Biodiesel Process Intensification by Using Static Mixers Tubular

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Biodiesel Process Intensification by Using Static Mixers Tubular Reactors E. Santacesaria,*,† R. Turco,† M. Tortorelli,† V. Russo,† M. Di Serio,† and R. Tesser† †

Department of Chemistry of the University of Naples “Federico II”, Complesso Universitario Monte S. Angelo, Via Cinthia 4, IT 80126 Naples, Italy ABSTRACT: Biodiesel is usually produced by reacting triglycerides, contained in vegetable oils, with methanol in the presence of KOH, NaOH, or related alkoxides as catalysts. In industry, the reaction is performed in stirred tank reactors and requires 12 h of reaction time being the reaction rate strongly affected by mass transfer limitation. We have recently shown, by using a Corrugated Plates Heat Exchanger Reactor, that a very high productivity (about 2 tons/day L) can be obtained by working at 60100 °C thanks to the presence of an intense local “micromixing”. Moreover, we have recently tested the performances obtained in a tubular reactor filled with stainless steel spheres of different diameters. By opportunely changing the spheres diameters it is possible to obtain microchannels in a range of 3001000 μm with an intense local micromixing. Again, thanks to micromixing we obtained very high productivities. However, in these last reactors the void portion of the reactor is low and the productivity per overall reactor volume is relatively low. It is possible to obtain better results, in terms of productivity, by filling the tubular reactor with stainless steel wool, being in this case the void fraction about 0.9. In the present work, some of the mentioned systems will be compared for their performances by using different amounts of KOH as catalyst (1 or 2% b.w. of oil). A dramatic change in biodiesel yield has been observed in all cases passing from 1 to 2% of catalyst independently of the reactants flow rate. These behaviors cannot be interpreted with the pseudomonophasic kinetic models, normally reported in the literature. At this purpose, for interpreting all the observed kinetic behaviors a new biphasic kinetic model, based on a reliable catalytic mechanism, has been developed. This model has been applied, first of all, to data reported in the literature related to runs performed in batch conditions with the scope of estimating the kinetic parameters, and then it has been applied to all the runs performed in continuous reactors with a satisfactory agreement.

1. INTRODUCTION Biodiesel is usually produced in industry by reacting vegetable oils with methanol in the presence of a homogeneous catalyst such as NaOH, KOH, or related alkoxides, at 60 °C (the methanol boiling point) and atmospheric pressure. The reagents are immiscible liquids and the interphase plays a fundamental role, as it will be seen in this paper. Many of the published works reporting kinetic data are related to batch reactors and are often affected by mass transfer limitation.17 In a recently published paper and patent8,9 we have shown that a very high productivity (2000 kg/day L) can be obtained by performing the transesterification reaction in a continuous corrugated plates heat exchanger reactor (CP-HEX reactor) working in a range of temperatures between 60 and 100 °C. The high performances observed are due to a very active local micromixing, because, the conversion increases by increasing the overall liquid flow rate. Therefore, it derives that all the reactors favoring a local micromixing or more generally increasing the liquidliquid interfacial area can give high performances in this reaction. This observation is confirmed by many results, reported in the literature, in which the transesterification reaction is performed in the following: reactors containing a static mixer, microwave irradiated reactors, ultrasonic irradiated reactors, centrifugal contactors, rotating packed bed reactors, and jet flow stirred reactors. It is then possible to increase the liquidliquid interfacial area also by using microreactors. Microchannels of sizes less than 300 μm allow to improve very much mass, heat, and momentum transfer.1012 The use of the different mentioned technologies r 2011 American Chemical Society

and related performances have recently been reviewed by Qiu, Zhao, and Weatherley.13 In other words, transesterification of triglycerides with methanol seems strongly promoted by the extension of the polar/apolar interface area. As the most commonly used catalyst KOH reacts in methanol giving place to KOCH3 according to the reaction K þ OH þ CH3 OH a K þ OCH3 þ H2 O

