Article pubs.acs.org/EF
Transesterification of Triglyceride to Fatty Acid Alkyl Esters (Biodiesel): Comparison of Utility Requirements and Capital Costs between Reaction Separation and Catalytic Distillation Configurations Aashish Gaurav, Mateus Lenz Leite, Flora T. T. Ng,* and Garry L. Rempel Department of Chemical Engineering, University of Waterloo, Waterloo, Ontario, N2L 3G1 Canada ABSTRACT: The efficacy of catalytic distillation (CD) and the economic merits that it would bring into the biodiesel production process is studied. Process flow sheets depicting conventional and CD technology are modeled in Aspen Plus, and detailed operating conditions and equipment designs are provided for each process. The feedstock considered is soybean oil, and the transesterification reaction for the triglyceride is considered for biodiesel production. After optimal design of both process flow sheets to produce 10 million gallons of biodiesel per year, adhering to ASTM purity standards, a detailed cost analysis is carried out using the Aspen economic analyzer tool to predict capital, operating, and utility costs and to calculate the cost of production per gallon of biodiesel. Results depict CD to be a promising candidate to replicate the conversion and product purity of conventional biodiesel processes while having significant savings in capital (41.42% cheaper than the conventional process) and utility (18.12% less than the conventional process) costs, thereby making it a very competitive alternative. The total operating costs and price of production per gallon of biodiesel are only meagerly cheaper for a CD process because the most significant factor to the biodiesel production process is the raw material cost. For both processes, the price of production per gallon of biodiesel after accounting for revenue generated from glycerol product is predicted to be around 1.7 dollars/gallon. The Aspen model is flexible to accommodate higher flow rates for scale-up of operations, add or remove stages of operation into the biodiesel process, modify feedstock and stream prices, and predict associated capital and production costs.
1. INTRODUCTION The challenge to meet the world’s growing energy demands in a sustainable manner is a key global priority, necessitating intensive research and development of advanced renewable energy systems for oil-based energy and other fossil fuels. Energy security, environmental concerns, foreign exchange savings, and socioeconomic issues have made clean biofuels derived from biomass or biological sources a hot topic. Among the various alternative fuel options, biodiesel has spurned much interest and popularity. Chemically, biodiesel is a mixture of fatty acid alkyl esters (FAAEs), derived from triglyceride molecules. Triglycerides and alcohol are converted to alkyl esters (biodiesel) via a catalyzed transesterification reaction, as illustrated in Figure 1. Glycerol
attractive as a renewable and environmentally friendly fuel option for domestic engines; apart from being nontoxic and biodegradable, biodiesel results in lesser air pollutants per net energy than diesel [reduction of greenhouse gases (GHGs) by 41%2] and an estimated 93% more energy than the energy invested in producing it.2 Biodiesel also has a higher flash point temperature of 150 °C, making its handling, use, and transport safer than petroleum diesel and better lubricating properties that minimize engine wear and tear. A promising alternative to petroleum, biodiesel use and production is hence expected to rise significantly in the near future. The objective of this research is to investigate the possibility of intensifying the conventional reactor separation technology for continuous production of biodiesel via catalytic distillation (CD). CD is a hybrid reactor technology that integrates heterogeneous catalytic reaction within a distillation column by immobilizing solid catalyst particles within discrete reactive sections.3 The distinguishing feature of CD is its ability to simultaneously carry out the chemical reaction and product purification within a single-stage operation, potentially resulting in significant energy savings as well as a reduction in operating and capital expenditures because of process intensification. The continuous removal of product from the reactive section via the distillation action can lead to an increased product yield and increased productivity, particularly for reactions that are
Figure 1. Transesterification (methanolysis) of triglyceride molecule to FAAEs (biodiesel).
and FAAEs are then separated and purified. Rapeseed and soybean oils are the leading vegetable oils for biodiesel production. The conventional catalysts for the transesterification reaction of these vegetable oils are predominantly basic mostly alkaline or alkaline earth oxides or alkoxides.1 Biodiesel is very © 2013 American Chemical Society
Received: September 4, 2013 Revised: October 9, 2013 Published: October 10, 2013 6847
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production process, especially in terms of reduction of water usage for cleaning the products because water is now considered as a valuable resource. In addition, homogeneous catalysts carry the disadvantage of causing severe environmental hazards in the case of spillage. Another novelty of our research is use of single software (Aspen) comprehensively for both simulation and equipment sizing and cost analysis. Most past researchers have carried out economic analysis of their simulations by first sizing the equipment through non-iterative calculations and then employing published equipment cost correlations9 or employing another software (for example, Chemsoft) to estimate capital costs using the design data.6 Aspen Plus is one of most extensively used chemical engineering packages, and use of a single, consistent software environment for simulation, equipment sizing, and cost analysis reduces estimation variation and gives more accurate and reliable costing results. To highlight the merits of CD, two process configurations depicting continuous production of biodiesel are modeled in Aspen Plus. Configuration A (Figure 2) is a traditional reactor separation flowsheet, where transesterification takes place in the reactor and the products are separated in a column. Configuration B (Figure 3) is a CD column that incorporates both the reaction and separation in one column. For the same biodiesel production capacity, a comparison of the total energy requirements and the capital costs is then performed between the two process configurations. The savings in energy requirements and capital costs are then quantified to relate the effectiveness of CD compared to the conventional process.
