Process Intensification Using CO2 As Cosolvent under Supercritical

Article pubs.acs.org/IECR

Process Intensification Using CO2 As Cosolvent under Supercritical Conditions Applied to the Design of Biodiesel Production J. Maçaira,† A. Santana,*,‡ A. Costa,§ E. Ramirez,∥ and M. A. Larrayoz‡ †

LEPABE−Faculdade de Engenharia, Universidade do Porto, Rua Dr. Roberto Frias s/n , 4200-465 Porto, Portugal Department of Chemical Engineering. Universitat Politécnica de Catalunya. Av. Diagonal 647, 08028 Barcelona, Spain § Department of Chemical Engineering. Universidade Federal de Santa Catarina, Centro Tecnológico, 88040-900, Florianópolis, Brazil ∥ Chemical Engineering Department, Faculty of Chemistry, Universitat de Barcelona, Martí I Franquès 1, 08028, Barcelona Spain ‡

S Supporting Information *

ABSTRACT: In this work, a new process for biodiesel production under supercritical conditions in presence of cosolvent (CO2) is designed and simulated using the process simulator Aspen Plus. The model was developed using experimental reaction data of continuous catalytic biodiesel synthesis at a 74:25:1 CO2 to methanol to triglycerides molar ratio, temperature range between 150 and 300 °C, at 250 bar. To decrease the temperature and pressure of operation and increase the conversion efficiency of biodiesel, CO2 was added as cosolvent to the reactants. Triolein (C57H104O6) was chosen to represent the vegetal oil and methyl oleate (C19H36O2), biodiesel. A detailed kinetic model based on a three step reversible reaction scheme is used to describe the transesterification reaction in the process simulator. The simulated process resulted in full triolein conversion and a high purity (99.8%) fatty acid methyl esters product. The process plant was designed and simulated to operate in a continuous mode and the annual production capacity of the plant was set at 10000 tons. The total energy for the designed process was 2223 kW. On the basis of a preliminary economical study the minimum selling price of biodiesel is 0.83$/L for a plant capacity of 10000 tons. A profitability analysis was conducted and the payback time was estimated as a function of the selling price and the plant capacity.

1. INTRODUCTION Biodiesel is an interesting alternative diesel fuel that contributes to reduce the environmental impacts in the transportation sector since emissions of the most regulated pollutants are substantially lower in comparison to fossil diesel.1−6 Biodiesel is relatively safe for use in diesel engines and storage in diesel container because of its high flash point. It can be used alone as fuel, or mixed with fossil diesel in diesel engines without major adjustments. An engine using biodiesel has fuel consumption, torsion, and traction ratio similar to that of fossil diesel. Biodiesel offers similar performance and engine durability as fossil diesel7 and there is no limitation on the mixing ratio of biodiesel and fossil diesel. Its use may improve emissions levels of some pollutants and deteriorate others. However, for quantifying the effect of biodiesel it is important to take into account several other factors such as raw material used, driving cycle, vehicle technology, etc. Usage of biodiesel will allow a balance to be sought between agriculture, economic development and the environment.8 Biodiesel (alkyl ester) is usually produced by the transesterification reaction of a lipid feedstock, as shown in Figure 1. Transesterification is the reversible reaction of a fat or oil (which is composed of triglycerides) with an alcohol to form fatty acid alkyl esters and glycerol. Stoichiometrically, the reaction requires a 3:1 molar alcohol-to-oil ratio, but excess alcohol is usually added to drive the equilibrium toward the products side. Transesterification reactions may employ various types of alcohols, preferably, those with low molecular weight. Although methanol is commonly used in the conventional biodiesel production process, ethanol has been shown to © 2014 American Chemical Society

produce a more environmentally friendly fuel, and is less toxic than methanol.9,10 Furthermore, the use of longer chain alcohols such as ethanol, n-propanol, and n-butanol in biodiesel production has shown to lower the cloud-point of the resulting biodiesel11,12 There are many factors that determine the success of biodiesel as a final alternative transportation fuel. Among which, two major factors are substantial to determine the process economy for the biodiesel production processes. The first is that the process has to be able to use the cheapest raw materials as raw material costs are the major part of the total production cost. The second is that the process should be intensified with as few as possible processing steps and equipment requirements. This calls for a systematic approach for biodiesel production process. Conventional biodiesel production processes involve the use of a catalyst, such as sodium hydroxide or sulfuric acid; this generally implies high energy consumptions for the separation of the catalyst and, in the case of basic catalysis, the formation of undesirable products due to the saponification reaction. However, the supercritical process does not require the steps of neutralization, washing, and drying. These steps can be omitted from the process. Furthermore, the supercritical process can increase the yield and allow the waste oil to be used as a raw material, negating the need to pretreat free fatty acids or water. Received: Revised: Accepted: Published: 3985

