, verified by GC analysis. The small difference in the boiling point (only 2.5 °C) between the isoamyl alcohol and the active amyl alcohol difficults their separation. Therefore, a large number of trays is needed for distillation separation of these substances as well as a high reflux ratio, which makes this process impracticable. An alternative to a tray distillation process is the use of a packed column, which, due to the greater interfacial area, increases considerably the mass transfer, reducing the size of the distillation column. Liu et al.34 developed a structured packing for a distillation column in order to separate these two isomers, obtaining active an amyl alcohol yield of 80% with 99% purity. For the proposed configurations, it was established that the sum of the compositions of isoamyl alcohol and active amyl alcohol in the product stream should be equal to 0.996. This was maintained so as to compare the different configurations. For the three proposed configurations, a main column with 30 trays was simulated. In preliminary simulations, this number of trays proved to be reasonable to promote the desired separation. Two additional stages are included in each configuration, referring to the reboiler and the condenser. The normalized composition shown in Table 1 was used with a mass flow of 100 kg/h and fed in the middle of the column. The final product mass flow was fixed at 67.77 kg/h. This value corresponds to the sum of the isoamyl alcohol (55.7 kg/h) and active amyl alcohol mass flows (12.07 kg/h) in the feed stream, ensuring higher recoveries. 4.3. Configuration A. In one-column process simulations with bottom mass flow equal to or greater than 60 kg/h (profile A), the emergence of two liquid phases in the distillation column was noticed, which significantly affects the column performance.

This phenomenon is associated with the heterogeneous azeotropic point between water and isoamyl alcohol. In the ternary system, inclusion of the decanter led to a decrease in the number of trays in which there is the formation of two liquid phases in the column distillation. Thus, the same configuration was used for the multicomponent system. It was necessary to include a heat exchanger in the previously proposed configuration to cool the stream directed to the decanter, since this operates at 25 °C. A schematic for this proposed configuration, defined as configuration A, is shown in Figure 7. A distillation column with 30 trays and reflux ratio of 5.75 was used. In order to analyze the influence of sidestream mass flow on the process, simulations were carried out by varying this variable from 10 to 100 kg/h. Withdrawal of the sidestream was evaluated at three stages, 5, 15, and 25, corresponding to trays 4, 14, and 24, and the organic stream always returned to the distillation column one tray below the withdrawal tray. It was found that an increased mass flow of the withdrawn sidestream resulted in increased mass flow of water in the aqueous stream of decanter and, consequently, an increase in the purity of isoamyl alcohol in the bottom stream of the distillation column. However, it can be noted that when exceeding a determined value of the sidestream mass flow (approximately 50 kg/h), the mass flow of water in the aqueous stream and the purity of isoamyl alcohol in the bottom stream do not exhibit significant variation. There is a tendency for water content in the aqueous phase to remain constant at a value between 8 and 9.3 kg/h and the mass fraction of isoamyl alcohol in the bottom stream was maintained between 0.81 and 0.82. Considering the example where the withdrawn sidestream was located at stage 5 and comparing the results for sidestream mass flow rates of 60 and 100 kg/h, it was expected that with an increase of 40 kg/h in mass flow of the sidestream it would be possible to remove more water from the process via the aqueous stream from the decanter. However, the mass flow of water in this sidestream increased only from 9.23 to 9.28 kg/h, and the purity of the isoamyl bottom stream increased from 0.8191 to 0.8194. The aqueous and organic streams of the decanter are determined by mass balance using the conditions of LLE at 25 °C. The restricted solubility of water in higher alcohol rich phase can be emphasized in the LLE. Nevertheless, this solubility is sufficient to carry some water to the organic stream. As the increase in mass flow of sidestream also reflects the increased mass flow of isoamyl 2341



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Figure 7. Flowsheet of the configuration A.