equilibrium-limited. Other potential advantages of CD include the mitigation of catalyst hot spots, improved temperature control, and enhanced energy integration because of conduction of an exothermic chemical reaction in a boiling medium with in situ separation via distillation. The transesterification of triglyceride meets the design criteria of CD because the products glycerol and FAAEs have a significant volatility difference that makes separation by distillation favorable (properties listed in Table 1). Second, the exothermicity of the reaction as reported in Table 1. Boiling Point and Density of the Components at 1 atm and 25 °C molecular weight (g/mol) boiling point (°C) density (g/cm3)
triolein
methanol
methyl oleate
glycerol
885.43 846.5 0.91
32.04 64.7 0.79
296.49 349 0.87
92.09 287.71 1.26
the literature4 and also confirmed via Aspen calculations in section 2.1 favors CD because the energy liberated can be efficiently converted in situ to drive the distillation process and enhance energy integration. Third, the transesterification is a reversible reaction, and the constant removal of products should shift the equilibrium toward the products;5 hence, CD becomes a very encouraging option. To the knowledge of the authors, in the literature to date, no systematic, comprehensive work has been performed comparing the capital and operating costs of biodiesel production using CD to the traditional process while meeting the optimal process conditions and purity requirements. A number of authors6−10 have analyzed the cost of the biodiesel production process considering traditional reaction and separation technology and homogeneous catalysts. None of these simulations consider use of heterogeneous catalyst systems (that would save catalystwashing operations) or examine the suitability of CD as an alternative to the traditional process. We feel that, despite the slower reaction rate compared to homogeneous catalysts, heterogeneous catalysts could bring merits to the biodiesel
2. PROCESS DESCRIPTION Batch experiments on transesterification of soybean oil with methanol were recently carried out in our laboratory over calcium oxide supported on Al2O3 as solid base catalysts (heterogeneous catalyst system).11 The reaction was assumed to be pseudo-first-order with respect to the triglyceride (soybean oil molecule) in excess of methanol.12 Average overall reaction constants at different temperatures were calculated, and the activation energy was estimated to be 30 kJ/mol based on the Arrhenius equation.
ln k = −
Ea +C RT
(1)
Figure 2. Configuration A: traditional reactor separation flow sheet for biodiesel production. 6848
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Figure 3. Configuration B: CD flow sheet.
Figure 4. Conventional flow sheet for biodiesel production from biomass feedstock. conversion per unit of volume, and they also require lower maintenance and shutdown times.14 A schematic diagram for a conventional biodiesel production process is shown in Figure 4. Most researchers6,8,10 studying the biodiesel production process using homogeneous catalyst systems base their simulations on the shown schematic (Figure 4). The basic steps are the
The current study models the above reaction kinetics in a continuous reactor. The transesterification reaction for biodiesel production in a continuous process can be carried out in different reactors, such as a plug-flow reactor or combined stirred-tank reactor;8,13 accordingly, the reactor conditions, such as volume and residence time, to achieve the same reactant conversion could differ. Plug-flow and packed-bed reactors are known for achieving the highest 6849
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transesterification reaction followed by alcohol recovery in a distillation column and glycerol catalyst extraction with water from the oil phase in a washing column. Some processes also require neutralization of the catalyst. Separate distillation columns are used for fatty acid methyl ester (FAME) and glycerol purification. In our flow sheet for the conventional reactor separation configuration with a heterogeneous catalyst, no washing column is required for extraction of the catalyst from the oil/ biodiesel phase; we are considering a plug-flow packed-bed reactor with a catalyst lifetime of 1 year. Economics for the heterogeneous catalyst system is discussed in section 3. We have also added a flash distillation unit after the high-temperature, high-pressure reactor operation because it brings enormous reduction (4 times less) in utility usage for methanol separation. The process simulations of the complete flow sheets of two configurations (conventional reactor separation and CD) are carried out in the process modeling software Aspen Plus. Although there ought to be differences between process simulation results and actual process operations, simulation software tools predict reliable and pretty accurate information on process operation because of their comprehensive thermodynamic packages, rich component data, and astute calculation techniques. Process simulations generally involve defining chemical components, the choice of an appropriate thermodynamic model, determination of operating conditions (temperature, pressure, concentrations, etc.), and sizing the operation (production capacity). The composition of soybean oil is reported to be 6.8% palmitic acid, 33.4% oleic acid, 51.5% linoleic acid, and 2.3% linolenic acid.15,16 Because the major components, oleic acid, linoleic acid, and linolenic acid, are all composed of 18 carbon atoms, triolein (C57H104O6) was chosen to represent the triglyceride molecule in soybean oil. Hence, the input components are triolein and methanol as the reactants and glycerol and methyl oleate or FAME as the products. Figure 5 shows the reaction. All of the components, except
length, diameter, and number of tubes, we also kept a watch on the residence time for the reaction while working in the specified pressure and temperature ranges. We aimed at more than 95% conversion of triolein to biodiesel to ensure ease of separation and product purity. The reaction time for the transesterification kinetic data used in our Aspen models was reported in the range of 8−12 h.11 The residence time for our Aspen simulation for configuration A is 16 h. The reactor length affects the conversion very strongly, and diameter and number of tubes affect the residence time. The heat of reaction was calculated via Aspen by running a separate RStoic reactor model at constant temperature and pressure and computing the reaction heat. The heat of reaction was found to be −100.17 kJ/mol per mole of biodiesel produced at 25 °C and 1 atm. The reaction is therefore exothermic, making of it a good candidate for a CD operation. The heat of reaction calculations was not used anywhere as an input in the simulation but was calculated to provide an understanding of the temperature profiles, process results, and energy requirements better. The main processing equipment units for the