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Investigation to biodiesel production by the two-step homogeneous base-catalyzed transesterification
Bioresource Technology 101 (2010) 7368–7374
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Investigation to biodiesel production by the two-step homogeneous base-catalyzed transesterification Jianchu Ye, Song Tu, Yong Sha * College of Chemistry and Chemical Engineering, Xiamen University, Xiamen 361005, Fujian, China
a r t i c l e
i n f o
Article history: Received 27 January 2010 Received in revised form 24 March 2010 Accepted 30 March 2010 Available online 8 May 2010 Keywords: Biodiesel Simulation Two-step production Reaction kinetics Separation
a b s t r a c t For the two-step transesterification biodiesel production made from the sunflower oil, based on the kinetics model of the homogeneous base-catalyzed transesterification and the liquid–liquid phase equilibrium of the transesterification product, the total methanol/oil mole ratio, the total reaction time, and the split ratios of methanol and reaction time between the two reactors in the stage of the two-step reaction are determined quantitatively. In consideration of the transesterification intermediate product, both the traditional distillation separation process and the improved separation process of the two-step reaction product are investigated in detail by means of the rigorous process simulation. In comparison with the traditional distillation process, the improved separation process of the two-step reaction product has distinct advantage in the energy duty and equipment requirement due to replacement of the costly methanol–biodiesel distillation column. Ó 2010 Elsevier Ltd. All rights reserved.
1. Introduction The one-step biodiesel production process by means of the homogeneous base-catalyzed transesterification of the vegetable oil is widely utilized to produce biodiesel in the world. For the one-step biodiesel production process, the reversible transesterification takes place only once and is under the constraint of the reaction equilibrium, therefore the yield of biodiesel is about 91% (Cayhli and Kusefoglu, 2008) and the concentration of the transesterification intermediate product MG (monoglyceride) is about 2 wt% in biodiesel (Ma and Hanna, 1999; Bournay et al., 2005). For the two-step biodiesel production, the byproduct glycerol of the first transesterification is removed from the product with help of the immiscible liquid–liquid separation, and it can be favor of further conversion of raw materials. The following second transesterification can increase the yield of biodiesel to about 97% and reduce the composition of the intermediate product MG to 0.5 wt% (Goff et al., 2004; Bournay et al., 2005; Cayhli and Kusefoglu, 2008). There were some studies of the two-step biodiesel production about the influence of temperature, the reactor types and the overall methanol/oil ratio on the yield of biodiesel (Stiefel and Dassori, 2009; Mjalli et al., 2009), however, there are no work in relation to the process parameters of the two-step transesterification, such as the quantitative split of methanol and reaction time between the two transesterifications. For the industry application of the two-step biodiesel production, the optimization of these * Corresponding author. Tel.: +86 13606052265; fax: +86 592 2183054. E-mail address: [email protected] (Y. Sha). 0960-8524/$ - see front matter Ó 2010 Elsevier Ltd. All rights reserved. doi:10.1016/j.biortech.2010.03.148
technical parameters are important, and are worthy of a detailed investigation. The transesterification products, including biodiesel, glycerol and excessive methanol mainly, are separated from each other usually by means of distillation in the current industry processes (Fukuda et al., 2001; Van Gerpen, 2005; Dimian and Bildea, 2008). The technical information of some typical biodiesel distillation processes were investigated in detail by means of some process simulation softwares integrated with comprehensive thermodynamics package, reliable component libraries and advanced calculation techniques (Zhang et al., 2003; Haas et al., 2006; Zapata et al., 2007; West et al., 2008). However, those simulations focused on the separation process of the biodiesel production, and the transesterification kinetics and its influence on the composition were not considered. The intermediate products of transesterification, such as MG, were missing in the process simulation. Therefore, some technical information, such as the composition and the energy duty, were inconsistent with the actual process. For example, the reboiler temperature of the methanol– biodiesel distillation column in Zhang et al.’s work (2003) was up to 707.4 K because of absence of the intermediate products. Moreover, due to removal of glycerol between two transesterifications, there is a significant composition difference between the one-step reaction product and the two-step reaction product, and it may lead to different energy and equipment requirement of the subsequent separation process. It is necessary to carry out a detailed survey of the whole two-step reaction–separation process. The traditional distillation separation process of the biodiesel production has the high energy duty due to very high boiling
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points of methyl ester and glycerol. There exists the requirement to develop the new biodiesel separation operation with the low energy duty. According to investigation to the liquid–liquid phase equilibrium of the methanol–water–methyl ester ternary system (Stloukal et al., 1997), it is feasible to make use of water washing to extract methanol from the methyl ester phase into the water phase. Although glycerol exists in the methyl ester–methanol mixture, the content of glycerol is usually very small because of removal of most of glycerol before the second transesterification, and it has no significant effect on the liquid–liquid phase equilibrium of the methanol–methyl ester–water ternary system. Furthermore, this small amount of glycerol can also be removed by water washing because glycerol is apt to solve in water. Therefore, the water washing operation could be utilized to replace the traditional costly methyl ester–methanol distillation operation, and the corresponding improved separation process is worthy of further study. In this work, based on the kinetic model of the homogeneous base-catalyzed transesterification of the refined sunflower oil and the liquid–liquid phase equilibrium of the transesterification product, the two-step transesterification process is calculated in order to determine the suitable process parameters quantitatively. For the separation of the two-step transesterification product, the traditional distillation separation process and the improved separation process with the low energy duty are investigated by means of the rigorous process simulation. The methodology of this work is also suitable to the biodiesel production process using other raw material oil if its reaction kinetic is known. 2. The two-step homogeneous base-catalyzed transesterification process At present, the annual production capacity of most of biodiesel plants varies between 50,000 t and 100,000 t, so a biodiesel plant with the output of 80,000 t/a biodiesel made from the refined sunflower oil is considered in this work. The two-step transesterification process of this plant, including two PFR reactors and a decanter mainly, is shown in Fig. 1. Glycerol in the first transesterification product R1 is separated from biodiesel in the decanter 1 due to immiscibility between methyl ester and glycerol. The biodiesel phase P1, containing biodiesel, unreacted methanol and raw material oil, enters the second reactor to react together with additional methanol and catalyst NaOH. The glycerol phase S2 and the second transesterification product S1 are fed into the subsequent separation process for further separation. The refined sunflower oil can be regarded as the mixture of different TG (triglyceride) because chains of fatty acids in TG can be different. Most of chemical and physics properties of different TG are similar. Therefore, for the convenience of the process computa-
tion, according to the work of Zhang et al. (2003) and West et al. (2008), triolein is utilized to represent TG in spite of different types of TGs, and methyl oleate is utilized to represent biodiesel despite of the complicated composition of biodiesel. As a result, on the basis of the material balance, the mass flow rate of the refined sunflower oil is 11,018 kg/h if this plant runs 300 days in one year. For the process shown in Fig. 1, it can be calculated if the transesterification kinetics of the sunflower oil and the liquid–liquid equilibrium of the methanol–glycerol–methyl ester ternary system in the decanter are determined. 2.1. The calculation of the reactor and decanter k1
TG þ M DG þ ME
ð1Þ
k2
k3
DG þ M MG þ ME
ð2Þ
k4
k5
MG þ M GL þ ME
ð3Þ
k6
r 1 ¼
dC TG ¼ k1 C TG C M k2 C DG C ME dt
ð4Þ
r 2 ¼
dC DG ¼ k3 C DG C M k4 C MG C ME dt
ð5Þ
r 3 ¼
dC MG ¼ k5 C MG C M k6 C GL C ME dt
ð6Þ
According to the kinetics model of the homogeneous base-catalyzed transesterification of the sunflower oil (Freedman et al., 1986; Noureddini and Zhu, 1997; Darnoko and Cheryan, 2000; Kusdiana and Saka, 2001; Komers et al., 2002; Veljkovic et al., 2006; Bambase et al., 2007), there are three stepwise reversible reactions with formation of intermediate products DG (diglyceride) and MG. As shown in Formulas (1–3), 1 mol TG and 3 mol M (methanol) can result in the maximum production of 3 mol ME (methyl ester) and 1 mol GL (glycerol). The best kinetics model of this reversible transesterification appears to be a pseudo secondorder model expressed as Formulas (4–6), and the transient composition in the reactor can be obtained through calculating Formulas (4–6) by means of the Runge–Kutta method if the initial composition in the reactor is known. The reaction rate constants ki shown in Table 1 were determined by Bambase et al. (2007) under the common industrial conditions of 333 K, 1.0 wt% catalyst NaOH (by weight of oil), 6:1 methanol/oil mole ratio, and these kinetics parameters can be considered as constants if NaOH concentration is higher than 0.5 wt% (by weight of oil) or the methanol/oil molar ratio is higher than 6:1 (Bambase et al., 2007). According to these parameters and reaction conditions, the extent of transesterification can be valuated in the form of the biodiesel yield which is expressed as the percent of ME/ (3 TG) (based on the mole). Results from kinetics calculation show that the equilibrium yield of biodiesel is 96.62%, and the transesterification achieves almost 90% after 10 min and approaches the reaction equilibrium after 15 min. However, the calculated concentration of the intermediate product MG is as high
Table 1 The base-catalyzed transesterification rate constant ki of the refined sunflower oil (Bambase et al., 2007).