August 13, 2013 November 22, 2013 February 19, 2014 February 19, 2014 dx.doi.org/10.1021/ie402657e | Ind. Eng. Chem. Res. 2014, 53, 3985−3995

Industrial & Engineering Chemistry Research

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Figure 1. Transesterification reaction scheme.

systems are a very interesting alternative to other processes in terms of energy requirements and capacity for achieving high reactant conversions and product purities. Marulanda28 reported a supercritical process with a 9:1 molar ratio at 400 °C which presented very low energy consumption (573 kW). Glisic et al.27analyzed several different schemes for fatty acid methyl ester (FAME) production at sub- and supercritical conditions, with or without the use of a heterogeneous catalyst, using Aspen Plus, to reduce the energy consumption and improve the life cycle energy efficiency. The use of a heterogeneous catalyst allowed decreasing the energy consumption by half for the supercritical process working at subcritical conditions. The main limitation for industrial application of the heterogeneous process for biodiesel production could be the cost and catalyst lifetime as well as its separation from the reaction mixture. However the simulation of these authors was based on a kinetic model based on the simple assumption that methanolysis of triglycerides could be described as a first order and irreversible one step reaction. In fact the transesterification reaction follows a three step irreversible reaction scheme, so instead of a simple reaction constant, a six reaction constants kinetic model should be used. This way more reliable results should be drawn and thus a more realistic assessment of the process would be accomplished. This paper suggests applying process intensification to biodiesel production. The most common intensification alternative is to carry out the reaction using cosolvents, solid acid catalyst and the recovery of methanol. This work intends to demonstrate the feasibility of using a high pressure continuous reactor to produce biodiesel with methanol and carbon dioxide (CO2) as cosolvent at moderate pressure and temperature. The combined use of supercritical methanol, CO2 as cosolvent and heterogeneous catalysis in a continuous process allows decreasing the temperature and pressure of operation usually employed at supercritical conditions. The residence time is much lower than conventional processes which traduces directly in a decrease in the reactor volume needed. After the transesterification reaction there are only two purification steps needed for obtaining high purity biodiesel and glycerol streams; this allows removing the high cost purification stages needed for removing water from process. The unreacted methanol is reintroduced with high purity into the process, allowing that, in steady state, the only methanol consumed is the one that reacts to form methyl esters. The CO2 used as cosolvent is introduced initially in the desired molar ratio to the other components in

However, because of the high temperature and pressure requirements for the supercritical reaction, high capital and manufacturing costs are expected. Some cosolvents could also act as an agent to reduce the operating temperature and pressure required for the transesterification process particularly for supercritical transesterification process13−18 Currently, the supercritical production process is at the research stage, using mostly a laboratory-scale batch or continuous reactor. To commercialize the process, there are still several challenges remaining in research and development. In particular, high temperature and pressure conditions, as well as excessive amounts of methanol that result in high energy consumption in a supercritical process need to be alleviated. In addition, research on behavior of the reaction system in large industrial plants needs to be done. One of the best ways to conduct an economic assessment of biodiesel production is to develop a simulation model of the process. Simulating a complete biodiesel production process allows investigation of technological feasibility and limitations of the process before economic aspects are examined. Furthermore, the effect of process parameters such as operating temperatures and pressure can be studied and optimized. In this way, a reliable economic analysis of a biodiesel production process can be conducted. Several techno-economic analyses on biodiesel industry have been developed examining its economic prospective and profitability.19−23 Zhang et al.19 have used simulations in HYSYS to compare four different biodiesel processes by studying two different catalysts and two different feedstocks. Moreover, Hu et al.24have implemented life cycle analysis and economic assessment for soybean biodiesel and conventional biodiesel. West et al.25 compared homogeneous base and acid catalyzed, heterogeneous and supercritical processes for biodiesel production from waste vegetable oil using HYSYS Plant. This study showed that the heterogeneous process was the most economically feasible, followed by supercritical process. Glisic and Skala26,27 performed Aspen Plus simulations using supercritical process at temperature and pressure of 300 °C and 200 bar, respectively, with the molar ratio of alcohol to triglycerides was 42:1 and the conventional base catalyzed method without free fatty acids (FFA) pretreatment. The results showed that the amount of energy required for the supercritical process was almost the same as that for the conventional base catalyzed process without FFA pretreatment. Gómez-Castro et al.21−23 have studied the thermally coupled reactive distillation systems as an intensified alternative to conventional biodiesel production. They showed that these 3986

dx.doi.org/10.1021/ie402657e | Ind. Eng. Chem. Res. 2014, 53, 3985−3995


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were presented in databank of Aspen Plus (see Table 1); (2) reacting mixture contained two polar components (methanol

the system; at steady state there is no production, consumption or losses in CO2, so there are no harmful emissions to the environment. Finally the total energy consumption is determined to be lower than most conventional processes. It is assumed a three step reversible reaction scheme with first order kinetics to represent the transesterification reaction. The accuracy of the kinetic model was compared to experimental data provided by Maçaira et al. (2011).17 Simulations were performed on the ASPEN plus process simulator. CO2 emissions and the methanol recovery have also been computed for the transesterification steps. The main aim of this study is therefore to design the supercritical process for biodiesel production considering actual experimental data, and to make a preliminary study on the profitability of the process.