flow sheet are shown in Figure 6. The flow sheet elements are a mixer, a plug-flow fixed-bed reactor (PFR), a flash separator, a decanter, and two distillation columns (RADFRAC). The mixer functions to enable the recycle of the overhead methanol streams from the flash and distillation columns back into the reactor. The flash separates methanol from glycerol and methyl oleate. The decanter splits methyl oleate (biodiesel product) from glycerol. The purpose of the first distillation RADFRAC column (FAME purification, DISTL1) is to purify the biodiesel from methanol, while the second column (glycerol purification, DISTL2) separates methanol from glycerol. A constant feed with a molar flow rate of 100 lb mol h−1 in a methanol/oil optimum molar ratio of 9:1 (for maximum conversion11) is fed to the reactor. The reactor configuration is shown in Table 3. A conversion of 99.83% is achieved. The output stream from the reactor is at 160 °C and 15 atm. Pure component physical properties obtained from the Aspen library suggest that the product separation should be easy because they have appreciable boiling point and density difference (Table 1). Because the boiling point of methanol (64.7 °C at 1 atm) is significantly lower than that of methyl oleate (343.85 °C at 1 atm) and glycerol (287.85 °C at 1 atm), a flash separator was used to isolate methanol from glycerol and methyl oleate. The flash drum works on the principle of a high-pressure gradient to achieve a good separation. The steps were methanol separation followed by the decantation of glycerol and final distillation of the products. The flash drum final operating conditions were 125 °C and 1 atm. The top stream from the flash has a methanol purity of 99%, which was recycled back to the reactor via the mixer. Design and tear stream specifications were specified in Aspen to estimate the fresh feed composition to the mixer to keep the methanol/triolein ratio of 9:1 at the mixer output. The bottom stream from the flash separation unit was taken to a decanter for separation of glycerol and methyl oleate via gravity separation. The decanter functions to purify the biodiesel product (methyl oleate) from methanol and glycerol based on the density difference and intermolecular interaction between them. Separation using a gravity settler has also been proposed.17 The decanter operates at 125 °C and 1 atm. The lighter stream from the decanter is composed of 90.89 mol % methyl oleate (biodiesel) and 8.77 mol % methanol. The denser stream is composed of a glycerol mole fraction (85.32 mol %) and methanol (14.66 mol %). Two RADFRAC distillation columns (FAME purification and glycerol purification) are employed for methanol recovery as well as purification of biodiesel and glycerol. The FAME purifier had five stages (condenser, three trays, and reboiler); the feed was fed to stage 3. The glycerol purifier only had three stages (condenser, tray, and reboiler), with feed on stage 2. Running simulations at varying reflux ratios resulted in minimum energy requirements at 0.5; therefore, this was set as the optimum reflux ratio for both columns. The column stage efficiency was assumed to be 1, and no pressure drop was considered to take place in the columns. The distillate was totally condensed. The FAME purifier distillation column achieved a separation of 99.99% mole fraction of methanol on the top stream (99.99% mass fraction)
Figure 5. Transesterification of the triolein molecule to methyl oleate (biodiesel). for triolein, had their property available in the Aspen pure component database. The triolein molecule file was imported from the National Institute of Standards and Technology (NIST) chemistry web book database, and its properties were downloaded using the NIST thermo data engine. All missing pure component properties and binary interaction parameters were estimated by Aspen. The process-type properties were chosen on the basis of the process. For both flowsheets, the process type was ALL and the base method was UNIQUAC, which uses ideal gas and Henry’s law, best fitting the process conditions. 2.1. Reactor Separation Model (Configuration A). The kinetic data for the transesterification reaction was incorporated from a previous experimental study in our laboratory.11 The activation energy and the pre-exponential factor were calculated using linear regression of the experimental data reported. The kinetic parameters are shown in Table 2.
Table 2. Arrhenius Parameters for the Transesterification Reaction11 pre-exponential factor (A) (s−1) activation energy (Ea) (J/mol)
0.5574 30700
Because results in ref 11 indicated higher biodiesel yields at higher temperatures (150−200 °C), an isothermal operation with no pressure drop at 160 °C was chosen. Reaction temperatures higher than 150 °C do not significantly affect the biodiesel yield but increase the cooling water utility requirements. To obtain the optimized reactor temperature operation, reactor simulations were run in the temperature range of 150−200 °C and, at 160 °C, a maximum conversion of 99.81% was achieved. For sizing the reactor, the Aspen reactor model configuration provides length, diameter, and number of tubes as adjustable parameters. This is a continuous process; therefore, while the reactor was sized varying 6850
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Figure 6. Flow sheet units: reactor separation flow sheet (configuration A). column. This 6:1 optimal ratio was chosen for best conversion after a set of trial simulation runs of the CD column. A total holdup of 50 lb mol was imposed for the liquid phase. The distillate purity was above 99% mole fraction of methanol and could be recycled. The mixer functioned to merge the recycle into the fresh feed. The total conversion for triolein in the process is 99.63%, and the bottom stream from the RADFRAC column was sent to a decanter to settle and separate glycerol from methyl oleate. The operation of the CD column at lower pressures results in very little methanol in the bottom product (0.0044 mole fraction); this is a notable benefit of the CD operation. The bottom product from the CD column is composed mostly of glycerol and methyl oleate, which are easily separated by operating the decanter, owing to a significant difference between their densities (glycerol density of ∼1.26 g/cm3 and methyl oleate density of ∼0.87 g/cm3). The decanter temperature and pressure conditions were set close to the bottom stream from the RADFRAC column to minimize energy expenses while maintaining the purity standards. At an operation of 150 °C and 1 atm, its heat duty was the lowest for a separation of 99% methyl oleate. A final stream of 99.10% mole fraction purity for methyl oleate was obtained from the decanter as well as another stream with a purity of 99.38% for glycerol. These are very high product purity standards, and hence, additional distillation columns were not required. The composition and flow rate of all constituent streams for the reactor separation configuration are shown in Table 5.