Fig. 1. Process flow diagram of the two-step transesterification process.
i
1
2
3
4
5
6
ki
0.2314
0.0166
0.4488
0.1068
0.877
0.06232
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as 2.067 wt% at the reaction equilibrium due to constraint of the reversible chemical equilibrium. It is much higher than the specification of the biodiesel standard, such as 0.8 wt% of MG in Germany Standard (DIN V 51606). The product of the transesterification reaction mainly includes glycerol, methanol and methyl ester. In the biodiesel production, glycerol is usually separated from methyl ester through a decanter because glycerol is basically immiscible with methyl ester. Some thermodynamics methods, such as UNIFAC–Dortmund method, can give the good prediction for the liquid–liquid phase separation of the glycerol–methanol–methyl ester ternary system (Chiu et al., 2005; Negi et al., 2006). However, a small quantity of MG in this ternary system can seriously affect the computational accuracy of these thermodynamics methods (Zhou et al., 2006). It is difficult to predict this liquid–liquid phase separation of the transesterification product theoretically since the calculated concentration of MG in the transesterification product is about 2.0 wt%. In order to avoid this difficulty, with respect to each component in the transesterification product, Stiefel and Dassori (2009) directly set the mass partition ratio of the component mass in the glycerol phase to that in the whole liquid–liquid system, and the mass partition ratios of TG, DG, MG, ME, methanol and glycerol was set to 0.001, 0.001, 0.001, 0.001, 0.4 and 0.94, respectively. This setting is adopted to calculate the liquid phase separation of the decanter between two reactors in this work.
2.2. The technical parameters of the two-step transesterification process For the two-step transesterification process shown in Fig. 1, some technical parameters, including operating temperature and pressure of the reactors, the reaction time in two reactors and the amount of methanol and NaOH in two reactors, need to be determined. The reaction pressure is set to 1 atm because the reaction occurs in the liquid phase and the pressure has little effect on the reaction rate (Dimian and Bildea, 2008). The operating temperature in two reactors is set at 333 K below the boiling point 337.8 K of methanol at 1 atm. In addition, most of the sunflower oil should be consumed in the first reactor, and the amount of the sunflower oil left to react in the second reactor is small, so only a small amount of NaOH are complemented to the second reaction. Therefore, the amount of NaOH added into the first reactor is set at 1.0 wt% (by weight of oil), and the 0.20 wt% (by weight of oil) NaOH is set to complement into the second reactor. It is known that the amount of NaOH has little effect on the reaction rate if the concentration of NaOH is higher than 0.5 wt% (Bambase et al., 2007), so this setting makes the concentration of NaOH in both reactors higher than 0.5 wt%, and this ensures that the kinetics parameter in Table 1 is applicable in the two-step process computation. Finally, there are only four process technical parameters left to be determined in Fig. 1, which are the overall mole ratio N of total methanol fed into two reactors to raw oil, the total reaction time t which is the sum of the reaction time in two reactors, and the split ratio x of methanol supplied in the first reactor to the total methanol, the split ratio y of the reaction time of the first reactor to the total reaction time. According to the current study, for the one-step biodiesel production, the appropriate methanol/oil mole ratio is 6:1 and the transesterification can achieve the equilibrium after 30 min. Therefore, N is considered to vary in the range of 5–9, t is considered to vary in the range of 20–40 min. Obviously, x and y should vary in the range of 0–1. The two-step transesterification process in Fig. 1 can be computed if N, t, x and y are known, and it is programmed by means of the software Matlab. According to the change of the biodiesel
yield in the variation ranges of N, t, x and y, a group of suitable process parameters, i.e. N, t, x and y, can be obtained. The effect of the total methanol/oil mole ratio N and the total reaction time t on the biodiesel yield is shown in Fig. 2, and data of the biodiesel yield in Fig. 2 are corresponding to the maximum value calculated in the 0–1 range of x and y at fixed N and t. It indicates apparently that increase of both N and t are in favor of enhancing the biodiesel yield, but the increase of methanol amount has little influence on the yield of biodiesel when N is greater than 7. Almost all of the yields of biodiesel can be beyond 99% in the whole range of 20–40 min if N is greater than 7. It should be noted that the increase of N can result in the more separation cost because the more excess methanol has to be recovered in the following separation process, and the extension of t also need the bigger reactor volume for the continuous production process. Therefore, N = 7 and t = 30 min are considered to be appropriate reaction process parameters according to the trade-off between the biodiesel yield and the process cost. Though the amount of methanol in the two-step reaction process is more than the one-step reaction process, the biodiesel yield of the two-step process at N = 7 and t = 30 min is 99.41% and almost all of material oil is converted into biodiesel. At N = 7 and t = 30 min, Fig. 3 indicates how the yield of biodiesel changes with the variation of the split ratios of the methanol and the reaction time between two reactors. The computational result shows that the maximum yield of biodiesel is 99.41 at x = 0.84 and y = 0.53, so in the two-step reaction process, the first reactor consumes more methanol than the second reactor while the respective reaction time of the two reactors is almost same. Moreover, as shown in Fig. 3, the change of x and y at the vicinity of the maximum value has no obvious impact on the yield of biodiesel, and it means that the yield of biodiesel is not sensitive to the split ratios of the methanol and the reaction time between two reactors in the two-step reaction process. 2.3. The stream information of the two-step transesterification process At N = 7, t = 30 min, x = 0.84 and y = 0.53, compositions of the main streams of the two-step reaction process in Fig. 1 are summarized in the Table 2. The concentration of the intermediate product MG in the two-step reaction product decreases significantly from 2.067 wt% to 0.3364 wt% in comparison with the one-step reaction. The yield of biodiesel is 99.41% after the second reaction, and it also has a significant increase in comparison with 96.62% of the
Fig. 2. Effect of N and t on the biodiesel yield.
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Fig. 3. Effect of x and y on the biodiesel yield.
one-step reaction. It should be noted that the yield calculated is a little greater than the experimental yield obtained by other researchers because the occurrence of the side reactions leading to the loss of raw oil is ignored in this work. Moreover, most of data reported about the yield of biodiesel were related to the yield of the whole reaction–separation process, however, the yield in this work is only related to the reaction process and the loss of biodiesel in the separation process is not considered. 3. The separation process of the two-step biodiesel production The two different separation processes is investigated rigorously by means of the process simulation software Aspen Plus, one for the traditional distillation separation process, another for the improved separation process. The difference between two separation processes locates in the separation method of methyl ester and methanol. As discussed in introduction of this work, the improved process utilizes the water washing operation to replace the costly methyl ester–methanol distillation operation, and its feasibility is investigated by the rigorous process simulation in this work. The feeds of the two separation processes, S1 and S2, come from the two-step reaction process shown in Fig. 1, and its information is listed in Table 2, so the intermediate products of the transesterification, i.e. MG and DG, can emerge in the separation process. The composition of the biodiesel product in both separation processes is specified to meet the Germany specification for biodiesel (DIN V 51606) as shown in Table 3 (Dimian and Bildea, 2008). The information of the separation process simulation together with the simulation information of the two-step reaction process can provide the whole process information of the two-step biodiesel production. 3.1. The setup of the process simulation The main equipments in the separation process of the biodiesel production include the liquid–liquid gravity settlers, distillation
Components
Unit
Min.
Max.