ð1Þ

The high interface area initially occurs to give the possibility to the anion CH3O to react with a molecule of triglyceride so transferring the anion charge in the oil phase in the form of diglycerolate.14 Higher is the initial interface area faster is the CH3O disappearance and the formation in situ of the true catalyst, that is, a mixture of di- and monoglycerolate.14 Equilibrium (1) is clearly shifted to the right as a consequence of the methanolate disappearance. Then, the reaction rate abruptly declines, because of the ultimate formation of glycerolate anion. This anion migrates in the polar phase being insoluble in the oil phase and is poorly active in the reaction. According to us, this behavior is responsible also of a never explained experimental phenomenon observed in batch kinetic runs, that is, by using Special Issue: CAMURE 8 and ISMR 7 Received: July 28, 2011 Accepted: October 12, 2011 Revised: October 11, 2011 Published: October 12, 2011 8777

dx.doi.org/10.1021/ie201640w | Ind. Eng. Chem. Res. 2012, 51, 8777–8787

Industrial & Engineering Chemistry Research different amounts of catalyst a different plateau is observed in the conversion5,6 not corresponding to the chemical equilibrium composition. As it will be seen in this work, this aspect, already observed in batch conditions, appeared dramatic in continuous reactors and passing from 1 to 2% b.w. of catalyst, the conversion changes from 8085 to 9597% by using the same reactor and operative conditions of flow rate, temperature, and methanol/oil ratio. In a recently published paper15 we proposed the use of three different tubular reactors respectively containing spheres of uniform size (2.5 mm of diameter) having microchannels of about 1000 μm, spheres of two different sizes of respectively 2.5 and 1 mm of diameters having microchannels of about 500 μm and two different sizes of respectively 2.5 and 0.39 mm of diameters having microchannels of about 300 μm, for studying the methanol-soybean oil transesterification reaction. The performances obtained were so high that none of the pseudohomogeneous kinetic models proposed by the literature, derived from runs performed in batch conditions, is suitable to describe the observed kinetic behaviors. Other packing materials can be used in tubular reactors for obtaining a more intense micromixing. In particular, in this work, some threaded rods of the same length of the tubular reactor have been used as static mixer elements. However, all these tubular reactors could be useful as laboratory reactors for orienting on the choice of the best static mixer or for modeling microreactors but are not useful for industrial purposes being the productivity, for reactor volume unity, relatively low. By filling the reactor with stainless steel wool15 it is possible to obtain a much higher void fraction accompanied by an increased productivity that renders the reactor interesting also for industrial scale up. The performances obtained, with all the mentioned types of reactors, are compared and quantitatively interpreted in this paper. At this scope, a new biphasic model, based on a reliable catalytic mechanism, more adherent to the physical behavior of the system, has been developed. This model is able to describe both batch and continuous runs with a reasonable number of parameters. A discussion on the catalytic mechanism, the biphasic kinetic model, and the obtained results will conclude the paper.

2. EXPERIMENTAL SECTION 2.1. Reagents. Soybean oil used in the runs was purchased in a local food-store (the fatty acid composition of the used soybean oil, determined by gas chromatographic analysis, was (%, w/w) as follows: palmitic = 11, stearic = 4, oleic = 23, linoleic = 56, linolenic = 5, others = 1). All other employed reagents (when not specified) were supplied by Aldrich at the highest level of purity available and were used as received without further purification. 2.2. Analytical Method. The FAME (fatty acid methyl esters) yields were determined by using 1H NMR technique (GEMINI 200 Mz), measuring the area of the H NMR signal related to methoxylic (A1, single signal at δ = 3.7 ppm) and methylenic groups (A2, triplet signal at δ = 2.3 ppm), respectively. The FAME yields can be calculated by using the following equation 2:A1 100 ð2Þ yieldFAME ¼ 3:A2 3

The 1H NMR spectra were obtained with GEMINI-200 equipment in deuterated chloroform.16

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Table 1. Characteristics of All the Used Reactorsa volumes (cm3)

void fractions

RP1

7.15

0.45

RP21

6.05

0.38

RP22

3.50

0.22

RWTR

13.9

0.89

TRR

6.00

0.38

reactors

a

The void fractions are referred to an empty tubular reactor of 15.7 cm3.