Table 3. Reactor Configuration number of tubes reactor total length (m) single tube diameter (m) reactor total volume (m3)
100 10.50 0.37 112.9
and 99.56% of FAME on the bottom stream (99.71% mass fraction). The FAME distillation column operates at 0.3 atm (vacuum distillation) to reduce the temperature of the product FAME stream (biodiesel starts thermal degradation via isomerism, polymerization, and pyrolysis at temperatures exceeding 275 °C).18 Low-pressure distillation for biodiesel has also been reported in the literature.8 Methanol separation from biodiesel is utmost necessary to meet ASTM standards. Most biodiesel standard allows only 0.2% (v/v) methanol in the final product.19 Residual methanol in the biodiesel fuel is a major environmental and health hazard because of a number of reasons. Methanol is toxic (ingestion of 10 mL causes permanent blindness), has cold-start problems and lower energy density, and evaporates quickly when exposed to air. Excess methanol can also make the fuel flammable and more dangerous to handle and store besides, corroding metal components of the engine.13,19,20 The glycerol purification column achieved a separation of 99.87% mole fraction for methanol on the top stream (99.29% mass fraction), which is recycled back into the mixer, and a purity of 98.77% of glycerol on the bottom stream (99.53% mass fraction). The glycerol purification column operated by 0.5 atm to have a product glycerol stream less than its boiling point, 287.71 °C. This low-pressure distillation operation at 0.5 atm was also featured in the work performed in ref 8. The composition and flow rate of all constituent streams for the reactor separation configuration are shown in Table 4. 2.2. CD Model (Configuration B). One of the most significant merits that a CD brings to a process is simplification of the flow sheet and savings in equipment cost and operation. In this case, the biodiesel process is intensified by removal of the plug-flow reactor and two distillation columns from configuration A by a single CD column in configuration B, where both reaction and separation occur. For the CD configuration, the process equipment required was a mixer for enabling the recycle of methanol into the reactor, a reactive distillation column (RADFRAC) for the reaction and methanol separation, and a decanter for glycerol and methyl oleate separation. The elements are shown in Figure 7. The CD column was modeled in Aspen via an equilibrium-based rigorous two-phase fractionation model (RADFRAC) with a total number of seven stages (five trays, condenser, and reboiler) and a reflux ratio of 0.6. The column was operated at a pressure of 3 atm and a per stage pressure drop of 0.1 atm. This is another advantage demonstrated by a CD operation. Because the heat of the exothermic reaction is consumed to separate out the products, the column is able to operate at higher pressures while maintaining the reboiler product biodiesel at temperatures less than its degradation temperature (275 °C). Two separate feed streams were added to the column. The lighter component alcohol was fed at stage 6 close to the reboiler, and the heavier component oil was fed at stage 2 close to the condenser, to enable the reaction to take place between these stages. Design criteria were specified for the flow rate of the fresh alcohol feed to maintain the 6:1 methanol/triolein ratio for the stream exiting the mixer and entering the
3. PROCESS COMPARISONS (COST AND ENERGY) Both process configurations were optimized for desired biodiesel ASTM purity standards while minimizing energy requirements. The final streams from both biodiesel process configurations A and B yield the same mass flow rate (8900 lb mol h−1) and percentage purity (99.00 mol %) of biodiesel (methyl oleate). Operating conditions for both columns were set to have less than 0.2% (v/v) methanol in the final product to comply with ASTM standards. The aim of this research is to compare the total capital and operating costs between the two optimized process configurations for obtaining the same purity and flow of biodiesel to predict quantitatively the more cost-efficient process. To achieve this, the Aspen process economic analyzer tool (formerly, Aspen Icarus Process Evaluator) was used. The process economic analyzer tool is most valuable to compare competing technologies and/or evaluate alternative process configurations that are able to calculate the preliminary size for process equipment and generate operating and capital costs using built-in design and cost models directly from simulation data. Tables 6 and 7 enlists the total capital costs, total operating costs, equipment purchase costs, and yearly utility and raw material costs for configurations A and B, respectively, for an annual biodiesel production of around 10 million gallons/year. These values were evaluated using generated mass and energy results from Aspen process simulators, plugging in stream prices and then mapping and sizing the equipment using the process 6851
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6852
a
90.00 10.00 0.55 0.01 0.8949 0.0994 0.0055 0.0001 0.2445 0.7509 0.0043 0.0003 100.56 11792.31 491.68
30.09 10.00
0.7505 0.2495
0.0982 0.9018
40.09 9818.35 4.93
1.00
68.27 1.00 0.38 0.62
R-IN
25.00 1.00
Stream names are specified in Figure 2.