rate in one side of the DWC is also reduced; therefore in these preliminary simulations, it was not possible to remove a sufficient amount of water with the sidestream. The water is a limiting factor in this distillation process because it generates heterogeneous azeotropes with various compounds present in the fusel oil. Consequently, the isoamyl alcohol was not obtained with high purity in the bottom stream of this design. Further investigations must be carried out, and the results from this paper provide the basis for future developments of the use of DWC for the fusel oil distillation. 4.4. Configuration B. As mentioned previously, it is possible to achieve a high purity of isoamyl alcohol using a low flow rate of the bottom product in a process that consists of only one distillation column, as seen in profile B of Figure 6. This system presents low efficiency due to the great loss of higher alcohols in the distillate of the column. A flowsheet was proposed including a second distillation column in order to recover the remaining isoamyl alcohol lost in the distillate of the first column. Preliminary tests were carried out and it was found that the top stream of the first distillation column showed phase separation at 25 °C. Therefore, it was decided to include a decanter to promote partial removal of water from this stream, facilitating the second distillation column process. A schematic for this proposed configuration, defined as configuration B, is shown in Figure 8. In this configuration, fusel oil is fed into column 1. The bottom stream (bottom 1) is isoamyl alcohol-rich, and the distillate (dist 1) goes to a decanter at 25 °C, where there is the formation of an organic phase (organic) and an aqueous solution (aqueous). The second column (column 2) is fed with the organic stream (organic), where this column has the function of recovering the remaining isoamyl alcohol in the bottom stream (bottom 2). The final product (bottom) is the sum of the bottom streams from each column (bottom 1 + bottom 2). Initially a simulation was performed using only the first column with a reflux ratio of 5.75 and bottom mass flow of 59 kg/h. By analyzing the results of this simulation, it can be observed that the bottom stream is a high-purity isoamyl alcohol product (0.822 in mass fraction). Furthermore, the organic stream

alcohol in this stream, then greater transfer of water to the organic stream occurs in the decanter, therefore limiting the amount of water that can be removed from the process via the aqueous stream from the decanter. For the withdrawn sidestream at stage 5, sidestream mass flow rates greater than 60 kg/h are sufficient to eliminate the effect of formation of two liquid phases in the column, permitting purification of the isoamyl alcohol. On the other hand, when the sidestream is removed at stages 15 and 25, the tested flow rates were not sufficient to eliminate formation of two liquid phases in the column, leading to a lower purity of isoamyl alcohol in the bottom stream. Due to the best results, it was defined that the sidestream be withdrawn at stage 5 in configuration A, with a mass flow of 60 kg/h. In configuration A the decanter operated at 25 °C; therefore, it is necessary to cool the sidestream to that temperature. Water at 20 °C was also used as a cooling fluid. The mass flow input of the cooling fluid was 120 kg/h, double that of the sidestream flow rate. In the simulator, the output temperature of the sidestream leaving the heat exchanger (sidestream* in Figure 7) was specified at 25 °C. In this configuration, the reflux ratio was adjusted in order to obtain the desired composition of isoamyl alcohol and active amyl alcohol in the product, reaching a value of 3.3, with SSC equal to 1.1321 kg steam/kg product. Since the configuration A has a sidestream distillation column, the dividing wall columns (DWC) could be used to achieve high purity of the sidestream. Nowadays, several studies have shown that DWC represent a very promising technology, allowing a significant energy requirement reduction as well as savings in the operating costs.35−37 Dividing wall columns technology has already been used before for homogeneous and heterogeneous azeotropic distillations, presenting good results.38−40 Then the DWC can be an attractive alternative for the distillation of fusel oil. However, preliminary results suggest that it is not possible to eliminate enough water from the distillation column. The DWC liquid flow rate is lower than in a traditional distillation column, which contributes to smaller equipment and consequently provides a capital saving.41 As this flow is lower, the liquid flow 2342



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Figure 8. Flowsheet of configuration B.

Figure 9. Flowsheet of the configuration C.

obtained from the top still had a high concentration of higher alcohols including isoamyl alcohol (0.261 in mass fraction), active amyl alcohol (0.058), isobutanol (0.228), and butanol (0.021), as well as water (0.123). The purification of this stream

in column 2 is more complicated, since the relative volatility of the other higher alcohols and isoamyl alcohol are close to unity, obtaining a bottom product with a higher content of other higher alcohols such as isobutanol. From this observation it can be 2343