Table 1. Physical and Thermodynamic Properties of Components

molecular weight (g mol−1) normal boiling temperature (K) liquid volume at Tb (m3 kmol−1) liquid molar volume (m3 kmol−1) critical temperature (K) critical pressure (Mpa) critical volume (m3 kmol−1) acentric factor

2. METHODS 2.1. Process Simulation. The process simulation software ASPEN plus version 2007 was used to carry out the simulation of the industrial scaled transesterification process and the design of the reaction and separation units. The procedures for process simulation mainly involve defining the chemical components, selecting a thermodynamic model, selecting proper operating units and their operating conditions (flow rate, temperature, pressure, and other conditions). Information for chemical components, such as CO2, methanol, or glycerol, is available in the ASPEN plus component library. However, sunflower oil, which is a complex mixture of palmitic, stearic, oleic and linoleic oils, is not available in the library. The major component present in triglycerides of sunflower oil is the oleic acid, so triolein (C57H104O2) was chosen to represent the sunflower oil in the ASPEN plus simulation since this component is available in the process simulator database. The product of the process, a mixture of various fatty acid methyl esters (FAME), was determined by the fatty acid composition in the final stream. Methyl oleate (C19H36O2), available in the ASPEN plus component library, was used for representing biodiesel. Diolein (C39H72O5) and monoolein (C21H40O4) were chosen to represent the intermediate species dyglycerides and monoglycerides. The structure of these species were generated with ChemSketch ACD software and imported to the ASPEN plus component library. Their physical properties were estimated from the molecular structures using the group contribution method.29,30 Mass and energy balances for each unit were obtained. Although equipment pressure drops are inevitable, in this study they were neglected. The process plant was designed and simulated to operate in a continuous mode and the annual production capacity of the plant was set at 10000 tons which corresponds to 1225 kg/h of biodiesel (considering 340 working days per year). In all simulations, the FAME final product purity was kept constant. 2.2. Thermodynamic Property Model Selection and Physical Properties Validation. Selection of the appropriate physical property methods, which successfully describes physical properties of the reaction system to be used in the simulation, is essential.31 The key factors considered to choose thermodynamic property model were: the nature of the properties of interest, the composition of the mixture, the pressure and the temperature range and the availability of parameters. To ease the selection of the right physical property method, it was used a decision tree found at Aspen Plus User Guide.32 Taking into account the following considerations: (1) no pseudocomponents in the reacting system, all components

monoolein

methyl oleate

triolein

diolein

885

621

356

296

827

765.03

674.82

595

2.71

1.11

0.53

0.49

0.958

0.623

0.36

0.341

977 0.334 3.25

920 0.505 2.83

835 1.056 1.254

721 1.103 1.108

1.97

1.76

1.53

1.04

and CO2) but no electrolytes; and (3) working pressure of the reactor was high (250 bar), and the mixture was considered in supercritical state; it was chosen a method based on equations of state for modeling. Redlich−Kwong (RK) equation of state was used to correlate the thermodynamic behavior for the system CO2−methanol−triglycerides. Validating the physical properties was the next step. That involved tabulating/reporting mixture properties and comparing the predicted results to experimental data. Experimental phase equilibrium data for ternary system carbon dioxide + methanol + vegetable oil was collected by means of a highpressure variable volume view cell (Phase Monitor, Supercritical Fluids Technologies, Inc., USA, maximum working temperature 403 K) employing the static method. Pressure and temperature of reacting mixture (1:25:74 oil/methanol/CO2 molar composition) were varied to find critical values, which assure the presence of a unique single-phase into the reactor. Figure 2 shows mixture phase equilibrium evolution with increasing pressure under isothermal conditions (393 K). In Figure 2a, it can be observed the presence of liquid−vapor equilibrium at 160 bar. At 190 bar (Figure 2b), mixture starts to be in supercritical state whereas at 210 bar (Figure 2c) mixture is totally under supercritical conditions. Thermodynamic behavior predicted by Redlich−Kwong, as well as Peng−Robinson equations of state was compared to experimental data (Figure 3). It can be noted that RK-EOS P-T envelope fits better than that generated from PR-EOS (5−15% standard deviation). As a result, the former EOS was chosen for the simulation project. For the separation and purification steps that were not carried out at high pressure and temperatures the thermodynamic models used were UNIQUAC and UNIQUAC-LL. 2.3. Kinetic Model. Several kinetic mechanisms have been proposed in literature by different researchers for the transesterification of vegetable oil; pseudo-first-order,33 second-order,33,34and pseudo-second-order.14 This kinetic study was based on our previous work.17 Transesterification consists of three consecutive reversible reactions, where a mole of FAME (E) is produced in each step, and monoglycerides (MG) and diglycerides (DG) are intermediate products of the triglycerides (TG) conversion as presented by eqs 1−3. At each reaction step, one molecule of methylated compound is produced for each molecule of methanol (M) consumed. As a 3987