Water MG/DG/TG Free glycerol Methanol
wt% wt% wt% wt%
– – – –
0.03 0.8/0.4/0.4 0.02 0.3
columns, extraction columns, flash drums and heat exchangers. The gravity settler and flash drum are set to work under the adiabatic condition. The sieve-tray column is chosen to carry out distillation and extraction operation in the separation process. Due to the high viscosity of both biodiesel and glycerol, in the distillation operation, the sieve-tray efficiency is set to 0.6, and the pressure drop of the single tray is set to 0.7 kPa. The tray efficiency of the extraction column is set to be 0.8. The boiling points of both biodiesel and glycerol are very high, and their decomposition temperature is lower than the boiling point, so the vacuum distillation is adopted to recover methanol from the biodiesel phase and glycerol phase. The pressure of methanol–biodiesel distillation column is set to 11.65 kPa in order to prevent decomposition of biodiesel. But for the methanol–glycerol distillation column, it is difficult to prevent the decomposition of glycerol only by the vacuum operation. In the biodiesel industry, a water steam is usually introduced into the methanol–glycerol distillation column to reduce the bottom temperature, and the purification of glycerol is still a very tough problem so far. Because the study on purification of glycerol is not a key point of this work, similar to the work of Zhang et al. (2003), the recovery of methanol from the glycerol phase is conducted only in a simple distillation column under the atmosphere pressure. The concentration of methanol recovered from biodiesel and glycerol in the recycling methanol stream must meet the requirement of the transesterification because the impurities, especially water, have serious influence on the transesterification. In this work the concentration of methanol in the recycling methanol streams is specified to be higher than 99.8 wt%. For both of the methanol–glycerol distillation column and the methanol–biodiesel distillation column, the tray number and reflux ratio are adjusted to meet to the requirement of the methanol concentration, and the tray numbers of 9–10 and the reflux ratio of 1–2 can achieve the separation target due to the big difference of boiling points between methanol and biodiesel/glycerol. In the separation process of the biodiesel production, except heat exchangers, flash drums and distillation columns, other equipments usually work under the normal temperature and pressure, and there is no corrosive and harmful material except NaOH. Pumps and pipes mainly change the pressure of streams, so they are absent in the process simulation. The equipment investment of the heat exchanger and flash drum is much lower than that of the distillation column, so it is supposed that the equipment cost in the separation process is mainly composed of the distillation column and its accessory equipments. The energy consumption of the separation process also mainly locates in the distillation operation due to high boiling points of methyl ester and glycerol. The catalyst NaOH is not taken into account in the simulation, and there are two reasons for neglect of NaOH. One reason is that
Table 2 Information of the main streams in the two-step transesterification process. Stream name
TG (wt%)
DG (wt%)
MG (wt%)
ME (wt%)
M (wt%)
G (wt%)
Mass rate (kg/h)
R1 P1 S2 S1
0.0825 0.0928 <0.01 0.0166
0.8075 0.9082 <0.01 0.1499
2.065 2.322 0.0185 0.3364
79.94 89.92 0.7185 89.36
9.213 6.223 32.95 9.009
7.888 0.5328 66.30 1.125
13,399 11,900 1498.2 12,299
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the presence of the little catalyst does not affect the thermodynamic equilibrium, and another reason is that in the industrial process the free glycerol in biodiesel is removed in a water washing column where NaOH can be also removed easily by means of the sour water washing at the same time. The water washing column still exists in the simulation even if the catalyst NaOH is ignored, so the reliability of the comparison between two different separation processes can be assured. In addition, the generation of soap in the biodiesel production is ignored because there is little water and free fatty acids in the refined sunflower oil. The neglect of NaOH and soap in the separation process just decrease the consumption of washing water, and it is little influence on other results of the process simulation. For some compounds emerging in the separation process, such as DG and MG, they are not contained in the database of the process simulation software Aspen Plus, therefore their properties are calculated in the simulation by means of the group contribution method according to their molecular structure. For the methanol–glycerol–methyl ester system, the computational accuracy of the UNIFAC–Dortmund thermodynamics method can meet the requirement of the simulation very well because the concentration of the intermediate product MG is very low after the second transesterification. For the liquid–liquid equilibrium of the methanol–water–methyl ester ternary system, the computational results by means of the UNIFAC–Dortmund method are also in good accordance with experimental results of Stloukal et al. (1997), and the maximum deviation between experimental and computational data of the composition in the water phase and methyl ester phase is no more than 5%. Therefore, the UNIFAC–Dortmund method is utilized to calculate the thermodynamic equilibrium in the simulation of two separation processes. 3.2. The traditional distillation separation process According to works of Van Gerpen (2005) and Fukuda et al. (2001), the process flow diagram of the traditional distillation sep-
aration is presented in Fig. 4. The second transesterification product S1 enters the gravity settler D-2. By means of the liquid– liquid phase separation, S1 are separated to the methyl ester phase MEP and glycerol phase GLP in D-2. The upper methyl ester phase MEP of D-2 is fed into the methanol–biodiesel distillation column T-2. The overhead product of T-2 is methanol and it is a part of the recycling methanol S5. The bottom product S7 of T-2 is the crude biodiesel containing impurities such as free glycerol. The glycerol phase S2 and GLP of the first and second transesterification products are fed into the methanol–glycerol distillation column T-1. The overhead product of T-1 is another part of the recycling methanol S5, and the bottom product of T-1 is crude glycerol. The impurity in crude biodiesel stream S7 is removed by 55 kg/h washing water W in the washing column E-1, and a very small amount of water in S6 is removed from the top of the flash drum F-1 after S6 is heated up to 473.1 K in the heat exchanger H-2. Finally, the biodiesel product S3 which meets the Germany specification for biodiesel (DIN V 51606) can be obtained after being cooled in the heat exchanger H-4. The composition and mass rate of the main streams involved in Fig. 4 is summarized in Table 4 and the technical information of the separation process is shown in Table 6. 3.3. The improved separation process As shown in Fig. 5, in the improved separation process, the second transesterification product S1 is fed into the water washing column E-1 at first, and is washed by 55 kg/h water WW. E-1 has four theoretical trays, and can remove most of methanol and all of free glycerol. After water washing, the concentration of methanol in the crude biodiesel stream S7 is 1.751 wt% in comparison with initial 9.009 wt% in S1. The small amount of methanol and water in S7 can be removed easily by the adiabatic flash drum F-1 because their boiling points are much lower than that of biodiesel. To order to improve the recovery of methanol, the vapor product S6 of F-1 is cooled, then mixed with the bottom product S8 of E-1 and the glycerol phase S2, and finally fed into the
Fig. 4. Process flow diagram of the traditional distillation separation process.
Table 4 Information of the main streams in Fig. 4. Stream name
TG (wt%)
DG (wt%)
MG (wt%)
ME (wt%)
M (wt%)
G (wt%)
W (wt%)
P (kPa)
T (K)
Mass rate (kg/h)
S3 S4 S5 S6 S7
0.0184 <0.01 <0.01 0.0184 0.0184
0.1673 <0.01 <0.01 0.1663 0.1667
0.3711 0.0647 <0.01 0.3691 0.3700
99.41 1.061 0.1059 99.13 99.38
<0.01 0.8565 99.81 0.0212 0.0239
<0.01 98.02 0.0834 <0.01 0.03894
0.0280 0 0 0.2944 0
101.3 101.3 101.3 101.3 101.3
298.1 298.1 313.2 331.6 522.8
11,022 1148.9 1595.3 11,084 11,055
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Fig. 5. Process flow diagram of the improved separation process.