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(0.996). In these simulations, both distillations columns operated with reflux ratio of 5.75 and feed at the middle of the column. It was concluded that with 16 trays the desired content in the bottom product (bottom) is achieved. The energy consumed by the reboiler was 3.68 kg steam/kg product. It is observed that not only does this configuration present higher energy requirements compared to configuration A, configuration B also includes a second distillation column with a considerable number of trays, which makes the implementation of this proposed configuration impracticable. 4.5. Configuration C. As discussed earlier, the reduction of water content in the feed stream facilitates the fractionation of fusel oil. Therefore, it is necessary to investigate a process capable of substantially removing water from the fusel oil feed stream. The stream of higher alcohols removed from the bioethanol rectification column is submitted to a washing process for the recovery of ethanol which it contains, where the aqueous stream returns to the process and the organic stream is commercialized as fusel oil. Therefore, fusel oil obtained from the distilleries is found at the solubility limit, and the addition of a small amount of water is sufficient to promote phase separation. The addition of water to the fusel oil induces the formation of two liquid phases, an aqueous phase composed mainly of water with low concentration of organics, and an organic phase rich in isoamyl alcohol with a reduced water content. In order to avoid that the addition of water to the process results in the loss of higher alcohol in the aqueous stream, an operation is necessary for their recovery. Configuration C is illustrated in Figure 9. A decanter was used at 25 °C receiving 100 kg fusel oil/h and 1 kg water/h. The organic stream is fed to a distillation column. Stripping of organics (stripping column) was suggested to recover the alcohols present in the aqueous stream. The aqueous stream from the decanter is fed into this column at the first stage. The higher alcohols are at infinite dilution in this solution in which repulsive attractions prevail, resulting in a considerable increase in volatility of these compounds. Thus, the higher alcohols are concentrated in the vapor phase, which allows for obtaining a stream rich in higher alcohols in the top of this column and a stream practically free of organics in its bottom. The conditions used for the stripping column were three stages and a reflux ratio of 1. Mass flow of the distillate was varied to obtain a high recovery of isoamyl alcohol in this stream, assuming a value of 0.5 kg/h in this simulation, which resulted in 99.26% recovery of isoamyl alcohol present in the aqueous stream. Constructive and operating conditions used for the distillation column were identical to those obtained in configuration A,

Table 2. Results Obtained for the Streams of the Decanter Configuration C composition in mass fraction components

feed

isoamyl alcohol active amyl alcohol isobutanol butanol pentanol propanol ethanol methanol water

0.5570 0.1207 0.0671 0.0057 0.0003 0.0094 0.0866 0.0002 0.1530

mass flow (kg/h)

100

aqueous phase

organic phase

0.0123 5.30 × 10−05 0.0049 9.64 × 10−06 6.80 × 10−06 0.0010 0.0048 0.0008 0.9762

0.6102 0.1325 0.0731 0.0063 0.0003 0.0102 0.0946 0.0002 0.0726

9.92

91.08

Table 3. Comparison of the Three Configurations config A

config B

config C

number of columns number of trays

1 30

number of decanters mass fraction of isoamyl alcohol + active amyl alcohol in final product mass fraction of isoamyl alcohol in final product (IA) mass fraction of active amyl alcohol in final product recovery of isoamyl alcohol (RE) (%) specific steam consumption (SSC) (kg steam/kg product)

1 0.996

2 2 30 (column 1) 30 (column) 16 (column 2) 3 (stripping column) 1 1 0.996 0.996

0.818

0.818

0.818

0.178

0.178

0.178

99.53 1.1321

99.53 3.6800

99.53 0.9111

concluded that the mass flow of the bottom product in column 1 must be greater than the mass flow of the bottom product in column 2. Thus, it is possible to obtain a final product (bottom 1 + bottom 2) with high purity of isoamyl alcohol. The bottom product mass flow acquired in column 1 (59 kg/h) was the maximum possible while ensuring that liquid phase separation does not occur inside the column. Therefore, the bottom mass flow of the second column was set at 8.77 kg/h. The sum of these two flow rates represents the total mass flow of isoamyl alcohol and amyl active alcohol fed into the process (67.77 kg/h). A series of steady-state simulations was performed to determine the minimum number of trays in the second distillation column (column 2) needed to attain the content of isoamyl alcohol and active alcohol stipulated for the bottom product

Table 4. Composition of Input and Output Streams of All Configurations config A

isoamyl alcohol active amyl alcohol isobutanol butanol pentanol propanol ethanol methanol water

config B

config C

feed

distillate

bottom

aqueous

dist 2

bottom

aqueous

dist

bottom

aqueous 2

0.5570 0.1207 0.0671 0.0057 0.0003 0.0094 0.0866 0.0002 0.1530

0.011 0.067 0.276 0.015 1.1 × 10−03 0.040 0.369 7.2 × 10−04 0.289

0.818 0.178 1.9 × 10−04 0.003 4.4 × 10−04 0.001 2.8 × 10−05 1.9 × 10−12 8.6 × 10−13