temperature (°C) pressure (atm) vapor fraction liquid fraction component mole flow (lb mol h−1) methanol triolein glycerol methyl oleate component mole fraction methanol triolein glycerol methyl oleate component mass fraction methanol triolein glycerol methyl oleate total mole flow (lb mol h−1) total mass flow (lb h−1) volume flow (m3 h−1)
FEED
0.1632 0.0014 0.0823 0.7532 100.56 11792.50 7.03
0.5972 0.0002 0.1048 0.2979
60.05 0.02 10.54 29.96
1.00
160.00 15.00
R-OUT
Table 4. Flow Sheet for the Reactor Separation: Configuration Aa
0.0148 0.0016 0.0923 0.8913 44.54 9961.42 5.62
0.1031 0.0004 0.2241 0.6724
4.59 0.02 9.98 29.95
1.00
125.00 1.00
FL-BOT
0.0564 4.09 × 10−11 0.9429 0.0007 11.59 965.88 0.38
0.1466 3.85 × 10−12 0.8532 0.0002
1.70 4.47 × 10−11 9.89 2.32 × 10−3
1.00
125.00 1.00
DEC-BOT
0.0103 0.0018 0.0010 0.9869 32.95 8995.55 5.10
0.0877 0.0006 0.0028 0.9089
2.89 0.02 0.09 29.94
1.00
125.00 1.00
DEC-TOP
0.0001 0.0018 0.0010 0.9971 30.08 8903.59 5.81
0.0007 0.0006 0.0031 0.9956
0.02 0.02 0.09 29.94
1.00
258.17 0.30
FAME
0.0043 1.81 × 10−21 0.9953 0.0005 10.01 914.96 0.37
0.0121 1.87 × 10−22 0.9877 0.0001
0.12 1.87 × 10−21 9.89 1.49 × 10−3
1.00
188.44 0.50
GLYCEROL
0.9706 1.65 × 10−5 0.0278 0.0016 56.03 1831.08 830.28
0.9900 6.08 × 10−7 0.0099 0.0002
55.47 3.41 × 10−5 0.55 0.01
125.00 1.00 1.00
RECYCLE1
1.0000 7.11 × 10−6 4.11 × 10−7 8.58 × 10−8 2.87 91.96 0.05
1.0000 2.57 × 10−7 1.43 × 10−7 9.27 × 10−9
2.87 7.38 × 10−7 4.11 × 10−7 2.66 × 10−8
1.00
36.80 0.30
RECYCLE2
0.9929 7.76 × 10−10 0.0022 0.0048 1.58 50.92 0.03
0.9987 2.83 × 10−11 0.0008 0.0005
1.58 4.47 × 10−11 1.24 × 10−3 8.31 × 10−4
1.00
47.94 0.50
RECYCLE3
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product stream price for technical-grade glycerol was set as 0.9 dollars/lb22 (the glycerol product stream in both configurations is more than 98.7 mol % in purity). The total capital corresponds to the investment required for purchase of equipment, costs of labor and materials (direct installation costs), costs for site preparation and buildings, and certain other costs (indirect installation costs). It also includes costs for land, working capital, and off-site facilities. The total operating costs include labor, maintenance, utilities, and raw material costs. Detailed information regarding parameters and evaluation basis for operating and capital costs can be found via generating the economics report file from the Aspen process economic analyzer toolbar. The implementation of heterogeneous catalyst systems for the biodiesel simulations in this study deserves special attention. Despite slower reaction kinetics and higher costs, they constitute an interesting area of study in the biodiesel process. Use of heterogeneous catalyst systems would significantly simplify the separation process and reduce equipment and utility costs.
Figure 7. Flow sheet units: CD flow sheet (configuration B).
economic analyzer tool. The feed stream prices were set as 26 cents/lb for soybean oil and 24 cents/lb for methanol.21 The Table 5. Flow Sheet for the CD: Configuration Ba stream
OIL
temperature (°C) pressure (atm) vapor fraction liquid fraction component mole flow (lb mol h−1) methanol triolein glycerol methyl oleate component mole fraction methanol triolein glycerol methyl oleate component mass fraction methanol triolein glycerol methyl oleate total mole flow (lb mol h−1) total mass flow (lb h−1) volume flow (m3 h−1) a
ALCOHOL
ALCOHOL2
DISBOT
DISTOP
FAME
GLYCEROL
25.00 3.20
25.00 1.00
62.00 3.70
273.37 3.70
95.43 3.00
150.00 1.00
150.00 1.00
1.00
1.00
1.00
1.00
1.00
1.00
1.00
30.07
60.00 2.88 × 10−3 0.06 1.33 × 10−3
0.18 0.04 9.96 29.89
29.93 2.88 × 10−3 0.06 1.33 × 10−3
0.12 0.04 0.12 29.89
0.06 4.57 × 10−11 9.85 1.70 × 10−3
1.00
0.9989 4.79 × 10−5 0.0010 2.21 × 10−5
0.0044 0.0009 0.2487 0.7460
0.9978 0.0001 0.0021 4.42 × 10−5
0.0039 0.0012 0.0039 0.9910
0.0060 4.61 × 10−12 0.9938 0.0002
1.00
0.9955 0.0013 0.0030 0.0002 60.07 1931.26 1.17
0.0006 0.0033 0.0935 0.9027 40.07 9817.91 6.51
0.9910 0.0026 0.0060 0.0004 30.00 967.85 0.62
0.0004 0.0036 0.0012 0.9947 30.16 8908.80 5.17
0.0021 4.45 × 10−11 0.9973 0.0006 9.91 909.10 0.35
10.00
1.00
1.00
10.00 8854.30 2.83
30.07 963.42 0.55
Stream names are specified in Figure 3.