In all cases the conversion values were confirmed by gaschromatographic analysis [UNI 10946:2001], using a gas chromatograph (Perkin-Elmer model Clarus 500), equipped with a flame ionization detector (FID), an on-column injector, and employing a FS-HP5 column (10 m 0.32 mm, I.D. 0.1 μm film). Before the analysis, the samples were derivatized with BSTFA (N,Obis (trimethylsilyl) trifluoroacetamide with trimethylchlorosilane). As internal standards, methylheptadecanoate was used for methyl ester and 1,2,4-butantriol was used for glycerol. The fluid dynamic tests have been performed analyzing the samples by UVvisible measurements. 2.3. Description of the Reactors Considered for the Comparison and Related Kinetic Runs. Five different reactors have been tested starting from a stainless steel cylindrical tube of 20 cm length and 12.7 mm (1/2”) of external diameter with an internal diameter of 10 mm. The internal void volume of the reactor was 15.7 cm3. This tube was filled by using different static mixer elements. It has been filled, for example, with 8 stainless steel threaded rods, of the same length of the reactor, having 3 mm of diameter and 0.5 mm of thread, forced into the cylinder (the acronym for this reactor is TRR). In a previous work,15 we filled the same tube with stainless steel spheres of different diameters. At this purpose, three different kinds of filling were used: (i) a single size packing (RP1) in which the reactor was filled with spheres of 2.5 mm diameter; (ii) a dual size packing (RP21) in which spheres of two different diameters (2.5 and 1.0 mm, respectively) were used; (iii) again, a dual size packing (RP22) in which spheres of two different diameters (2.5 and 0.39 mm, respectively) were used. A theoretical hexagonal packing of the bigger spheres RP1 leads to two different kinds of cavities called octahedral (the bigger ones) and tetrahedral (the smaller ones). If the ratio between the radius of the spheres is r/R = 0.414, the smaller spheres occupy the octahedral cavities, while, the tetrahedral cavities are filled just if r/R = 0.227. In our case, we have chosen first of all a r/R = 0.156 in order to occupy both the cavities; this value has been obtained choosing a diameter for the smaller spheres of 0.39 mm. Therefore, the RP1 packing structure allows the introduction of the smaller spheres of 0.39 mm of diameter without altering the geometrical framework created by the bigger spheres. In the case of the RP22 structure the introduction of spheres of 1.0 mm of diameter forces the bigger spheres to assume a geometrical cubic framework. An estimation of the size of the microchannels has been made considering an ideal packing, so obtaining an average value of about 1000 μm for the structure RP1, 500 μm for the RP22, and 300 μm for the RP21. The void fractions obviously change in the three mentioned cases, as it can be appreciated in Table 1, where, all the characteristics of the used reactors are reported for comparison. The described tubular reactors, filled in different ways, are 8778

dx.doi.org/10.1021/ie201640w |Ind. Eng. Chem. Res. 2012, 51, 8777–8787

Industrial & Engineering Chemistry Research

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Figure 1. Scheme of the used laboratory plants.