0.003 1.3 × 10−04 0.081 0.004 2.6 × 10−04 9.7 × 10−03 0.082 4.4 × 10−04 0.819

0.006 1.2 × 10−04 0.335 0.019 6.2 × 10−05 0.046 0.412 3.7 × 10−04 0.182

0.818 0.178 6.1 × 10−04 0.003 4.4 × 10−04 0.018 0.002 3.1 × 10−08 2.7 × 10−08

0.012 0.005 0.040 4.1 × 10−04 9.5 × 10−04 0.007 0.081 9.5 × 10−04 0.858

0.012 0.223 0.281 0.014 1.6 × 10−07 0.039 0.363 6.3 × 10−04 0.290

0.818 0.178 2.1 × 10−04 3.5 × 10−03 4.4 × 10−04 1.2 × 10−03 3.8 × 10−05 1.6 × 10−12 4.0 × 10−13

0.096 0.023 2.5 × 10−04 5.9 × 10−03 1.4 × 10−03 0.865 1.4 × 10−03 5.2 × 10−04 0.998

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of isoamyl alcohol occurs, whose mass fraction was increased from 0.5570 to 0.6102. Using configuration C the established purity was obtained with 99.53% of recovery of isoamyl alcohol and 0.9111 of SSC. The low value found for the specific steam consumption is highlighted. This occurs mainly due to the fact that column operates with a lower total liquid mass flow when compared to other configurations, resultant of the preremoval of the aqueous stream in the decanter. Therefore, a lower evaporation rate in the reboiler is required, reducing the energetic cost. By analyzing the results, it can be concluded that the stripping column plays a fundamental role in the recovery of isoamyl alcohol, since a small number of trays was sufficient to recover the higher alcohols in the top of the column. 4.6. Economic Comparison between Configurations A, B, and C. Table 3 presents the equipments involved in each proposal along with the main results, Table 4 shows the composition results, and Figure 10 provides the operating conditions for all configurations. It is observed that in all configurations the active amyl alcohol composition in the final product is similar, which ensures the same proportion of the isomers isoamyl alcohol and active amyl alcohol. Active amyl alcohol is recovered mostly in the bottom stream, with recovery of 99.99% in the three proposed configurations. Configuration C was the most energetically efficient configuration. In relation to the amount of equipment, this proposal also seems to be the most viable. In order to evaluate the cost of these three configurations, an economic analysis was also performed as part of this study. The same operating conditions (input, outputs, cooling water temperature, and steam conditions) were used in all configurations. Total annual cost (TAC) was calculated in order to perform a comparison of the three configurations, where the TAC includes the annualized capital costs and the operating costs. Estimation of equipment costs are based on the procedure of Douglas,42 including column shells, internal trays of the column, and also the reboiler and condenser. Details for the economic basis used in the calculations are based on the study by Luyben43 and are given in Table 5. The payback period was Table 5. Basis of Economics and Equipment Sizing column diameter: Aspen tray sizing length: number of trays with 0.61 m spacing plus 20% extra length condensers heat-transfer coefficient (Uc) = 0.852 kW·K−1·m−2 differential temperature (ΔTc) = reflux-drum temperature − 310 K condenser area (m2) = condenser duty/(UcΔTc) reboilers heat-transfer coefficient (Ur) = 0.568 kW·K−1·m−2 differential temperature (ΔTr) = 424.25 − base temperature reboiler area (m2) = reboiler duty/(UrΔTr) utilities cost LP steam $13.28/GJ cooling water $0.354/GJ TAC = (capital cost/payback period) + operating cost payback period = 3 years Marshall and Swift index45 = 1448.3

Figure 10. Operating conditions for configurations A, B, and C.

except for the reflux ratio, which was 2.6 (the minimum value necessary to achieve the mass fraction of 0.996 of isoamyl alcohol and active amyl alcohol in the product stream). The stream of recovered organics was fed in the middle of the main column (stage 16). Table 2 shows the result for the input and output streams of the decanter. Starting with a feed containing 0.1530 of water in mass fraction and utilizing the decantation process proposed, it was possible to obtain a column feed stream with a water composition of 0.0726. It was also observed that a preconcentration

assumed to be 3 years. Additionally, unit costs of low-pressure steam and cooling water were assumed as 13.28 and 0.354 $/GJ, respectfully, according to Turton et al.44 The TAC for the three configurations are compared in Table 6. 2345