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Figure 2. Mixture phase evolution with increasing pressure: (a) 160, (b)190, (c) 210 bar. Oil/methanol/CO2 molar composition = 1:25:74, T = 393 K.

d[DG] = k1[TG][A] − k 2[DG][E] − k 3[DG][A] dt + k4[MG][E]

(5)

d[MG] = k 3[DG][A] − k4[MG][E] − k5[MG][A] dt + k6[GL][E]

(6)

d[GL] = k5[MG][A] − k6[GL][E] dt

(7)

d[E] = k1[TG][A] − k 2[DG][E] + k 3[DG][A] dt − k4[MG][E] + k5[MG][A] − k6[GL][E] Figure 3. Predicted P−T envelopes vs experimental data.

d[A] −d[E] = dt dt

k1

(1)

k2 k3

DG + A ⇄ MG + E

(2)

k4 k5

MG + A ⇄ GL +E

(3)

k6

The governing set of differential equations characterizing the stepwise reactions involved in the trasesterification triglyceride can be written as described in eqs 4−9. d[TG] = −k1[TG][A] + k 2[DG][E] dt

(9)

where k1−k6 are reaction rate constants and [TG], [DG], [MG], [GL], [A], and [E] are molar concentration of each component in the reaction mixture. 2.4. Process Design. The flowsheet that represents the developed process is shown in Figure 4. The transesterification process of vegetable oil was performed in a supercritical continuous reactor, followed by downstream purification, which consisted of three steps: methanol recovery by flash separation, glycerol separation by gravity separation, and methyl ester purification by flash separation. On the basis of experimental information17 the ratio of oil/methanol/CO2 used in this process design was 1:25:74. The cosolvent (CO2) was used in this supercritical process because of its proved ability to decrease the temperature and pressure of the supercritical reaction.18 The reaction pressure was set to 250 bar and the temperature range was varied between 50 and 300 °C so its effect could be evaluated. In all simulations the mass fraction of

result, six different rate constants of the reaction are reported for the whole reaction (Table 2). TG + A ⇄ DG + E

(8)

(4)

Table 2. Reaction Rate Constants Used in the Kinetic Model17 reaction direction

rate constants (L·mol−1·min−1)

Ea (kJ·mol−1)

k0 (L·mol−1·min−1)

TG → DG DG → TG DG → MG MG → DG MG → GL GL → MG

k1 k2 k3 k4 k5 k6

16.20 20.48 13.02 3.29 24.56 17.69

10.24 14.95 8.83 4.78 19.87 14.72

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Figure 4. Process flowsheet for biodiesel production used for the developed process.

3. RESULTS AND DISCUSSION 3.1. Temperature Influence in the Reaction Products and in Total Energy Consumption. The transesterification reaction is favored by the increase of the reactor temperature as can be seen in Figure 5. Here the reaction products mass

methyl oleate in the product stream was maximized, so the effect of temperature in the global process could be optimized. The feed stream is a mixture of methanol, carbon dioxide and triolein. This stream also includes the unreacted methanol that is recovered after the reaction takes place further ahead in the process and mixed with fresh methanol in Mixer I. The methanol and carbon dioxide streams are pressurized to 80 bar in Pump I and in Compressor, respectively, before being cooled to 8 °C by Heat Exchanger I. This means that the mixture of methanol/carbon dioxide is in the liquid phase which allows it to be pressurized by Pump III to the desired reaction pressure, 250 bar. After this stream is properly pressurized it is mixed, in Mixer III, with the previously pressurized triolein feed in Pump II. This mixture of methanol/carbon dioxide/triolein forms the reaction inlet stream that is heated to the desired reaction temperature in Heat Exchanger II. Heat exchangers before the reactor allow the preheating of the reacting mixture by the stream exiting the reactor. Then, the reaction takes place in a fixed bed reactor containing the solid acid catalyst. After the transesterification reaction the outlet stream is then depressurized in Expansion Valve to the operating pressure of the next flash unit, Separator I. Here, at 1.7 bar and 87 °C, the unreacted methanol and carbon dioxide are separated from the reactor outlet mixture of remaining triglycerides, diglycerides, monoglycerides, glycerol, and the methyl esters. This stream is depressurized to ambient pressure and is sent to a gravity separator where glycerol is separated at 20 °C. The resulting stream is sent to purification in Flash Separator II where fatty acid methyl esters can be obtained in compliance with the European standard (EN14124). In the Separator II, besides the biodiesel stream obtained in the bottoms, the upper stream contains traces of methanol/carbon dioxide that is sent to be mixed in Mixer III with the recovered stream of methanol/ carbon dioxide in the flash evaporator after the reactor.