Table 2. Dispersion Numbers Obtained for All the Different Used Reactors reactor

flow rates (cm3 min‑1)

1/Pe = D/uL

RP1

1.45.6

0.0330.062

RP21

5.68.3

0.0350.050

RP22

1.45.68.3

0.0830.10

RWTR

5.68.3

0.0360.042

TRR

5.68.3

0.070.091

useful as laboratory reactors to foresee the behavior of the transesterification reaction in the presence of static mixers (intense local micromixing) or microreactors (presence of microchannels) but cannot be considered the base for the development of an industrial size reactor, being the void fraction of the reactor too low as it can be appreciated always in Table 1. For this reason, in a paper recently submitted for publication15 we proposed a tubular reactor filled with a stainless steel ribbon wool (RWTR) having a void fraction of 0.89 (see Table 1). The laboratory plant is the same for all the described reactors changing only the type of internal static mixer. A scheme of this plant is reported in Figure 1. Kinetics runs have been made by feeding to the reactors two different liquid flows: one of methanol containing the dissolved catalyst (KOH, normally 1 or 2 wt % with respect to the oil) and another of refined soybean oil. The oil was preheated at the reaction temperature, whereas methanol was fed at room temperature. Two HPLC pumps have been used for feeding the reagents. The typical experimental conditions of the runs performed, using the different described reactors, are as follows: molar ratio methanol/soybean oil = 6:1 considering for soybean oil an average molecular weight of 885, a KOH catalyst concentration of 1 or 2% b.w. referred to the oil and a temperature of 60 °C. 2.4. Fluid Dynamic Characterization Tests. All the described reactors have been subjected to a fluid dynamic characterization with a stimulus-response step test to evaluate the RTD (retention time distribution). Water has been used as flowing medium, and cobalt nitrate(II) hexahydrate, having a characteristic pink color, has been used in the test as a tracer.

Figure 2. Methylesters yield as a function of the residence time obtained with different reactors by using 1.0 wt.% and 2.0 wt.% of catalyst.

A solution of known concentration (C0) of the tracer was sent to the reactor at time = 0; from this time samples have been periodically withdrawn from the outgoing flow in order to record the step response. Then, the samples have been analyzed by an UVvisible spectrophotometer, monitoring the absorbance of cobalt nitrate at 511 nm. The tests have been repeated at different flow rates. An HPLC pump was used for regulating the cobalt solution flow rate. The results have been interpreted by plotting the dimensionless tracer concentration F = C/Co as a function of the dimensionless time Θ = t/tm, being that tm is the residence time, obtaining a series of sigmoidal curves. The interpretation of these curves in the way reported in a previous work15 allows the evaluation of the dispersion number, corresponding to the reciprocal of the Peclet number 1 D ¼ ð3Þ Pe u3L The ND numbers for all the tested reactors are reported in Table 2. As it can be seen, the estimated ND values are close to the ones expected for the static mixers 1/ND = Pe = 70 3 L17 (where L stands for the length of the reactor expressed in meters). It is important to underline that for all the described systems the local turbulence is not given mainly by the flow rate, but by the static mixer elements that warrants an intimate ND ¼

8779



Figure 3. Anions population in agreement with the catalytic mechanism via enolates.

mixing between the two immiscible phases, because of the eddies formation. However, as these numbers are near to zero the flows, in all cases, can be approximated to a plug flow. Moreover, from our previous studies, we have demonstrated that the static mixer reactors present low Reynold’s number (>100), so it could be possible to conclude that they work in laminar flow, but despite this the very low dispersion numbers found approximate the fluid-dynamic plug flow behavior.

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catalyst amount gives place to much higher yields. A so large increase in the yields, simply by increasing the amount of used catalyst, is unexpected, and, as it will be seen later, this behavior cannot be described by using a monophasic pseudohomogeneous kinetic model, normally, adopted in the literature for describing batch runs.5,6 3.2. Mechanism of the Methanol-Soybean Oil Transesterification Reaction. From the observation of the kinetic data collected by using all the described reactors it is possible to conclude that the increase of the liquidliquid interface area is beneficial for the reaction. Moreover, when the interface area is high the reaction consistently occurs also for very low residence time. The same observation can be derived also considering the data collected in the literature13 with different techniques like the following: oscillatory flow reactors, microwave irradiated reactors, ultrasonic irradiated reactors, centrifugal contactors, rotating packed bed reactors, and jet flow stirred reactors. The conclusion could be that methanolate anions, formed with the reaction 1, react at the liquidliquid interface with triglycerides transferring the charge from the polar to the apolar phase. Greater is the initial interface area, greater is the methanolate disappearance rate. After this reaction, occurring at the interface, the transesterification takes place in the oil phase through three different consecutive steps, according to the schemes 4-6