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Table 6. TAC Comparison of Configurations config A column shell cost ($) column trays cost ($) condenser cost ($) reboiler cost ($) steam cost ($/year) cooling water cost ($/year) total annualized capital cost ($/year) total annualized operating cost ($/year) TAC ($1000/year)

config B

config C

col 1

col 1

col 2

col

stripping col

60865.06 1632.65 11815.27 43576.24 18826.76 320.18 39296.41 19146.93 58.44

98004.99 3263.65 25259.02 78181.19 46045.13 1097.38 68236.28 47142.50

31635.94 699.89 13098.67 37040.83 15093.74 371.29 27491.78 15465.03

53194.71 1342.25 11429.40 36703.70 14634.86 304.87 34223.35 14939.73

974.81 5.87 726.85 2307.34 516.10 4.90 1338.29 521.00

158.34

As expected, configuration B presents the largest TAC value since the presence of two columns causes an obvious increase in capital cost as well as an increase in operating cost. Configuration C has the least TAC, being 12.7% less than configuration A. The main reason for the lower TAC is because of the capital saving in the main column of this configuration. As previously mentioned, this column operates with a lower total liquid mass flow, resulting in a column with smaller diameter, which consequently requires less capital investment. The lower total liquid mass flow also directly affects the heat duties of the process. As a result, the TAC of configuration C is lowest. Among the three configurations, configuration C was the most economically advantageous, and therefore a procedure was used to achieve an improved design for further reducing its TAC. 4.7. Optimal Design of Configuration C. Total annual cost (TAC) is used as the economic parameter to be minimized for configuration C. The design variable was the number of trays and the reflux ratio the manipulating variable. The reflux ratio of the column was adjusted to obtain a product with a minimum composition of 0.996 in mass fraction of the two alcohols, isoamyl alcohol and active amyl alcohol. All simulations are carried out using the same specification: the bottom mass flow of the column was fixed at 67.77 kg/h with feed in its middle. The stripping column with three trays operates with distillate mass flow of 0.5 kg/h and reflux ratio of 1. Figure 11 shows the results of the TAC for some simulations performed when varying the number of trays from 25 to 80. It can

51.02

Table 7. Composition of Input and Output Streams of Optimal Configuration C isoamyl alcohol active amyl alcohol isobutanol butanol pentanol propanol ethanol methanol water

feed

distillate

bottom

aqueous

0.5570 0.1207 0.0671 0.0057 0.0003 0.0094 0.0866 0.0002 0.1530

0.010 2.7 × 10−06 0.281 0.015 6.3 × 10−08 0.039 0.363 6.3 × 10−04 0.290

0.818 0.178 1.4 × 10−04 0.003 4.4 × 10−04 3.4 × 10−08 2.7 × 10−10 6.6 × 10−19 2.1 × 10−20

9.6 × 10−05 2.3 × 10−06 2.5 × 10−04 5.9 × 10−07 1.4 × 10−07 8.6 × 10−05 1.4 × 10−03 5.2 × 10−04 0.998

streams. Results for the mass flows are the same as shown in Figure 10. In the next section, a plant-wide control system is developed and tested for configuration C with the aforementioned specifications.

5. PROCESS DYNAMICS AND CONTROL The proper overall control strategy for configuration C was investigated. The Aspen Plus steady-state simulation was exported to the dynamic simulation of Aspen Dynamics. Volumes of the column bases were specified with a holdup time of 10 min and 50% liquid level. The decanter was sized for a 20 min holdup to allow for the two liquid phases to separate. 5.1. Selection of Temperature Control Trays. The slope criterion analysis was used to determine the tray temperature control points. Figure 12 shows the difference between temperatures on adjacent trays for the column and stripping column. There is a large change at the feed stage (stage 22) of the column; however, the feed is not a good location for temperature control.43 A large change in temperature is also observed in stage 11, which was selected for temperature control of the column. The slope analysis suggests the use of stage 3 for temperature control of the stripping column. 5.2. Control Structure. Figure 13 shows the control structure developed for configuration C consisting of eight inventory control loops (six levels and two pressures). In each column the reflux drum level is controlled by manipulating the distillate flow, and the column bottom level is controlled by manipulating the bottom flow. In the decanter, the aqueous phase level is controlled by manipulating the aqueous outlet flow, and the organic phase level is controlled by manipulating the organic outlet flow. Top pressures of the two columns are controlled by manipulating the condenser duties. All level controllers are proportional only with a gain of 2. The temperature of one tray in each column was controlled by manipulating the reboiler heat input in that column. A 1 min dead time is inserted in each temperature loop. Relay-feedback