Table 5 Information of the main streams in Fig. 5. Stream name
TG (wt%)
DG (wt%)
MG (wt%)
ME (wt%)
M (wt%)
G (wt%)
W (wt%)
P (kPa)
T (K)
Mass rate (kg/h)
S3 S4 S5 S6 S7 S8 S9
0.0184 <0.01 <0.01 <0.01 0.0181 <0.01 <0.01
0.1672 <0.01 <0.01 <0.01 0.1638 <0.01 <0.01
0.3199 0.5357 <0.01 <0.01 0.3135 0.5490 0.5132
99.41 2.925 <0.01 4.331 97.52 1.763 3.222
0.0768 0.2049 99.81 84.26 1.751 81.98 0.4904
<0.01 94.90 <0.01 <0.01 <0.01 12.45 91.05
0.0108 1.431 0.1871 11.40 0.2373 3.260 4.727
101.3 101.3 101.3 20.26 101.3 101.3 107.6
298.1 298.1 351.3 423.1 336.6 335.7 459.9
11,027 1190.6 1590.2 223.72 11,250 1111.4 1243.1
Table 6 Technical information of the two separation processes. Items Heat exchanger
H-1 H-2 H-3 H-4
Separation unit
T-1
T-2
E-1 E-2 F-1 F-2
Stream in/out temperature Energy duty (MW) Stream in/out temperature Energy duty (MW) Stream in/out temperature Energy duty (MW) Stream in/out temperature Energy duty (MW) Reboiler temperature (K) Pressure (kPa) Reboiler duty (MW) Condenser duty (MW) Diameter (m) Tray number/reflux ratio Reboiler temperature (K) Pressure (kPa) Reboiler duty (MW) Condenser duty (MW) Diameter (m) Tray number/reflux ratio Theoretical stages Temperature (K) Theoretical stages Temperature (K) Temperature (K) Pressure (kPa) Temperature (K) Pressure (kPa)
(K) (K) (K) (K)
Fig. 4
Fig. 5
528.9/333.1 1.414 331.6/473.1 0.9883 557.5/298.1 0.2226 473.1/298.1 1.158
336.0/423.1 0.6357 423.1/298.1 0.7873 431.7/298.1 0.03410 431.7/298.1 0.1099
557.5 101.3 0.6588 0.4564 0.5 10/1 528.9 11.65 1.900 0.5560 1.0 10/1 5 333.1 –
459.9 101.3 1.480 1.418 0.9 10/2 –
473.1 20.26 –
423.1 20.26 459.9 101.3
4 333.1 –
Separation effect
Recovery efficiency of methanol (%) Recovery efficiency of ME (%) Biodiesel output (kg/h)
99.22 99.70 11,022
99.09 99.73 11,027
Number of the main separation equipments
Distillation column Washing column Flash evaporator Heat exchanger
2 1 1 4
1 1 2 4
Energy duty
The total heat duty (MW) The total cold duty (MW)
3.547 3.534
2.226 2.239
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glycerol–methanol distillation column T-1. The overhead product of T-1, the methanol stream S5, is recycled to the reaction process. The bottom product of T-1 contains the glycerol and water, and water is removed through the flash operation so that crude glycerol is obtained. The information of streams and equipments in the improved separation process are shown in Tables 5 and 6 . The compositions of biodiesel produced by the improved separation process can meet the Germany specification for biodiesel very well, so the improved separation process can achieve the separation target of the biodiesel production. Although it could be difficult to conduct the liquid–liquid extraction operation in water washing column E-1 due to the small density difference between the heavy and light phase (Tizvar et al., 2009), the centrifugal extraction operation can be utilized to solve this problem. 3.4. The comparison between the traditional and improved separation process The parameters of main equipments, energy duty and separation effect of the traditional and improved separation processes are summarized in the Table 6. In order to compare the separation effect of the two processes, the recovery efficiency of methanol and biodiesel are defined as following formulas:
The recovery efficiency of methanol ¼
F S5 wM;S5 100% F S1 wM;S1 þ F S2 wM;S2
The recovery efficiency of methyl ester ¼
F S3 wME;S3 100% F S1 wME;S1
where, F is the mass flow rate of streams, wM and wME are the mass fraction of methanol and methyl ester in the corresponding stream. As shown in Table 6, the two investigated separation processes have similar separation effect, and both can meet the specification of biodiesel. For the traditional distillation process, the equipment investment and energy duty are high because of the presence of two distillation columns in the process and high boiling points of biodiesel and glycerol. However, the improved separation process has significant advantage. There is only one distillation column in the improved separation process, and its energy duty decreases by about 35% in contrast with the traditional distillation process. The results from simulation show that it is feasible for the improved separation process in the further application. 4. Conclusion By means of the rigorous process simulation, the two-step biodiesel production process made from the refined sunflower oil is investigated in detail. The total methanol/oil mole ratio, the total reaction time and the split ratios of both methanol and reaction time between two reactors are determined quantitatively. The technical information about the traditional distillation separation process and the improved separation process are acquired and
compared. The improved separation process, where the water washing operation is utilized to replace the costly methyl ester– methanol distillation operation, is verified to be of advantage in reducing energy consumption and equipment investment. References Bournay, L., Casanave, D., Delfort, B., Hillion, G., Chodorge, J.A., 2005. New heterogeneous process for biodiesel production: a way to improve the quality and the value of the crude glycerin produced by biodiesel plants. Catal. Today 106 (1), 190–192. Bambase, M.E., Nakamura, N., Tanaka, J., Matsumura, M., 2007. Kinetics of hydroxide-catalyzed methanolysis of crude sunflower oil for the production of fuel-grade methyl esters. J. Chem. Technol. Biotechnol. 82 (3), 273–280. Cayhli, G., Kusefoglu, S., 2008. Increased yields in biodiesel production from used cooking oils by a two step process: comparison with one step process by using TGA. Fuel Proc. Tech. 89 (2), 118–122. Chiu, C.W., Goff, M.J., Suppes, G.J., 2005. Distribution of methanol and catalysts between biodiesel and glycerin phases. AIChE J. 51 (4), 1274–1278. Dimian, A.C., Bildea, C., 2008. Chemical Process Design: Computer-Aided Case Studies. Wiley-VCH, New York. Darnoko, D., Cheryan, M., 2000. Kinetics of palm oil transesterification in a batch reactor. JAOCS 77 (12), 1263–1267. Fukuda, H., Kondo, A., Noda, H., 2001. Biodiesel fuel production by transesterification of oils. J. Biosci. Bioeng. 92 (5), 405–416. Freedman, B., Butterfield, R.O., Pryde, E.H., 1986. Transesterification kinetics of soybean oil. JAOCS 63 (10), 1375–1380. Goff, M.J., Bauer, N.S., Lopes, S., Sutterlin, W.R., Suppes, G.J., 2004. Acid-catalyzed alcoholysis of soybean oil. JAOCS 81 (4), 415–420. Haas, M.J., McAloon, A.J., Yee, W.C., Foglia, T.A., 2006. A process model to estimate biodiesel production costs. Bioresource Technol. 97 (4), 671–678. Kusdiana, D., Saka, S., 2001. Kinetics of transesterification in rapeseed oil to biodiesel fuel as treated in supercritical methanol. Fuel 80 (5), 693–698. Komers, K., Skopal, F., Stloukal, R., Machek, J., 2002. Kinetics and mechanism of the KOH-catalyzed methanolysis of rapeseed oil for biodiesel production. Eur. J. Lipid Sci. Technol. 104 (11), 718–737. Mjalli, F.S., San, L.K., Yin, K.C., 2009. Dynamics and control of a biodiesel transesterification reactor. Chem. Eng. Tech. 32 (1), 13–26. Ma, F.R., Hanna, M., 1999. A biodiesel production: a review. Bioresource Technol 70 (1–4), 1–15. Noureddini, H., Zhu, D., 1997. Kinetics of transesterification of soybean oil. JAOCS 74 (11), 1457–1463. Negi, D.S., Sobotka, F., Kimmel, T., 2006. Liquid–liquid phase equilibrium in glycerol–methanol–methyl oleate and glycerol–monoolein–methyl oleate ternary systems. Ind. Eng. Chem. Res. 45 (10), 3693–3696. Stiefel, S., Dassori, G., 2009. Simulation of biodiesel production through transesterification of vegetable oil. Ind. Eng. Chem. Res. 48 (3), 1068– 1071. Stloukal, R., Komers, K., Machek, J., 1997. Ternary phase diagram biodiesel fuel– methanol–water. J. Prakt. Chem. 339 (5), 485–487. Tizvar, R., McLean, D.D., Kates, M., Dube, M.A., 2009. Optimal separation of glycerol and methyl oleate via liquid–liquid extraction. J. Chem. Eng. Data 54 (5), 1541– 1550. Van Gerpen, J., 2005. Biodiesel processing and production. Fuel Process. Technol. 86 (10), 1097–1107. Veljkovic, V.B., Lakicevic, S.H., Stamenkovic, O.S., Todorovic, Z.B., Lazic, M.L., 2006. Biodiesel production from tobacco (Nicotiana tabacum L.) seed oil with a high content of free fatty acids. Fuel 85 (17–18), 2671–2675. West, A.H., Posarac, D., Ellis, N., 2008. Assessment of four biodiesel production processes using HYSYS. Plant. Bioresource Technol. 99 (14), 6587–6601. Zapata, C.D., Martinez, L.D., Castiblanco, E.A., Uribe, C.A.H., 2007. Biodiesel production from crude palm oil: 1. Design and simulation of two continuous processes. Dyna – Colombia 74 (151), 71–82. Zhang, Y., Dube, M.A., Mclean, D.D., Kates, M., 2003. Biodiesel production from waste cooking oil: 1. process design and technological assessment. Bioresource Technol. 89 (1), 1–6. Zhou, H., Lu, H.F., Liang, B., 2006. Solubility of multicomponent systems in the biodiesel production by transesterification of Jatropha curcas L. oil with methanol. J. Chem. Eng. Data 51 (3), 1130–1135.