Figure 5. Mass fractions of the reaction products for different reaction temperatures.

fractions were plotted as a function of the operating temperature. After 90 °C the reaction reaches equilibrium and the reaction products are essentially the same; thus could be expected that the best operating temperature for the reactor would be 90 °C to minimize the energy consumption. However this is not verified. To check the best operating temperature of the reactor, the plant total energy consumption has been computed for several reactor working temperatures. 3989



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Figure 6. Heater II and Separator I relative energy consumption, relative total energy consumption, and methyl oleate mass fraction as a function of the reactor operating temperature.

Figure 7. (a) Influence of Flash Separator I operating temperature in the recovery of methanol, carbon dioxide and methyl oleate; at 87 °C (vertical dashed line) there is maximum recovery of the three components. (b) Expansion valve stream outlet temperature and liquid and vapor fraction as a function of the reactor temperature; at reactor working temperature of 200 °C the depressurization at the expansion valve causes a decrease in the outlet stream (stream 13) temperature from 200 to 87 °C, which is the optimum working temperature of Flash separator I, as seen in panel a.

the energy necessary to separate and purify the reaction products further ahead is the independent of the reactor temperature. In this second region, the energy consumption is limited by the energy spent in Separator I. The total energy consumption decreases in the second region down a minimum verified at reactor temperature of 200 °C; after this the energy consumption starts to increase again linearly with the reactor temperature. To understand the behavior of this curve let us look to the process happening at Separator I. The operation temperature of Separator I has been optimized using with the two following objectives: complete FAME recovery in the liquid stream, and maximize the methanol/carbon dioxide separation. Figure 7a shows the recovery of each component, in mass weight percentage (wt %), for several values of operating temperatures of the flash separator. As the temperature of the flash separator increases, the recovery of methanol increases up to complete recovery at around 150 °C. The recovery of FAME is complete up to 87 °C; however for temperature higher than 87 °C the recovery of

Figure 6 presents the data of the total energy consumed in the plant computed for several reactor working temperatures; energy consumption data of Separator I and Heater II is also shown, as well the relative mass fraction of methyl oleate at the outlet stream of the reactor. As can be seen, the plant total energy consumption does not have a liner relation with reactor temperature. As expected the energy consumption of Heater II increases with the reactor temperature, because this unit is responsible for preheating the reaction mixture to the operating temperature of the reactor. However, the total energy consumption of the plant does not accompany this trend. In the plant energy consumption curve can be distinguished three regions: in the first, for reactor temperatures up to 75−80 °C, the total energy consumption increases exponentially with the reactor temperature; in this region the total energy consumption is manly limited by the Reactor and Heater II energy consumptions that depend on their respective temperatures. However, after 80 °C, the conversion of triolein is complete and the reaction products are essentially the same, so 3990



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FAME starts to decrease. So the best operating temperature for the separator is chosen according to the premise that all the FAME in the product must be recovered. At the chosen operating temperature there is 99.99% recovery of methyl oleate, 97.6% of methanol, and 99.9% of carbon dioxide. The depressurization from 250 to 1.7 bar occurred at the expansion valve causes an abrupt temperature decrease. For each working temperature of the reactor the expansion valve outlet stream temperature is lower than the reactor temperature and influenced differently by each case in the depressurization process. In Figure 7b, the reaction products stream temperature after the expansion valve is plotted for each reactor working temperature. For example, for reactor temperature of 50 °C the expansion causes a temperature decrease to −32 °C. So in this case it is necessary to heat the stream at the next separation unit, Separator I from −32 to 87 °C. For reactor temperature 100 °C the expansion causes the stream to be at 11 °C, so it is necessary to heat the separator from 11 to 87 °C, this is 76 °C. This fact remains valid up to 200 °C, with decreasingly less energy required to heat the separator. In the case of 200 °C the temperature decrease is about 113 °C, which causes the reactor outlet stream to be at 87 °C, which is the same conditions that is necessary to separate the methanol/carbon dioxide from the rest of the components at the separation unit, Separator I. In theory at this working temperature of the reactor there is no energy consumption in separating these components (see Figure 6). So, in conclusion, below 200 °C the depressurization lowers the temperature to a value that is less than the temperature necessary in the separation unit, 87 °C, so it is necessary to spend energy to heat the separator. Above 200 °C, the opposite situation is verified: the reaction outlet stream needs to be cooled to 87 °C, and so, more energy is required. In summary, the use of 200 °C as the reactor temperature, which favors reaction rate and triolein conversion, results in the the depressurization to 1.7 bar causing the (optimum) separation of methanol and carbon dioxide from methyl esters and glycerol, without the need of any heat exchangers and no extra energy consumption. Although the operating temperature chosen is below the maximum allowed temperature of the resin Nafion SAC-13, 210 °C, the stability and duration of this resin should be investigated when used at high temperatures for long periods of time. Further studied about the durability of this resin would be very interesting in order to validate their use in industrial applications. 3.2. Process Simulation Results. According to the designs principles found at several chemical process equipment’s selection and design books,35,36 the fixed bed reactor was designed to have an aspect ratio between its length and diameter of 3 and a residence time that allowed the full conversion of triglycerides. The catalyst mass was calculated to fill all the reactor volume. The results show that there is full triglyceride conversion with reactor dimensions of 3.0 m × 1.0 m and residence time just over 4 min. Because there are always deviations from calculations and real results, the safe design of the reactor’s dimension should be higher than the theoretically predicted for the complete conversion of the products. Figure 8 shows the simulated concentration profiles of the reaction species, including the intermediate species. The kinetic model allows confirming if the designed reactor allows the equilibrium to be reached and most important checking if there is any intermediate species at the reactor outlet. As can be seen