Figure 11. Effect of the number of stage on TAC’s.

be observed that the number of stages equal to 40 corresponds to the minimum TAC. In this simulation the reflux ratio was 1.3 and SSC = 0.66. Table 7 shows the results for the input and output 2346



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Figure 12. Configuration C temperature profile and slope: (A) column and (B) stripping column.

Figure 13. Proposed overall control strategy.

tuning constants for the third stage temperature loop of the striping column are Kc = 10.12 and τI = 7.92 min. The reflux ratio is held constant, but the reflux flow can still be adjusted in the case of feed composition disturbances.

tests and Tyreus−Luyben tuning were used to obtain controller tuning parameters in the column temperature loops. The resulting PI tuning constants for the 11th stage temperature loop of the column are Kc = 6.41 and τI = 15.84 min, and the 2347



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Figure 14. Responses with feed flow changes (solid line, −20%; dashed line, +20%).

5.3. Dynamic Performance Results. To evaluate the control performance of this control structure, disturbances in feed flow rate and feed composition are made at time = 1 h. Figure 14 shows the dynamic responses for ±20% feed flow rate changes. It is observed that the two tray temperatures are

controlled back to their set point values within 2 h, and the compositions of isoamyl alcohol and active amyl alcohol in the product are held quite close to the desired values. It is also noticed that the bottom mass flow of the column decreases or increases correspondingly with the changes in the isoamyl 2348



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Figure 15. Responses with feed composition changes (solid line, −10% of isoamyl alcohol and active amyl alcohol; dashed line, +10% of isoamyl alcohol and active amyl alcohol).

active amyl alcohol concentrations by ±10%, with the others seven components retaining their in initial ratios to each other. The two tray temperatures were again controlled back to their set point values. A small decrease in purity of the bottom product is

alcohol and active amyl alcohol feed composition, maintaining high recoveries. Figure 15 provides the results for changes in the feed composition by increasing or decreasing the isoamyl alcohol and 2349



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noticed with the composition disturbance of −10%. The bottom outlet mass fraction of isoamyl alcohol and active amyl alcohol deviated from 0.996 to 0.994. This occurs because with the decrease of isoamyl alcohol and active amyl alcohol in feed there was a parallel increase in other higher alcohols such as butanol and isobutanol, thus resulting in an increase of these two alcohols in the final product and decreasing purity of the desired isoamyl alcohol. These simulation results demonstrate that the proposed control structure provides effective control for different disturbances.