Table 6. Detailed Cost Analysis for Optimum Design and Operating Conditions for the Reactor Separation Configuration (Configuration A) configuration A: reactor + distillation optimal specification
optimal operation condition
heat duty (kW)
volume (112.9 m3) volume (3.27 m3) volume (1.80 m3)
temperature (160 °C), pressure (15 atm) temperature (125 °C), pressure (1 atm) temperature (125 °C), pressure (1 atm)
−127.63544
FAME distillation column
stages (5), feed (3)
mole RR (0.5), pressure (0.3 atm)
glycerol distillation column
stages (3), feed (2)
mole RR (0.5), pressure (0.5 atm)
equipment reactor flash separator decanter
62.4769071 −14.899454 −20.044019 386.883111 −11.698744 28.4992635
utility cost ($/year) raw material cost ($/year)
equipment utility cost ($/year)
total operating cost ($/year)
equipment purchase cost ($)
9618.60
204800
4708.26
19200
1122.82
16100
30667.00 3029.32
25509398.00
70500
total capital cost ($)
5871380.00
52800
49145.00 22222700.00 6853
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Table 7. Detailed Cost Analysis for Optimum Design and Operating Conditions for the CD Configuration (Configuration B) configuration B: CD
optimal operation condition
heat duty (kW)
equipment utility cost ($/year)
equipment
optimal specification
decanter
volume (1.80 m3)
temperature (150 °C), pressure (1 atm)
−415.90
13423.32
RADFRAC
stages (7), feed (2, 6) (8.53 m high)
mole RR (0.6), pressure (3 atm), per stage pressure drop (0.1 atm)
−210.30 620.56
26815.98
utility cost ($/year) raw material cost ($/year)
total operating cost ($/year)
equipment purchase cost ($)
total capital cost ($)
13500.00 25001100.30
89700.00
3439170.00
40239.30 22207200.00
Table 8. Utility Usage and Costs Corresponding to Individual Process Equipment for the Reactor Separation Configuration (Configuration A) energy consumption/h (kW) mixer reactor flash separator decanter glycerol column FAME column total heating total cooling net energy requirement total utilities cost ($/year) total energy fraction price fraction ($/year)
−127.63544 62.4769071 −14.899454 −11.698744 28.4992635 −20.044019 386.883111 477.8592816 174.277657 652.1369386
heating fraction
cooling fraction
equipment heating cost ($/year)
0.73 0.13
equipment cooling cost ($/year)
equipment total utility cost ($/year)
9618.60 1122.82 881.62
9618.60 4708.26 1122.82 3029.32
1510.52
30667.00
4708.26 0.09 0.07
0.06
2147.70 0.12
0.81
29155.49
49145.00 heating 0.73 36011.45
cooling 0.27 13133.55
Figure 8. Liquid-phase concentrations of each component in the CD column (configuration B). The transesterification reaction (demonstrated by a decreasing concentration of triolein) mostly takes place on trays 1−3.
considerably benefit the cost of the overall process. Other merits include higher selectivity, water tolerance, and lifetime of the catalyst. In this work, we try to approximate the catalyst
Because glycerol is a valuable byproduct of the biodiesel production process, we believe that a relatively purer supply of glycerol from use of heterogeneous catalyst systems would 6854
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preparing the catalyst bed. Hence, the total catalyst costs were approximately 190 751 dollars for configuration A and 301 210 dollars for configuration B. These values are rough estimates for catalyst requirements and costs and will vary depending upon packing, bed characteristics, equipment geometry, temperature and pressure conditions, and flow rates of the process. The objective of calculating catalyst requirement in this research is to gauge an idea of the probable costs associated with changing the biodiesel production process to heterogeneous catalysis because no cost estimate for a heterogeneous catalyzed process for biodiesel was available in the literature. It is of note that the catalyst requirement for configuration B appears to be more than that for configuration A. Further investigation of composition and temperature profiles for the CD column (Figures 8 and 9) demonstrate that most of the reaction is taking place between trays 1 and 3. Hence, in actual running of the CD column, the required catalyst loading for the bottom trays should be less than the calculated value and both configurations would have relatively closer catalyst requirements. Results demonstrate that the CD configuration (B) is significantly more economical compared to the reactor separation configuration (A) in terms of capital and utility costs. There is only meagerly savings in terms of total operating and raw material costs per year. The total capital cost in dollars for the CD configuration (B) is 3.44 million dollars (41.42% less) compared to 5.87 million dollars for the reactor separation configuration (A). These capital costs are in agreement with reported capital costs for 10 million gallon annual production biodiesel plants that use soybean oil as feedstock using homogeneous catalysts.6,23,24 The total operating cost in dollars per year for the CD configuration (B) is 24.95 million dollars (1.46% less) compared to 25.32 million dollars for the reactor separation configuration (A). Numbers for operating costs for an annual 10 million gallon biodiesel soybean oil facility closely match reported literature.6,25 Because we are working at the same flow rates and achieving the same conversions, the raw material costs per year for both configurations are nearly identical. The utility costs per year for the CD configuration (B) are 18.12% lower compared to the reactor separation configuration (A). Aspen used built-in heat integration techniques (pinch technology) to minimize the utility costs that can be accessed using the energy analysis icon on the analysis toolbar. Utility usage and utility costs corresponding to each equipment in flow sheet configurations A and B are listed in Tables 8 and 9, respectively. The production cost per gallon of biodiesel is a significant factor to predict the profitability of the production process. From the annual production capacity and the total operating cost per year for the plant, the production cost per gallon of biodiesel was calculated. The calculations are shown in Table 10. The production cost per gallon of biodiesel for both configurations
Figure 9. Temperature profile (liquid phase) in the CD column (configuration B). A steeper temperature profile between trays 1 and 3 indicates the region of the exothermic transesterification reaction.