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Bioresource Technology journal homepage: www.elsevier.com/locate/biortech
Investigation to biodiesel production by the two-step homogeneous base-catalyzed transesterification Jianchu Ye, Song Tu, Yong Sha * College of Chemistry and Chemical Engineering, Xiamen University, Xiamen 361005, Fujian, China
a r t i c l e
i n f o
Article history: Received 27 January 2010 Received in revised form 24 March 2010 Accepted 30 March 2010 Available online 8 May 2010 Keywords: Biodiesel Simulation Two-step production Reaction kinetics Separation
a b s t r a c t For the two-step transesterification biodiesel production made from the sunflower oil, based on the kinetics model of the homogeneous base-catalyzed transesterification and the liquid–liquid phase equilibrium of the transesterification product, the total methanol/oil mole ratio, the total reaction time, and the split ratios of methanol and reaction time between the two reactors in the stage of the two-step reaction are determined quantitatively. In consideration of the transesterification intermediate product, both the traditional distillation separation process and the improved separation process of the two-step reaction product are investigated in detail by means of the rigorous process simulation. In comparison with the traditional distillation process, the improved separation process of the two-step reaction product has distinct advantage in the energy duty and equipment requirement due to replacement of the costly methanol–biodiesel distillation column. Ó 2010 Elsevier Ltd. All rights reserved.
1. Introduction The one-step biodiesel production process by means of the homogeneous base-catalyzed transesterification of the vegetable oil is widely utilized to produce biodiesel in the world. For the one-step biodiesel production process, the reversible transesterification takes place only once and is under the constraint of the reaction equilibrium, therefore the yield of biodiesel is about 91% (Cayhli and Kusefoglu, 2008) and the concentration of the transesterification intermediate product MG (monoglyceride) is about 2 wt% in biodiesel (Ma and Hanna, 1999; Bournay et al., 2005). For the two-step biodiesel production, the byproduct glycerol of the first transesterification is removed from the product with help of the immiscible liquid–liquid separation, and it can be favor of further conversion of raw materials. The following second transesterification can increase the yield of biodiesel to about 97% and reduce the composition of the intermediate product MG to 0.5 wt% (Goff et al., 2004; Bournay et al., 2005; Cayhli and Kusefoglu, 2008). There were some studies of the two-step biodiesel production about the influence of temperature, the reactor types and the overall methanol/oil ratio on the yield of biodiesel (Stiefel and Dassori, 2009; Mjalli et al., 2009), however, there are no work in relation to the process parameters of the two-step transesterification, such as the quantitative split of methanol and reaction time between the two transesterifications. For the industry application of the two-step biodiesel production, the optimization of these * Corresponding author. Tel.: +86 13606052265; fax: +86 592 2183054. E-mail address: [email protected] (Y. Sha). 0960-8524/$ - see front matter Ó 2010 Elsevier Ltd. All rights reserved. doi:10.1016/j.biortech.2010.03.148
technical parameters are important, and are worthy of a detailed investigation. The transesterification products, including biodiesel, glycerol and excessive methanol mainly, are separated from each other usually by means of distillation in the current industry processes (Fukuda et al., 2001; Van Gerpen, 2005; Dimian and Bildea, 2008). The technical information of some typical biodiesel distillation processes were investigated in detail by means of some process simulation softwares integrated with comprehensive thermodynamics package, reliable component libraries and advanced calculation techniques (Zhang et al., 2003; Haas et al., 2006; Zapata et al., 2007; West et al., 2008). However, those simulations focused on the separation process of the biodiesel production, and the transesterification kinetics and its influence on the composition were not considered. The intermediate products of transesterification, such as MG, were missing in the process simulation. Therefore, some technical information, such as the composition and the energy duty, were inconsistent with the actual process. For example, the reboiler temperature of the methanol– biodiesel distillation column in Zhang et al.’s work (2003) was up to 707.4 K because of absence of the intermediate products. Moreover, due to removal of glycerol between two transesterifications, there is a significant composition difference between the one-step reaction product and the two-step reaction product, and it may lead to different energy and equipment requirement of the subsequent separation process. It is necessary to carry out a detailed survey of the whole two-step reaction–separation process. The traditional distillation separation process of the biodiesel production has the high energy duty due to very high boiling
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points of methyl ester and glycerol. There exists the requirement to develop the new biodiesel separation operation with the low energy duty. According to investigation to the liquid–liquid phase equilibrium of the methanol–water–methyl ester ternary system (Stloukal et al., 1997), it is feasible to make use of water washing to extract methanol from the methyl ester phase into the water phase. Although glycerol exists in the methyl ester–methanol mixture, the content of glycerol is usually very small because of removal of most of glycerol before the second transesterification, and it has no significant effect on the liquid–liquid phase equilibrium of the methanol–methyl ester–water ternary system. Furthermore, this small amount of glycerol can also be removed by water washing because glycerol is apt to solve in water. Therefore, the water washing operation could be utilized to replace the traditional costly methyl ester–methanol distillation operation, and the corresponding improved separation process is worthy of further study. In this work, based on the kinetic model of the homogeneous base-catalyzed transesterification of the refined sunflower oil and the liquid–liquid phase equilibrium of the transesterification product, the two-step transesterification process is calculated in order to determine the suitable process parameters quantitatively. For the separation of the two-step transesterification product, the traditional distillation separation process and the improved separation process with the low energy duty are investigated by means of the rigorous process simulation. The methodology of this work is also suitable to the biodiesel production process using other raw material oil if its reaction kinetic is known. 2. The two-step homogeneous base-catalyzed transesterification process At present, the annual production capacity of most of biodiesel plants varies between 50,000 t and 100,000 t, so a biodiesel plant with the output of 80,000 t/a biodiesel made from the refined sunflower oil is considered in this work. The two-step transesterification process of this plant, including two PFR reactors and a decanter mainly, is shown in Fig. 1. Glycerol in the first transesterification product R1 is separated from biodiesel in the decanter 1 due to immiscibility between methyl ester and glycerol. The biodiesel phase P1, containing biodiesel, unreacted methanol and raw material oil, enters the second reactor to react together with additional methanol and catalyst NaOH. The glycerol phase S2 and the second transesterification product S1 are fed into the subsequent separation process for further separation. The refined sunflower oil can be regarded as the mixture of different TG (triglyceride) because chains of fatty acids in TG can be different. Most of chemical and physics properties of different TG are similar. Therefore, for the convenience of the process computa-
tion, according to the work of Zhang et al. (2003) and West et al. (2008), triolein is utilized to represent TG in spite of different types of TGs, and methyl oleate is utilized to represent biodiesel despite of the complicated composition of biodiesel. As a result, on the basis of the material balance, the mass flow rate of the refined sunflower oil is 11,018 kg/h if this plant runs 300 days in one year. For the process shown in Fig. 1, it can be calculated if the transesterification kinetics of the sunflower oil and the liquid–liquid equilibrium of the methanol–glycerol–methyl ester ternary system in the decanter are determined. 2.1. The calculation of the reactor and decanter k1
TG þ M DG þ ME
ð1Þ
k2
k3
DG þ M MG þ ME
ð2Þ
k4
k5
MG þ M GL þ ME
ð3Þ
k6
r 1 ¼
dC TG ¼ k1 C TG C M k2 C DG C ME dt
ð4Þ
r 2 ¼
dC DG ¼ k3 C DG C M k4 C MG C ME dt
ð5Þ
r 3 ¼
dC MG ¼ k5 C MG C M k6 C GL C ME dt
ð6Þ
According to the kinetics model of the homogeneous base-catalyzed transesterification of the sunflower oil (Freedman et al., 1986; Noureddini and Zhu, 1997; Darnoko and Cheryan, 2000; Kusdiana and Saka, 2001; Komers et al., 2002; Veljkovic et al., 2006; Bambase et al., 2007), there are three stepwise reversible reactions with formation of intermediate products DG (diglyceride) and MG. As shown in Formulas (1–3), 1 mol TG and 3 mol M (methanol) can result in the maximum production of 3 mol ME (methyl ester) and 1 mol GL (glycerol). The best kinetics model of this reversible transesterification appears to be a pseudo secondorder model expressed as Formulas (4–6), and the transient composition in the reactor can be obtained through calculating Formulas (4–6) by means of the Runge–Kutta method if the initial composition in the reactor is known. The reaction rate constants ki shown in Table 1 were determined by Bambase et al. (2007) under the common industrial conditions of 333 K, 1.0 wt% catalyst NaOH (by weight of oil), 6:1 methanol/oil mole ratio, and these kinetics parameters can be considered as constants if NaOH concentration is higher than 0.5 wt% (by weight of oil) or the methanol/oil molar ratio is higher than 6:1 (Bambase et al., 2007). According to these parameters and reaction conditions, the extent of transesterification can be valuated in the form of the biodiesel yield which is expressed as the percent of ME/ (3 TG) (based on the mole). Results from kinetics calculation show that the equilibrium yield of biodiesel is 96.62%, and the transesterification achieves almost 90% after 10 min and approaches the reaction equilibrium after 15 min. However, the calculated concentration of the intermediate product MG is as high
Table 1 The base-catalyzed transesterification rate constant ki of the refined sunflower oil (Bambase et al., 2007).