(9) Pérez, E. R.; Cardoso, D. R.; Franco, D. W. Análise dos álcoois, ́ ésteres e compostos carbonilicos em amostras de óleo fúsel. Quim. Nova 2001, 24, 10 [in Portuguese]. (10) Aucejo, A.; Burguet, M. C.; Montón, J. B.; Muñoz, R.; Sanchotello, M.; Vázquez, M. I. Vapor−liquid-equilibria for systems of 1-butanol with 2-methyl-1-butanol, 3-methyl-1-butanol, 2-methyl-2-butanol, and 3methyl-2-butanol at 30 and 100 kPa. J. Chem. Eng. Data 1994, 39, 271. (11) Aucejo, A.; Burguet, M. C.; Montón, J. B.; Muñoz, R.; Sanchotello, M.; Vázquez, M. I. Isothermal vapor−liquid-equilibria of 1-pentanol with 2-methyl-1-butanol, 2-methyl-2-butanol, and 3-methyl-2-butanol. J. Chem. Eng. Data 1994, 39, 578. (12) Brunjes, A. S.; Bogart, M. J. P. Vapor−liquid equilibria for commercially important systems of organic solvents: The binary systems ethanol−n-butanol, acetone−water and isopropanol−water. Ind. Eng.Chem. 1943, 35, 255. (13) Cho, T. H.; Ochi, K.; Kojima, K. Isobaric vapor−liquid-equilibria for binary S-systems with limited miscibility, water−normal-amyl alcohol and water−isoamyl alcohol. Kagaku Kogaku Ronbun 1984, 10, 181. (14) Gmehling, J.; Onken, U. Vapor−liquid equilibrium data collection. In DECHEMA Chemistry Data Series; Behrens, D., Eckermann, R., Eds.; DECHEMA: Frankfurt, Germany, 1981. (15) Hellwig, L. R.; Winkle, M. V. Vapor−liquid equilibria for ethyl alcohol binary systems. Ind. Eng.Chem. 1953, 45, 624. (16) Hill, W. D.; Winkle, M. V. Vapor−liquid equilibria in methanol binary systems: Methanol−propanol, methanol−butanol, and methanol−pentanol. Ind. Eng.Chem. 1952, 44, 205. (17) Kurihara, K.; Nakamichi, M.; Kojima, K. Isobaric vapor−liquid equilibria for methanol + ethanol + water and the three constituent binary systems. J. Chem. Eng. Data 1993, 38, 446. (18) Lladosa, E.; Montón, J. B.; Burguet, M. C.; Muñoz, R. Vapor− liquid equilibria in the ternary system dipropyl ether plus 1-propanol plus 1-pentanol and the binary systems dipropyl ether plus 1-pentanol, 1-propanol plus 1-pentanol at 101.3 kPa. Fluid Phase Equilib. 2006, 247, 175. (19) Mohsen-nia, M.; Memarzadeh, M. R. Isobaric (vapour plus liquid) equilibria for the (1-propanol+1-butanol) binary mixture at (53.3 and 91.3) kPa. J. Chem. Thermodyn. 2010, 42, 792. (20) Resa, J. M.; González, C.; Moradillo, B.; Ruiz, A. Isobaric vapor− liquid equilibria of 3-methyl-1-butanol with methanol and vinyl acetate at 101.3 kPa. Fluid Phase Equilib. 1997, 132, 205. (21) Resa, J. M.; González, C.; Goenaga, J. M.; Iglesias, M. Density, refractive index, and speed of sound at 298.15 K and vapor-liquid equilibria at 101.3 kPa for binary mixtures of ethyl acetate+1-pentanol and ethanol+2-methyl-1-propanol. J. Chem. Eng. Data 2004, 49, 804. (22) Resa, J. M.; González, C.; Goenaga, J. M. Density, refractive index, speed of sound at 298.15 K, and vapor−liquid equilibria at 101.3 kPa for binary mixtures of methanol plus 2-methyl-1-butanol and ethanol plus 2-methyl-1-butanol. J. Chem. Eng. Data 2005, 50, 1570. (23) Resa, J. M.; Goenaga, J. M.; Iglesias, M. Vapor−liquid equilibria at 101.3 kPa for binary mixtures containing 2-methyl-1-propanol+2methyl-1-butanol, 2-methyl-1-propanol+3-methyl-1-butanol, and 2methyl-1-propanol+1-pentanol. J. Chem. Eng. Data 2006, 51, 1892. (24) Resa, J. M.; González, C.; Goenaga, J. M. Density, refractive index, speed of sound at 298.15 K, and vapor−liquid equilibria at 101.3 kPa for binary mixtures of propanol+2-methyl-1-butanol and propanol+3methyl-1-butanol. J. Chem. Eng. Data 2006, 51, 73. (25) Tamir, A.; Wisniak, J. Binary vapor−liquid equilibria of some amyl alcohols. J. Chem. Eng. Data 1976, 21, 182. (26) Thiede, S.; Horstmann, S.; Meisel, T.; Sinnema, J.; Gmehling, J. Experimental determination of vapor−liquid equilibria and excess enthalpy data for the binary system 2-methyl-1-butanol+3-methyl-1butanol as a test mixture for distillation columns. Ind. Eng. Chem. Res. 2010, 49, 1844. (27) Wisniak, J.; Tamir, A. Association effects in the methanol−1pentanol system. J. Chem. Eng. Data 1988, 33, 432. (28) Bessa, L. C. B. A.; Batista, F. R. M.; Meirelles, A. J. A. Double-effect integration of multicomponent alcoholic distillation columns. Energy 2012, 45, 603.