Figure 8. Simulated concentration profile of glycerol (G), triglyceride (TG), diglyceride (DG), monoglyceride (MG), and fatty acid methyl ester (FAME) during transesterification vs reactor length at 200 °C and 250 bar.

in Figure 8, all mono and diglycerides are consumed and are reacted to form methyl oleate in the first half of the reactor. Not knowing the intermediate species concentration profiles could lead to wrong dimensioning of the reaction vessel and its operating conditions and thus higher amount of energy required for the purification step or even unwanted final product specifications. In Figure 9 the concentrations profiles along the reactor length are plotted for several operating temperatures. As can be seen temperature plays an important role in reaction progress. For example for operating temperature of 50 °C there is high monoglycerides content present at reactor outlet and diglycerides are consumed only at the very end of the reactor (L = 2.5 m). The content of mono and diglycerides tend to zero at lower reactor lengths by the increase the reactor temperature. At 200 °C, monoglycerides disappear at the beginning of the reactor and diglycerides are consumed at half the reactor length, ensuring maximum conversion of triolein to methyl oleate. With this information and bearing the considerations made in the previous section, 200 °C is the best operating temperature for the reactor. Results from the main process units and its global mass balances are presented in Table 3 and Table 4. Table S1 in Supporting Information gives an overview for the units operation in the process. Methanol recovery is 98%, which contributes to the minimization of the overall reagent consumption despite the excess molar ratio in which the transesterification reaction is carried out. The recovery of glycerol in the gravity separator is 99.5% with 90.0% of purity. The final stream of methyl esters is produced at 1225 kg/h with 99. 8% mass fraction and with glycerol content of 6.1 × 10−2 wt %. The feed and product information is shown in Table 4. The final product stream consists in 99.8% of methyl oleate with some traces of carbon dioxide and a mass fraction of glycerol that is below the maximum allowed by the European Norm (EN 14124) for biodiesel purity (0.25 wt %). Although the reaction is carried out with excess methanol of 25:1 molar ratio 3991



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Figure 9. Intermediate species, (a) monoglycerides and (b) diglycerides, mass composition profiles in the reactor for several operating temperatures.

Table 3. Global Mass Balances in Kilograms Per Houra

a

equipment

feed

outlet

Reactor Mixer I Mixer II Mixer III Separator I Separator II Gravity Separator

5676 4451 5676 4303 5676 1243 1385

5676 4451 5676 4303 5676 1243 1385

Accumulation = 0.0.

Table 4. Feed and Product Stream Information for the Simulated Process feed streams stream information

methanol

temperature (°C) pressure (kPa) mass flow (kg h−1) methanol CO2 triolein dioolein monolein methyl oleate glycerol

CO2

product streams triolein

20 20 20 101 101 101 148.7 0a 1225 component mass fraction 1 0 0 0 0 0 0

0 1 0 0 0 0 0

0 0 1 0 0 0 0

9.9 × 10−4 2.09 × 10−5 0 0 0 0.998 6.07 × 10−4

FAMES

glycerol

20 101 1225

20 101 126.7

Figure 10. Methyl esters and glycerol component mass fraction for the experimental laboratory unit and Aspen simulator for 10000 tons/year.