requirements and associated costs into the total annual operating costs of the whole process. It must be noted that our simulations are modeled on kinetic parameters taken from the performance of 20% calcium oxide supported on Al2O3 (20% calcium oxide by weight, with 80% alumina) as the solid base (discussed in section 211). Assuming a catalyst life of 1 year and catalyst loading of 3 wt % catalyst/oil ratio (maximum ester yield at this ratio11), costs associated with this heterogeneous catalyst system are added to the economic analysis results generated from the Aspen process economic analyzer tool. For configuration A, the residence times associated with the packed-bed reactor were used to approximate the triolein amount at any specific time inside the reactor and, accordingly, the catalyst requirement was calculated. For configuration B, the liquid holdup on each stage, feed composition, and flow rate were used to approximate the triolein amount and the corresponding catalyst requirements. Price specifications for alumina Al2O3 (Brockmann I, activated, 150 mesh size) and calcium nitrate terahydrate Ca(NO3)2·4H2O for synthesizing calcium oxide were taken from Sigma-Aldrich. The associated catalyst requirement for configuration A (the reaction + distillation process) is around 1907.51 kg/year, whereas for configuration B (CD process), it is 3012.0975 kg/year. As of August 2013, the Sigma-Aldrich price for alumina Al2O3 (Brockmann I, activated, 150 mesh size) is 387 dollars for 5 kg and the Sigma-Aldrich price for calcium nitrate terahydrate Ca(NO3)2·4H2O is 268 dollars for 2.5 kg. Using these prices, the cost of a catalyst system comprising 20% calcium oxide and 80% alumina by weight comes out to be 72.64 dollars/kg. Because batch and bulk costs for chemicals vary significantly, the catalyst was assumed to have an average cost of 100 dollars/kg. This price also allows for some compensation for the loss in material while synthesizing the catalyst and costs involved in
Table 9. Utility Usage and Costs Corresponding to Individual Process Equipment for the CD Configuration (Configuration B) energy consumption/h (kW) decanter distillation column total heating total cooling net energy requirement total utilities cost ($/year) total energy fraction price fraction ($/year)
−415.90298 −210.29701 620.559298 620.559298 −626.19999 1246.759288 40239.30 heating 0.50 20028.62
heating fraction
cooling fraction
equipment heating cost ($/year)
0.66 0.34 1.00
equipment cooling cost ($/year)
equipment total utility cost ($/year)
13423.32 6787.36
13423.32 26815.98
20028.62
cooling 0.50 20210.68 6855
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gratefully acknowledged. Mateus Lenz Leite acknowledges financial support for an exchange scholarship provided by the Brazilian National Council for Scientific and Technological Development (CNPq) through Science without Borders Program (CsF). Aashish Gaurav acknowledges a Trillium scholarship for Graduate Studies provided by the Ontario Government.
Table 10. Per Gallon Production Cost of Biodiesel for Configurations A and B reactor + distillation (configuration A)
CD (configuration B)
5871380.00 22222700.00
3439170.00 22207200.00
49145.00 190751 25509398.00 7214920
40,239.30 301210 25001100.30 7172290
8903.59 0.99 4027.089 870 40548.62 2.39
8908.80 0.99 4019.67 870 40473.92 2.36
1.71
1.67
total capital cost (USD) total raw materials cost (USD/year) (methanol + soybean oil) total utilities cost (USD/year) total catalyst cost total operating cost (USD/year) total product sales (USD/year) (glycerol) total “FAME stream” flow (lb/h) FAME mass fraction total FAME mass (kg/h) density (kg/m3) at 25 °C and 1 atm total volume (m3/year) price without considering glycerol revenue ($/gallon) price considering glycerol revenue ($/gallon)
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(1) Kreutzer, U. Manufacture of fatty alcohols based on natural fats and oils. J. Am. Oil Chem. Soc. 1984, 61 (2), 343−348. (2) Hill, J.; Nelson, E.; Tilman, D.; Polasky, S.; Tiffany, D. Environmental, economic, and energetic costs and benefits of biodiesel and ethanol biofuels. Proc. Natl. Acad. Sci. U. S. A. 2006, 103 (30), 11206−11210. (3) Ng, F. T. T.; Rempel, G. L. Catalytic distillation. Encyclopedia of Catalysis; John Wiley and Sons, Inc.: Hoboken, NJ, 2002. (4) Yokoyama, S. The Asian Biomass Handbook; Japan Institute of Energy: Tokyo, Japan, 2008. (5) Ma, F.; Hanna, M. A. Biodiesel production: A review. Bioresour. Technol. 1999, 70 (1), 1−15. (6) Haas, M. J.; McAloon, A. J.; Yee, W. C.; Foglia, T. A. A process model to estimate biodiesel production costs. Bioresour. Technol. 2006, 97 (4), 671−678. (7) Myint, L.; El-Halwagi, M. Process analysis and optimization of biodiesel production from soybean oil. Clean Technol. Environ. Policy 2009, 11 (3), 263−276. (8) Zhang, Y.; Dubé, M. A.; McLean, D. D.; Kates, M. Biodiesel production from waste cooking oil: 1. Process design and technological assessment. Bioresour. Technol. 