6. CONCLUSIONS All proposed processes were capable of obtaining high-purity isoamyl alcohol and active amyl alcohol products with high recovery, but the total annual cost of configuration C is less than the others. The results show that it is not possible to obtain only isoamyl alcohol by conventional distillation processes of fusel oil; thus, a product is obtained consisting of a mixture of the isomers active amyl and isoamyl alcohols. An overall control strategy of this configuration was proposed, and a control performance can be obtained using simple temperature control. The results presented in this work for configuration C provide support for development for fusel oil industrial purification plants to obtain the isoamyl alcohol. Configuration C presents an additional advantage compared to the conventional distillation process, since reduced water content in the column feed stream results in a lower energy requirement.

■

AUTHOR INFORMATION

Corresponding Author

*Fax: +55 19 3521 4027. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

■ ■

ACKNOWLEDGMENTS The authors are grateful to CNPq (131859/2010-2) and FAPESP (08/56258-8) for the financial support. REFERENCES

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(29) Kadir, S.; Decloux, M.; Giampaoli, P.; Joulia, X. Liquid−liquid equilibria of the ternary systems 3-methyl-1-butanol + ethanol plus water and 2-methyl-1-propanol + ethanol + water at 293.15 K. J. Chem. Eng. Data 2008, 53, 910. (30) Ghizellaoui, S.; Coquelet, C.; Richon, D.; Meniai, A. H. Liquid− liquid equilibrium of (water + 1-propanol + 1-pentanol) system at 298.15 and 323.15 K. Fluid Phase Equilib. 2010, 296, 42. (31) Gomis-Yagües, V.; Ruíz-Bevià, F.; Ramos-Nofuentes, M.; Fernández-Torres, M. J. The influence of the temperature on the liquid−liquid equilibrium of the ternary system 1-butanol−1-propanol− water. Fluid Phase Equilib. 1998, 149, 139. (32) Sørensen, J. M.; Arlt, W. Liquid−liquid equilibrium data collection. In DECHEMA Chemistry Data Series; Behrens, D., Eckermann, R., Eds.; DECHEMA: Frankfurt, Germany, 1980. (33) Reid, R. C.; Prausnitz, J. M.; Poling, B. E. The Properties of Gases and Liquids, 4th ed.; McGraw-Hill Book Co.: New York, 1987. (34) Liu, Y.; Gao, H.; Sun, J.; Wang, Y.; Wu, P.; Liu, Y. E; Ji, Y. Novel packing-enhanced distillation separation of isoamyl alcohol isomer: Experimental and pilot scale study. AIChE J. 2011, 57, 3037. (35) Asprion, N.; Kaibel, G. Dividing wall columns: Fundamentals and recent advances. Chem. Eng. Process. 2010, 49, 139. (36) Dejanović, I.; Matijašević, Lj.; Olujić, Ž . Dividing wall columnA breakthrough towards sustainable distilling. Chem. Eng. Process. 2010, 49, 559. (37) Yildirim, Ö .; Kiss, A. A.; Kenig, E. Y. Dividing wall columns in chemical process industry: A review on current activities. Sep. Purif. Technol. 2011, 80, 403. (38) Wang, S.-J.; Lee, C.-J.; Jang, S.-S.; Shieh, S.-S. Plant-wide design and control of acetic acid dehydration system via heterogeneous azeotropic distillation and divided wall distillation. J. Process Control 2008, 18, 45. (39) Kiss, A. A.; Suszwalak, D. J-.P. C. Enhanced bioethanol dehydration by extractive and azeotropic distillation in dividing-wall columns. Sep. Purif. Technol. 2012, 86, 70. (40) Sun, L.-Y.; Chang, X.-W.; Qi, C.-X.; Li, Q.-S. Implementation of ethanol dehydration using dividing-wall heterogeneous azeotropic distillation column. Sep. Sci. Technol. 2011, 46, 1365. (41) Olujic, Z.; Jödecke, M.; Shilkin, A.; Schuch, G.; Kaibel, B. Equipment improvement trends in distillation. Chem. Eng. Process. 2009, 48, 1089. (42) Douglas, J. M. Conceptual Design of Chemical Process; McGrawHill: Boston, 1988. (43) Luyben, W. L. Distillation Design and Control Using Aspen Simulation; John Wiley & Sons, Inc.: New York, 2006. (44) Turton, R.; Bailie, R. C.; Whiting, W. B.; Shaeiwitz, J. A. Analysis, Synthesis, and Design of Chemical Processes, 3rd ed.; Pearson Education: Boston, 2009. (45) Chemical Engineering, Economic Indicators, www.che.com. Accessed July 2012.

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