Fig. 1. Process flow diagram of the two-step transesterification process.
i
1
2
3
4
5
6
ki
0.2314
0.0166
0.4488
0.1068
0.877
0.06232
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as 2.067 wt% at the reaction equilibrium due to constraint of the reversible chemical equilibrium. It is much higher than the specification of the biodiesel standard, such as 0.8 wt% of MG in Germany Standard (DIN V 51606). The product of the transesterification reaction mainly includes glycerol, methanol and methyl ester. In the biodiesel production, glycerol is usually separated from methyl ester through a decanter because glycerol is basically immiscible with methyl ester. Some thermodynamics methods, such as UNIFAC–Dortmund method, can give the good prediction for the liquid–liquid phase separation of the glycerol–methanol–methyl ester ternary system (Chiu et al., 2005; Negi et al., 2006). However, a small quantity of MG in this ternary system can seriously affect the computational accuracy of these thermodynamics methods (Zhou et al., 2006). It is difficult to predict this liquid–liquid phase separation of the transesterification product theoretically since the calculated concentration of MG in the transesterification product is about 2.0 wt%. In order to avoid this difficulty, with respect to each component in the transesterification product, Stiefel and Dassori (2009) directly set the mass partition ratio of the component mass in the glycerol phase to that in the whole liquid–liquid system, and the mass partition ratios of TG, DG, MG, ME, methanol and glycerol was set to 0.001, 0.001, 0.001, 0.001, 0.4 and 0.94, respectively. This setting is adopted to calculate the liquid phase separation of the decanter between two reactors in this work.
2.2. The technical parameters of the two-step transesterification process For the two-step transesterification process shown in Fig. 1, some technical parameters, including operating temperature and pressure of the reactors, the reaction time in two reactors and the amount of methanol and NaOH in two reactors, need to be determined. The reaction pressure is set to 1 atm because the reaction occurs in the liquid phase and the pressure has little effect on the reaction rate (Dimian and Bildea, 2008). The operating temperature in two reactors is set at 333 K below the boiling point 337.8 K of methanol at 1 atm. In addition, most of the sunflower oil should be consumed in the first reactor, and the amount of the sunflower oil left to react in the second reactor is small, so only a small amount of NaOH are complemented to the second reaction. Therefore, the amount of NaOH added into the first reactor is set at 1.0 wt% (by weight of oil), and the 0.20 wt% (by weight of oil) NaOH is set to complement into the second reactor. It is known that the amount of NaOH has little effect on the reaction rate if the concentration of NaOH is higher than 0.5 wt% (Bambase et al., 2007), so this setting makes the concentration of NaOH in both reactors higher than 0.5 wt%, and this ensures that the kinetics parameter in Table 1 is applicable in the two-step process computation. Finally, there are only four process technical parameters left to be determined in Fig. 1, which are the overall mole ratio N of total methanol fed into two reactors to raw oil, the total reaction time t which is the sum of the reaction time in two reactors, and the split ratio x of methanol supplied in the first reactor to the total methanol, the split ratio y of the reaction time of the first reactor to the total reaction time. According to the current study, for the one-step biodiesel production, the appropriate methanol/oil mole ratio is 6:1 and the transesterification can achieve the equilibrium after 30 min. Therefore, N is considered to vary in the range of 5–9, t is considered to vary in the range of 20–40 min. Obviously, x and y should vary in the range of 0–1. The two-step transesterification process in Fig. 1 can be computed if N, t, x and y are known, and it is programmed by means of the software Matlab. According to the change of the biodiesel
yield in the variation ranges of N, t, x and y, a group of suitable process parameters, i.e. N, t, x and y, can be obtained. The effect of the total methanol/oil mole ratio N and the total reaction time t on the biodiesel yield is shown in Fig. 2, and data of the biodiesel yield in Fig. 2 are corresponding to the maximum value calculated in the 0–1 range of x and y at fixed N and t. It indicates apparently that increase of both N and t are in favor of enhancing the biodiesel yield, but the increase of methanol amount has little influence on the yield of biodiesel when N is greater than 7. Almost all of the yields of biodiesel can be beyond 99% in the whole range of 20–40 min if N is greater than 7. It should be noted that the increase of N can result in the more separation cost because the more excess methanol has to be recovered in the following separation process, and the extension of t also need the bigger reactor volume for the continuous production process. Therefore, N = 7 and t = 30 min are considered to be appropriate reaction process parameters according to the trade-off between the biodiesel yield and the process cost. Though the amount of methanol in the two-step reaction process is more than the one-step reaction process, the biodiesel yield of the two-step process at N = 7 and t = 30 min is 99.41% and almost all of material oil is converted into biodiesel. At N = 7 and t = 30 min, Fig. 3 indicates how the yield of biodiesel changes with the variation of the split ratios of the methanol and the reaction time between two reactors. The computational result shows that the maximum yield of biodiesel is 99.41 at x = 0.84 and y = 0.53, so in the two-step reaction process, the first reactor consumes more methanol than the second reactor while the respective reaction time of the two reactors is almost same. Moreover, as shown in Fig. 3, the change of x and y at the vicinity of the maximum value has no obvious impact on the yield of biodiesel, and it means that the yield of biodiesel is not sensitive to the split ratios of the methanol and the reaction time between two reactors in the two-step reaction process. 2.3. The stream information of the two-step transesterification process At N = 7, t = 30 min, x = 0.84 and y = 0.53, compositions of the main streams of the two-step reaction process in Fig. 1 are summarized in the Table 2. The concentration of the intermediate product MG in the two-step reaction product decreases significantly from 2.067 wt% to 0.3364 wt% in comparison with the one-step reaction. The yield of biodiesel is 99.41% after the second reaction, and it also has a significant increase in comparison with 96.62% of the
Fig. 2. Effect of N and t on the biodiesel yield.
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Fig. 3. Effect of x and y on the biodiesel yield.
one-step reaction. It should be noted that the yield calculated is a little greater than the experimental yield obtained by other researchers because the occurrence of the side reactions leading to the loss of raw oil is ignored in this work. Moreover, most of data reported about the yield of biodiesel were related to the yield of the whole reaction–separation process, however, the yield in this work is only related to the reaction process and the loss of biodiesel in the separation process is not considered. 3. The separation process of the two-step biodiesel production The two different separation processes is investigated rigorously by means of the process simulation software Aspen Plus, one for the traditional distillation separation process, another for the improved separation process. The difference between two separation processes locates in the separation method of methyl ester and methanol. As discussed in introduction of this work, the improved process utilizes the water washing operation to replace the costly methyl ester–methanol distillation operation, and its feasibility is investigated by the rigorous process simulation in this work. The feeds of the two separation processes, S1 and S2, come from the two-step reaction process shown in Fig. 1, and its information is listed in Table 2, so the intermediate products of the transesterification, i.e. MG and DG, can emerge in the separation process. The composition of the biodiesel product in both separation processes is specified to meet the Germany specification for biodiesel (DIN V 51606) as shown in Table 3 (Dimian and Bildea, 2008). The information of the separation process simulation together with the simulation information of the two-step reaction process can provide the whole process information of the two-step biodiesel production. 3.1. The setup of the process simulation The main equipments in the separation process of the biodiesel production include the liquid–liquid gravity settlers, distillation
Components
Unit
Min.
Max.