2003, 89 (1), 1−16. (9) Zhang, Y.; Dubé, M. A.; McLean, D. D.; Kates, M. Biodiesel production from waste cooking oil: 2. Economic assessment and sensitivity analysis. Bioresour. Technol. 2003, 90 (3), 229−240. (10) You, Y.-D.; Shie, J.-L.; Chang, S.-Y.; Huang, S.-H.; Pai, C.-Y.; Yu, Y.-H. Economic cost analysis of biodiesel production: Case in soybean oil. Energy Fuels 2007, 22 (1), 182−189. (11) Pasupulety, N.; Gunda, K.; Liu, Y.; Rempel, G. L.; Ng, F. T. T. Production of biodiesel from soybean oil on CaO/Al2O3 solid base catalysts. Appl. Catal., A 2013, 452, 189−202. (12) Oliveira, C. F.; Dezaneti, L. M.; Garcia, F. A. C.; de Maced, J. L.; Dias, J. A.; Dias, S. C. L.; Alvim, K. S. P. Esterification of oleic acid with ethanol by 12-tungstophosphoric acid supported on zirconia. Appl. Catal., A 2010, 372 (2), 153−161. (13) Van Gerpen, J.; Shanks, B.; Pruszko, R.; Clements, D.; Knothe, G. Biodiesel Production Technology; National Renewable Energy Laboratory (NREL): Golden, CO, 2004; NREL/SR-510-36244. (14) Fogler, H. S. Elements of Chemical Reaction Engineering, 4th ed.; Prentice Hall: Upper Saddle River, NJ, 2006. (15) Baughman, W. F.; Jamieson, G. S. The chemical composition of soya bean oil. J. Am. Chem. Soc. 1922, 44 (12), 2947−2952. (16) Smith, W. B. The composition of soy-bean oil. J. Ind. Eng. Chem. 1922, 14 (6), 530−530. (17) Krawczyk, T. Biodiesel. Int. News Fats, Oils Relat. Mater. 1996, 7 (8), 800. (18) Lin, R.; Zhu, Y.; Tavlarides, L. L. Mechanism and kinetics of thermal decomposition of biodiesel fuel. Fuel 2013, 106, 593−604. (19) Bondioli, P. The Biodiesel Handbook. By Gerhard Knothe, Jon Van Gerpen and Jürgen Krahl (Eds.). Biotechnol. J. 2007, 2 (12), 1571− 1572. (20) Berrios, M.; Skelton, R. L. Comparison of purification methods for biodiesel. Chem. Eng. J. 2008, 144 (3), 459−465. (21) Methanex Corporation. Methanol PriceMethanex Regional Posted Contract Prices; http://www.methanex.com/products/ methanolprice.html (accessed July 7, 2013). (22) ICIS. ICIS Indicative Chemical Prices A−Z; http://www.icis.com/ chemicals/channel-info-chemicals-a-z/ (accessed June 17, 2013). (23) Van Gerpen, J. H. Biodiesel economics. Proceedings of the Oilseeds and Biodiesel Workshop; Billings, MT, Jan 9−10, 2008.
comes out to be around 2.3−2.4 dollars/gallon. After accounting for the glycerol sales from the product stream, the production cost comes out to be around 1.6−1.7 dollars/gallon for both of the configurations. Values for production costs are in agreement with ref 6 and 26 and several white papers published. It may be noted that, in the literature,6 published in 2005, a biodiesel production cost of around 2 dollars/gallon was predicted. Our predicted costs are very close, if the cumulative inflation factor between years 2005 and 2013 (around 15.06% as predicted by Statistics Canada) is considered between the raw material costs.
4. CONCLUSION The present research established the commercial feasibility of replacing the traditional reactor separation technology by the CD technology. The Aspen Plus process models developed can calculate optimal process conditions and process economics for biodiesel production. A comparison of conventional reactor separation and CD flow sheet simulation data predicts that CD can lead to significant savings in capital and utility costs. One of the hallmarks of this research is the investigative study of heterogeneous catalyst systems for the biodiesel process and an approximation of the associated catalyst requirements and corresponding costs. Another advantage of CD configuration was the possibility of high-pressure operation while maintaining low product stream temperatures. The production cost per gallon of biodiesel calculated is another suitable tool to predict the profitability of the operation. The process model developed in Aspen Plus is flexible to accommodate higher flow rates for scale-up of operations, add or remove stages of operation, and predict associated capital and production costs. One can also manipulate the feedstock and methanol price in the model to obtain different profitability and cost scenarios.
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REFERENCES
AUTHOR INFORMATION
Corresponding Author
*E-mail:
[email protected]. Notes
The authors declare no competing financial interest.
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ACKNOWLEDGMENTS Financial support from the Natural Sciences and Engineering Research Council of Canada (NSERC) for this research is 6856
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(24) Tyson, K. S., Biodiesel technology, economics and case studies. Proceedings of the NAEMI Biomass and Business Training Workshop; Spokane, WA, May 15−19, 2006. (25) Fortenbery, T. R. Biodiesel Feasibility Study: An Evaluation of Biodiesel Feasibility in Wisconsin; University of WisconsinMadison: Madison, WI, 2004. (26) Wisner, R. Biodiesel economicsCosts, tax credits and coproduct. AgMRC Renewable Energy Newsletter, June 2009; http://www. agmrc.org/renewable_energy/biodiesel/biodiesel-economics-coststax-credits-and-co-product (accessed July 7, 2013).
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