3. RESULTS AND DISCUSSION 3.1. Methanol-Soybean Oil Transesterification, a Comparison of the Runs Performed in Different Packed Bed Tubular Reactors (TRR, RP1, RP21, RP22, RWTR). To examine the effect

of the presence of a packed bed, a first run was performed with the empty reactor using a total flow rate of oil and methanol of 8.33 cm3/min. The methylesters yield obtained in this case was only 3%. This value can be compared with the much higher values obtained in the presence of a packed bed as it can be seen in Figure 2, where the yields obtained, as a function of residence time, with the different tested reactors are reported for comparison. As already mentioned, the experimental runs were all performed at 60 °C with a methanol/oil molecular ratio equal to 6 with the aim to evaluate the effect of the residence time on the methylesters yields. First of all, it must be observed, that in any case the presence of a packed bed gives place to a very strong increase of the conversion due to the static mixing effect. Then, it must be pointed out that, in many cases, by decreasing the residence time, the yield in methylesters increases until reaching a maximum value. This behavior can be explained by considering that increasing the flow rate, in these systems, means to increase the liquidliquid interface. Therefore, the lower values of the yield, obtained at the highest residence time, are probably due to a less active micromixing, having as a consequence the reduction of the interface area. This has a negative consequence on both the methanolate anion draining and the methanol mass transfer rate from the polar to the apolar phase. The decrease of the yields, observed at the lowest residence time, seems the normal behavior of a reacting system in a kinetic regime requiring time to occur. At last, as it can be seen in Figure 2, several runs have been performed by using 2% of KOH catalyst. It is interesting to observe that the increase of the

1Þ T þ CH3 OH / D þ E

ð4Þ

2Þ D þ CH3 OH / M þ E

ð5Þ

3Þ M þ CH3 OH / G þ E

ð6Þ

where T represents the triglycerides, D represents the diglycerides, M represents the monoglycerides, G represents glycerol, and E represents the methylesters. These reactions are catalyzed by di- and monoacylglycerolate anions dissolved in the oil phase. Hence, together with the successive transesterification steps also the catalyst is subjected to change with the reaction starting from CH3O, dissolved in the polar phase, that is transformed in three successive steps to diacylglycerolates, monoacylglycerolates, and glycerolates anions, that is, many different catalytic species are formed like the ones reported in Figure 3. The existence of acylglycerolates anions is postulated by Dijkstra et al.14 However, recently Kouzu et al. have prepared pure calcium glycerolate.18 Glycerolates anions are not soluble in the oil phase and promptly migrate in the polar phase; consequently, the reaction rate rapidly slows down. A mechanism of this type suggests that it is very important to favor the formation of a high interface area in the initial reaction step (reaction 1) in order to consume methanolate anions as soon as possible, and then the reaction prosecutes in the oil phase in chemical regime until the catalytic species, soluble in oil, are present. Moreover, as the FAME yields change along the time from 0 to near 9599%, the composition of the two phases greatly changes; in particular, as diacylglycerolates and monoacylglycerolates have tenside properties, a stable well dispersed methanol/oil emulsion can be formed, i.e., a great interfacial area is formed. Some other physical properties of the mixture change, too, and all these aspects can influence the reaction rate. However, as it is extremely difficult to take into account all the mentioned aspects on the reaction rates, we have neglected them in the 8780



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model as a first approximation averaging the interface area for any adopted fluid dynamic condition. The reaction of methanolate with triglycerides, at the interface, can be schemed as follows CH3 O þ T f D þ E