0,102 1.94 × 10−3 0 0 0 5.03 × 10−15 0.896

was used in process simulator. The difference verified could be due to internal or external mass transfer limitations present in the real system and not considered in the simulation. Even though the results are good indication that the simulation engine coupled with an experimental valid kinetic model is a powerful tool for designing the process. The equipment energy balances are shown in Table 5. The most energy consuming process units are the reactor and Heater I, that is responsible by the condensation of the mixture carbon dioxide/methanol necessary to its pressurization to the reaction pressure. The total energy consumption is 2223 kW. Table 6 shows relative comparisons in energy consumption between the conventional and supercritical processes. Lee et al.37 simulated a conventional alkali catalyzed process and a supercritical methanol process and their results showed the calculated energy consumption was lower for the conventional process (2349 kW) than for the supercritical process (3927 kW). Glisic and Skala27 reported two different continuous processes (supercritical process and alkali catalyzed at low temperature and pressure) for analysis of total energy consumption in biodiesel production. This study showed that total consumption energy was very similar in both processes: 2326 kW for conventional and 2407 for supercritical. Marulanda28 investigated the simulation of a biodiesel

a

In steady state there is neither consumption nor losses in CO2. The actual quantity that flows in the process is 3321 kg h−1.

to triolein, it is necessary to feed only 149 kg/h of methanol to the system and 1225 kg/h of triolein. Recovery of 98% of the unreacted methanol explains this fact and means that although the high alcohol/vegetable oil molar ratio used the consumption of methanol is rather low. The simulated results for a 10000 ton/year of biodiesel production were compared with the results from or previous work based on an experimental laboratory unit of biodiesel production in the same operation conditions.17 Figure 10 shows the methyl esters and glycerol mass fraction at the reactor outlet for the Aspen plus simulation and the experimental laboratory unit. Although the difference in the production capacity, the results are very similar because the kinetic model determined in the experimental laboratory unit 3992



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yields at mild temperature conditions thereby reducing the total energy consumption. 3.3. Preliminary Economical Assessment. Estimations of the cost of equipment and other costs related to the capital investment play a crucial role in selecting a particular process from other design alternatives. In this section, a preliminary economical assessment of the developed process is conducted. The objective is to have some insight whether the process could be profitable or not. The methodology employed relies on the sizing of the major process units, with the configuration and operating conditions discussed in the previous section, and the selection of the construction materials. Then the total capital investment is determined, by using the equations of purchase cost for each specific equipment that can be found in equipment design literature.35 The purchase costs have been updated using the Chemical Engineering Plant Cost Index (CEPCI) of December 2012 (see Supporting Information for more details). The sizing and purchase costs of each equipment can be found in Tables S2−S9 in the Supporting Information. With this information the annual costs, annual earnings and profitability could be predicted. The total investment cost was estimated to be roughly 2.4 M$ for the 10 kton/year plant. This cost includes not only the equipment acquisition, installation, instrumentation, and tubing, but other costs that can be found in Table S11 in the Supporting Information. The production costs, presented in Supporting Information Table S13, were estimated to be about 15.3 M$/year and include not only production direct costs, such as raw materials,37 work labor, and maintenance, among others, but also fixed costs, such as amortization, local taxes, and insurance costs. The actual earnings were considered to be the sale of the biodiesel itself and the glycerol byproduct and were estimated in approximately 16 M$/year for the 1 kton/year plant. The prices of raw materials used were conservative values based on previous reported values20,37,38 and can be found in Table S12 in Supporting Information. The approximate profitability of the projected was computed using the payback period, which is the time required for the annual earnings to equal the original investment.35 Figure 11a presents the annual cash flow for the designed 1 kton/year plant, considering several selling prices of biodiesel. The selling price of the glycerol byproduct was set constant to 900 $/ton. As can be verified in Figure 11a), the final selling price of the product has a large influence in the payback time. For the designed process, the minimum selling price of biodiesel for profitability was determined to be $0.83 per liter. For an intermediate value of 0.89 $/L the original investment is expected to be paid in two years. When the demand is high, advantage can be taken out of the economy of scale. This principle assumes that every piece of equipment can be made larger as the production rate increases. Although it can be a considered a rough approximation, and

Table 5. Equipment Energy Balance and Energy Consumption (kW) in

out

energy consumption (kW)

Pump I Pump II Pump III Heater I

−313.7 −21.8 −10910.3 −10269.2

−311.8 49.0 −10832.3 −10978.3

1.90 70.8 78.0 709

Heater II

−10851.0

−10315.9

535

Reactor

−10315.9

−11041.6

726

Separator I

−11041.6

−11039.4

Separator II

−884.6

−847.4

37.2

Gravity Separator

−1104.2

−1167.7

63.5

equipment

total (kW)

2.10

energy carrier electricity electricity electricity refrigeration fluid heat transfer fluid heat transfer fluid medium pressure steam medium pressure steam medium pressure steam