Water MG/DG/TG Free glycerol Methanol
wt% wt% wt% wt%
– – – –
0.03 0.8/0.4/0.4 0.02 0.3
columns, extraction columns, flash drums and heat exchangers. The gravity settler and flash drum are set to work under the adiabatic condition. The sieve-tray column is chosen to carry out distillation and extraction operation in the separation process. Due to the high viscosity of both biodiesel and glycerol, in the distillation operation, the sieve-tray efficiency is set to 0.6, and the pressure drop of the single tray is set to 0.7 kPa. The tray efficiency of the extraction column is set to be 0.8. The boiling points of both biodiesel and glycerol are very high, and their decomposition temperature is lower than the boiling point, so the vacuum distillation is adopted to recover methanol from the biodiesel phase and glycerol phase. The pressure of methanol–biodiesel distillation column is set to 11.65 kPa in order to prevent decomposition of biodiesel. But for the methanol–glycerol distillation column, it is difficult to prevent the decomposition of glycerol only by the vacuum operation. In the biodiesel industry, a water steam is usually introduced into the methanol–glycerol distillation column to reduce the bottom temperature, and the purification of glycerol is still a very tough problem so far. Because the study on purification of glycerol is not a key point of this work, similar to the work of Zhang et al. (2003), the recovery of methanol from the glycerol phase is conducted only in a simple distillation column under the atmosphere pressure. The concentration of methanol recovered from biodiesel and glycerol in the recycling methanol stream must meet the requirement of the transesterification because the impurities, especially water, have serious influence on the transesterification. In this work the concentration of methanol in the recycling methanol streams is specified to be higher than 99.8 wt%. For both of the methanol–glycerol distillation column and the methanol–biodiesel distillation column, the tray number and reflux ratio are adjusted to meet to the requirement of the methanol concentration, and the tray numbers of 9–10 and the reflux ratio of 1–2 can achieve the separation target due to the big difference of boiling points between methanol and biodiesel/glycerol. In the separation process of the biodiesel production, except heat exchangers, flash drums and distillation columns, other equipments usually work under the normal temperature and pressure, and there is no corrosive and harmful material except NaOH. Pumps and pipes mainly change the pressure of streams, so they are absent in the process simulation. The equipment investment of the heat exchanger and flash drum is much lower than that of the distillation column, so it is supposed that the equipment cost in the separation process is mainly composed of the distillation column and its accessory equipments. The energy consumption of the separation process also mainly locates in the distillation operation due to high boiling points of methyl ester and glycerol. The catalyst NaOH is not taken into account in the simulation, and there are two reasons for neglect of NaOH. One reason is that
Table 2 Information of the main streams in the two-step transesterification process. Stream name
TG (wt%)
DG (wt%)
MG (wt%)
ME (wt%)
M (wt%)
G (wt%)
Mass rate (kg/h)
R1 P1 S2 S1
0.0825 0.0928 <0.01 0.0166
0.8075 0.9082 <0.01 0.1499
2.065 2.322 0.0185 0.3364
79.94 89.92 0.7185 89.36
9.213 6.223 32.95 9.009
7.888 0.5328 66.30 1.125
13,399 11,900 1498.2 12,299
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the presence of the little catalyst does not affect the thermodynamic equilibrium, and another reason is that in the industrial process the free glycerol in biodiesel is removed in a water washing column where NaOH can be also removed easily by means of the sour water washing at the same time. The water washing column still exists in the simulation even if the catalyst NaOH is ignored, so the reliability of the comparison between two different separation processes can be assured. In addition, the generation of soap in the biodiesel production is ignored because there is little water and free fatty acids in the refined sunflower oil. The neglect of NaOH and soap in the separation process just decrease the consumption of washing water, and it is little influence on other results of the process simulation. For some compounds emerging in the separation process, such as DG and MG, they are not contained in the database of the process simulation software Aspen Plus, therefore their properties are calculated in the simulation by means of the group contribution method according to their molecular structure. For the methanol–glycerol–methyl ester system, the computational accuracy of the UNIFAC–Dortmund thermodynamics method can meet the requirement of the simulation very well because the concentration of the intermediate product MG is very low after the second transesterification. For the liquid–liquid equilibrium of the methanol–water–methyl ester ternary system, the computational results by means of the UNIFAC–Dortmund method are also in good accordance with experimental results of Stloukal et al. (1997), and the maximum deviation between experimental and computational data of the composition in the water phase and methyl ester phase is no more than 5%. Therefore, the UNIFAC–Dortmund method is utilized to calculate the thermodynamic equilibrium in the simulation of two separation processes. 3.2. The traditional distillation separation process According to works of Van Gerpen (2005) and Fukuda et al. (2001), the process flow diagram of the traditional distillation sep-
aration is presented in Fig. 4. The second transesterification product S1 enters the gravity settler D-2. By means of the liquid– liquid phase separation, S1 are separated to the methyl ester phase MEP and glycerol phase GLP in D-2. The upper methyl ester phase MEP of D-2 is fed into the methanol–biodiesel distillation column T-2. The overhead product of T-2 is methanol and it is a part of the recycling methanol S5. The bottom product S7 of T-2 is the crude biodiesel containing impurities such as free glycerol. The glycerol phase S2 and GLP of the first and second transesterification products are fed into the methanol–glycerol distillation column T-1. The overhead product of T-1 is another part of the recycling methanol S5, and the bottom product of T-1 is crude glycerol. The impurity in crude biodiesel stream S7 is removed by 55 kg/h washing water W in the washing column E-1, and a very small amount of water in S6 is removed from the top of the flash drum F-1 after S6 is heated up to 473.1 K in the heat exchanger H-2. Finally, the biodiesel product S3 which meets the Germany specification for biodiesel (DIN V 51606) can be obtained after being cooled in the heat exchanger H-4. The composition and mass rate of the main streams involved in Fig. 4 is summarized in Table 4 and the technical information of the separation process is shown in Table 6. 3.3. The improved separation process As shown in Fig. 5, in the improved separation process, the second transesterification product S1 is fed into the water washing column E-1 at first, and is washed by 55 kg/h water WW. E-1 has four theoretical trays, and can remove most of methanol and all of free glycerol. After water washing, the concentration of methanol in the crude biodiesel stream S7 is 1.751 wt% in comparison with initial 9.009 wt% in S1. The small amount of methanol and water in S7 can be removed easily by the adiabatic flash drum F-1 because their boiling points are much lower than that of biodiesel. To order to improve the recovery of methanol, the vapor product S6 of F-1 is cooled, then mixed with the bottom product S8 of E-1 and the glycerol phase S2, and finally fed into the
Fig. 4. Process flow diagram of the traditional distillation separation process.
Table 4 Information of the main streams in Fig. 4. Stream name
TG (wt%)
DG (wt%)
MG (wt%)
ME (wt%)
M (wt%)
G (wt%)
W (wt%)
P (kPa)
T (K)
Mass rate (kg/h)
S3 S4 S5 S6 S7
0.0184 <0.01 <0.01 0.0184 0.0184
0.1673 <0.01 <0.01 0.1663 0.1667
0.3711 0.0647 <0.01 0.3691 0.3700
99.41 1.061 0.1059 99.13 99.38
<0.01 0.8565 99.81 0.0212 0.0239
<0.01 98.02 0.0834 <0.01 0.03894
0.0280 0 0 0.2944 0
101.3 101.3 101.3 101.3 101.3
298.1 298.1 313.2 331.6 522.8
11,022 1148.9 1595.3 11,084 11,055
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Fig. 5. Process flow diagram of the improved separation process.