2223

production process by supercritical transesterification at a continuous reactor working at 400 °C and his results showed that the total energy consumption of this process was 573 kW but with a final biodiesel purity of 88.9%. Our study showed a total of energy required for the supercritical process would be around 2223 kW with 99.8% of methyl esters mass fraction in the final product. Comparing our work with the literature, we obtained the second lowest value of the energy consumption while maintaining high biodiesel purity in compliance with the EN 14214 standard. These results should be considered in designing future industrial production processes for biodiesel synthesis not only because the good results here disclosed but also because they include a detailed kinetic model. The supercritical process is mechanistically much simpler than the conventional process. The conventional process requires more units such as mixers, pumps, phase separators, or even distillation columns, because of additional steps involved in the catalyst neutralization and separation steps. Therefore the supercritical process using CO2 as cosolvent may reduce the energy consumption due to simplified separation and purification steps. The difference for the energy consumption of our supercritical process with the other reported in the literature can be due to excess methanol employed in the process (42:1), which exhibits high energy consumption in methanol heating up and recycling; the fact that our process uses CO2 as cosolvent increasing the rate of supercritical transesterification making it possible to obtain high biodiesel

Table 6. Energy Consumption of the Designed Process and Some Published Results

a

process

temperature (°C)

methanol/oil molar ratio

energy consumption (kW)

purity (wt %)

ref

conventional supercritical conventional supercritical this worka supercritical

60 350

6:1 24:1

300 200 400

42:1 25:1 9:1

2349 3927 2326 2407 2223 573

99.5 99.8 99.8 99.8 99.8 88.9

Lee et al.37 Lee et al.37 Glisic and Skala27 Glisic and Skala27 this work Marulanda28

Supercritical/heterogeneous catalyzed, with CO2 as cosolvent. 3993



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literature, is the most environmentally favorable. Using a cosolvent allows reducing the reaction conditions. Although the supercritical methanol process is highly energy intensive, the downstream operations of methanol recovery and products purification are much simpler. If available, the use of a detailed kinetic model of the transesterification reaction should be employed as it is a powerful tool in designing the process units, mainly the reactor, allowing predicting if intermediate species of the transesterification reaction will be present in the outlet product stream. This allows maximizing the methyl esters yield in the final product while minimizing the necessary energy to purify the final stream. The preliminary economic assessment of the project indicates that the biodiesel selling price and the vegetable oil buying price have a great impact in the profitability and in the viability of the developed process. However even for a relatively small production capacity plant of 10 kton/year, the payback time can be achieved within two years of operation, for a biodiesel selling price of 0.89 $/L. To decrease the selling price of the product a higher production capacity plant should be used. It was estimated that a 50 kton/ year plant could achieve profitability in two years for a biodiesel price of 0.75$/L. Although the estimated values could be considered a rough preliminary approximation, they are an indication that the combination of heterogeneous catalysis with the supercritical process with cosolvent is an alternative process to conventional biodiesel synthesis that should be taken into account when designing biodiesel plants.

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ASSOCIATED CONTENT

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AUTHOR INFORMATION

S Supporting Information *

Overview of the main process units, pump specifications, flash separator specifications, heater specifications, and cost accounting and capital estimation. This material is available free of charge via the Internet at http://pubs.acs.org. Figure 11. (a) Annual cash-flow for several biodiesel selling prices and (b) payback time as a function of the biodiesel selling price, estimated for several plant production capacities.

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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considering a single train plant, with no or few equipment duplicates, the cost variation with capacity can be correlated by the following equation:35 ⎛ capacity ⎞m cost 2 2⎟ = ⎜⎜ ⎟ cost1 capacity ⎝ 1⎠

ACKNOWLEDGMENTS The authors would like to acknowledge Spanish Ministry of Science, Technology and Innovation (Grant No. ENE 200914502) for the financial support given.

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(10)

where the exponent factor m can vary from 0.38 to 0.9. In this study, the m factor was set to 0.9. Using eq 10, the payback time as a function of the biodiesel price in dollars per liter was computed for several plant capacities (Figure 11b). As could be expected, increasing the plant capacity has a positive impact in the profitably of the project. For instance, considering a 50 kton plant and a 0.75 $/L biodiesel price, the expected payback time is two years, which is a acceptable value for high-risk ventures.35,36

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4. CONCLUSION This work combines the technologies of heterogeneous catalysis and the supercritical process, making the most of the advantages of each to create a continuous process. The heterogeneous catalyzed supercritical transesterification, using CO2 as cosolvent, compared with the other process in the 3994



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NOTE ADDED AFTER ASAP PUBLICATION This paper was published ASAP on February 27, 2014, with a misspelling in the title. The corrected version was reposted on February 28, 2014.

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Process Intensification Using CO2 As Cosolvent under Supercritical

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