Table 5 Information of the main streams in Fig. 5. Stream name
TG (wt%)
DG (wt%)
MG (wt%)
ME (wt%)
M (wt%)
G (wt%)
W (wt%)
P (kPa)
T (K)
Mass rate (kg/h)
S3 S4 S5 S6 S7 S8 S9
0.0184 <0.01 <0.01 <0.01 0.0181 <0.01 <0.01
0.1672 <0.01 <0.01 <0.01 0.1638 <0.01 <0.01
0.3199 0.5357 <0.01 <0.01 0.3135 0.5490 0.5132
99.41 2.925 <0.01 4.331 97.52 1.763 3.222
0.0768 0.2049 99.81 84.26 1.751 81.98 0.4904
<0.01 94.90 <0.01 <0.01 <0.01 12.45 91.05
0.0108 1.431 0.1871 11.40 0.2373 3.260 4.727
101.3 101.3 101.3 20.26 101.3 101.3 107.6
298.1 298.1 351.3 423.1 336.6 335.7 459.9
11,027 1190.6 1590.2 223.72 11,250 1111.4 1243.1
Table 6 Technical information of the two separation processes. Items Heat exchanger
H-1 H-2 H-3 H-4
Separation unit
T-1
T-2
E-1 E-2 F-1 F-2
Stream in/out temperature Energy duty (MW) Stream in/out temperature Energy duty (MW) Stream in/out temperature Energy duty (MW) Stream in/out temperature Energy duty (MW) Reboiler temperature (K) Pressure (kPa) Reboiler duty (MW) Condenser duty (MW) Diameter (m) Tray number/reflux ratio Reboiler temperature (K) Pressure (kPa) Reboiler duty (MW) Condenser duty (MW) Diameter (m) Tray number/reflux ratio Theoretical stages Temperature (K) Theoretical stages Temperature (K) Temperature (K) Pressure (kPa) Temperature (K) Pressure (kPa)
(K) (K) (K) (K)
Fig. 4
Fig. 5
528.9/333.1 1.414 331.6/473.1 0.9883 557.5/298.1 0.2226 473.1/298.1 1.158
336.0/423.1 0.6357 423.1/298.1 0.7873 431.7/298.1 0.03410 431.7/298.1 0.1099
557.5 101.3 0.6588 0.4564 0.5 10/1 528.9 11.65 1.900 0.5560 1.0 10/1 5 333.1 –
459.9 101.3 1.480 1.418 0.9 10/2 –
473.1 20.26 –
423.1 20.26 459.9 101.3
4 333.1 –
Separation effect
Recovery efficiency of methanol (%) Recovery efficiency of ME (%) Biodiesel output (kg/h)
99.22 99.70 11,022
99.09 99.73 11,027
Number of the main separation equipments
Distillation column Washing column Flash evaporator Heat exchanger
2 1 1 4
1 1 2 4
Energy duty
The total heat duty (MW) The total cold duty (MW)
3.547 3.534
2.226 2.239
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glycerol–methanol distillation column T-1. The overhead product of T-1, the methanol stream S5, is recycled to the reaction process. The bottom product of T-1 contains the glycerol and water, and water is removed through the flash operation so that crude glycerol is obtained. The information of streams and equipments in the improved separation process are shown in Tables 5 and 6 . The compositions of biodiesel produced by the improved separation process can meet the Germany specification for biodiesel very well, so the improved separation process can achieve the separation target of the biodiesel production. Although it could be difficult to conduct the liquid–liquid extraction operation in water washing column E-1 due to the small density difference between the heavy and light phase (Tizvar et al., 2009), the centrifugal extraction operation can be utilized to solve this problem. 3.4. The comparison between the traditional and improved separation process The parameters of main equipments, energy duty and separation effect of the traditional and improved separation processes are summarized in the Table 6. In order to compare the separation effect of the two processes, the recovery efficiency of methanol and biodiesel are defined as following formulas:
The recovery efficiency of methanol ¼
F S5 wM;S5 100% F S1 wM;S1 þ F S2 wM;S2
The recovery efficiency of methyl ester ¼
F S3 wME;S3 100% F S1 wME;S1
where, F is the mass flow rate of streams, wM and wME are the mass fraction of methanol and methyl ester in the corresponding stream. As shown in Table 6, the two investigated separation processes have similar separation effect, and both can meet the specification of biodiesel. For the traditional distillation process, the equipment investment and energy duty are high because of the presence of two distillation columns in the process and high boiling points of biodiesel and glycerol. However, the improved separation process has significant advantage. There is only one distillation column in the improved separation process, and its energy duty decreases by about 35% in contrast with the traditional distillation process. The results from simulation show that it is feasible for the improved separation process in the further application. 4. Conclusion By means of the rigorous process simulation, the two-step biodiesel production process made from the refined sunflower oil is investigated in detail. The total methanol/oil mole ratio, the total reaction time and the split ratios of both methanol and reaction time between two reactors are determined quantitatively. The technical information about the traditional distillation separation process and the improved separation process are acquired and
compared. The improved separation process, where the water washing operation is utilized to replace the costly methyl ester– methanol distillation operation, is verified to be of advantage in reducing energy consumption and equipment investment. References Bournay, L., Casanave, D., Delfort, B., Hillion, G., Chodorge, J.A., 2005. New heterogeneous process for biodiesel production: a way to improve the quality and the value of the crude glycerin produced by biodiesel plants. Catal. Today 106 (1), 190–192. Bambase, M.E., Nakamura, N., Tanaka, J., Matsumura, M., 2007. Kinetics of hydroxide-catalyzed methanolysis of crude sunflower oil for the production of fuel-grade methyl esters. J. Chem. Technol. Biotechnol. 82 (3), 273–280. Cayhli, G., Kusefoglu, S., 2008. Increased yields in biodiesel production from used cooking oils by a two step process: comparison with one step process by using TGA. Fuel Proc. Tech. 89 (2), 118–122. Chiu, C.W., Goff, M.J., Suppes, G.J., 2005. Distribution of methanol and catalysts between biodiesel and glycerin phases. AIChE J. 51 (4), 1274–1278. Dimian, A.C., Bildea, C., 2008. Chemical Process Design: Computer-Aided Case Studies. Wiley-VCH, New York. Darnoko, D., Cheryan, M., 2000. Kinetics of palm oil transesterification in a batch reactor. JAOCS 77 (12), 1263–1267. Fukuda, H., Kondo, A., Noda, H., 2001. Biodiesel fuel production by transesterification of oils. J. Biosci. Bioeng. 92 (5), 405–416. Freedman, B., Butterfield, R.O., Pryde, E.H., 1986. Transesterification kinetics of soybean oil. JAOCS 63 (10), 1375–1380. Goff, M.J., Bauer, N.S., Lopes, S., Sutterlin, W.R., Suppes, G.J., 2004. Acid-catalyzed alcoholysis of soybean oil. JAOCS 81 (4), 415–420. Haas, M.J., McAloon, A.J., Yee, W.C., Foglia, T.A., 2006. A process model to estimate biodiesel production costs. Bioresource Technol. 97 (4), 671–678. Kusdiana, D., Saka, S., 2001. Kinetics of transesterification in rapeseed oil to biodiesel fuel as treated in supercritical methanol. Fuel 80 (5), 693–698. Komers, K., Skopal, F., Stloukal, R., Machek, J., 2002. Kinetics and mechanism of the KOH-catalyzed methanolysis of rapeseed oil for biodiesel production. Eur. J. Lipid Sci. Technol. 104 (11), 718–737. Mjalli, F.S., San, L.K., Yin, K.C., 2009. Dynamics and control of a biodiesel transesterification reactor. Chem. Eng. Tech. 32 (1), 13–26. Ma, F.R., Hanna, M., 1999. A biodiesel production: a review. Bioresource Technol 70 (1–4), 1–15. Noureddini, H., Zhu, D., 1997. Kinetics of transesterification of soybean oil. JAOCS 74 (11), 1457–1463. Negi, D.S., Sobotka, F., Kimmel, T., 2006. Liquid–liquid phase equilibrium in glycerol–methanol–methyl oleate and glycerol–monoolein–methyl oleate ternary systems. Ind. Eng. Chem. Res. 45 (10), 3693–3696. Stiefel, S., Dassori, G., 2009. Simulation of biodiesel production through transesterification of vegetable oil. Ind. Eng. Chem. Res. 48 (3), 1068– 1071. Stloukal, R., Komers, K., Machek, J., 1997. Ternary phase diagram biodiesel fuel– methanol–water. J. Prakt. Chem. 339 (5), 485–487. Tizvar, R., McLean, D.D., Kates, M., Dube, M.A., 2009. Optimal separation of glycerol and methyl oleate via liquid–liquid extraction. J. Chem. Eng. Data 54 (5), 1541– 1550. Van Gerpen, J., 2005. Biodiesel processing and production. Fuel Process. Technol. 86 (10), 1097–1107. Veljkovic, V.B., Lakicevic, S.H., Stamenkovic, O.S., Todorovic, Z.B., Lazic, M.L., 2006. Biodiesel production from tobacco (Nicotiana tabacum L.) seed oil with a high content of free fatty acids. Fuel 85 (17–18), 2671–2675. West, A.H., Posarac, D., Ellis, N., 2008. Assessment of four biodiesel production processes using HYSYS. Plant. Bioresource Technol. 99 (14), 6587–6601. Zapata, C.D., Martinez, L.D., Castiblanco, E.A., Uribe, C.A.H., 2007. Biodiesel production from crude palm oil: 1. Design and simulation of two continuous processes. Dyna – Colombia 74 (151), 71–82. Zhang, Y., Dube, M.A., Mclean, D.D., Kates, M., 2003. Biodiesel production from waste cooking oil: 1. process design and technological assessment. Bioresource Technol. 89 (1), 1–6. Zhou, H., Lu, H.F., Liang, B., 2006. Solubility of multicomponent systems in the biodiesel production by transesterification of Jatropha curcas L. oil with methanol. J. Chem. Eng. Data 51 (3), 1130–1135.