ð7Þ

(being that D is a diacylglycerolate anion). The rate of this reaction is reasonably proportional to the interface area, the catalyst concentration, and the concentration of T at the interface (this last we assumed proportional to the bulk concentration). Therefore, we can write rCH3 O ¼ k1 3 a 3 ½CH3 O 3 ½T

ð8Þ

where a is the liquidliquid interface area. The effect of reaction 7 is that of transferring the catalyst from the polar to the apolar phase where the transesterification reaction effectively occurs. As the rate of CH3O disappearance depends on the interface area, the limit of this rate will be the one corresponding to the contact of two monomolecular films. This is in agreement with the following observations: (i) kinetic data, reported in the literature, are often in disagreement probably because the stirring rates and consequently the interface area of the reactors are different; (ii) none of the monophasic kinetic models published in the literature can simulate the very high conversions obtained in a few minutes in different efficient mixing devices, because, in those cases the obtainable interface areas are much higher. Another important observation is the dramatic change of the conversion in well mixed continuous reactors obtained when the amount of used catalyst is increased by a factor of 2. These differences are poorly dependent on the residence time, that is, the conversion reached does not increase further with time for reaching a chemical equilibrium. This behavior can be observed also for many experimental batch runs reported in the literature5,6 where by increasing the catalyst concentration an increase of the reaction rate is observed but the final plateau changed with the amount of used catalyst. Therefore, the plateau is clearly a consequence of the catalyst deactivation not of the chemical equilibrium. The deactivation occurs, because the catalyst continues to change during the transesterification reaction according to the following simplified mechanism D þ CH3 OH f M þ E M

þ CH3 OH f G

þ E

ð9Þ

Figure 4. Parity plot obtained by simulating with the UNIFAC LLE method a MeOH/glycerol/FAME system. The experimental data are reported in the literature at 60 °C.20,21 The compositions are expressed in molar fractions.

action probably is due to the equilibrium G þ CH3 OH a CH3 O þ G

CH3O formed for exchange with G reacts again, at the interface, with mono- and diglycerides molecules according to the reactions

being that M = monoacylglycerolates anions and G = glycerolates anions. The corresponding reaction rates are ð11Þ

rM ¼ kM 3 ½M 3 ½CH3 OHoil

ð12Þ

G is not soluble in the oil phase and migrates in the polar phase. G is much less active than D and M that are the true catalysts formed in situ and soluble in the oil phase. Its catalytic

CH3 O þ D f M þ E

ð14Þ

CH3 O þ M f G þ E

ð15Þ

Then, reactions 14 and 15 prosecute more slowly being the equilibrium (13) shifted to the left, according to the experimental observations. It cannot be excluded also the charge exchange between CH3OH and M or G, but the extent of this reactions is probably negligible for the draining effect of the sequence of reactions bringing from methanolate to glycerolate anion. To use more catalyst means to have a greater concentration of the active components in any step of the reaction in the oil phase and also the catalyst deactivation is slowed. The reaction occurs in the oil phase in the three consecutive steps reported in the schemes 4-6, and the related reaction rates can be described in a simplified way as

ð10Þ

rD ¼ kD 3 ½D 3 ½CH3 OHoil

ð13Þ

rT ¼ kT 3 ½D þ M 3 ½T 3 ½CH3 OHoil

ð16Þ

rD ¼ kD 3 ½D þ M 3 ½D 3 ½CH3 OHoil

ð17Þ

rM ¼ kM 3 ½D þ M 3 ½M 3 ½CH3 OHoil

ð18Þ

Methanol is partitioned between the two liquid phases, and obviously only methanol dissolved in the oil phase is involved in the reaction. In some conditions, this last aspect can further slow down the reaction rate as a consequence of mass transfer limitation. The described mechanism is oversimplified, because, as recently reported by Dijkstra et al.,14 diacylglycerolate anion formed in the first catalytic reaction step (7) migrates in apolar 8781



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2Þ CH3 O þ D f M þ E

phase and promotes further transesterification steps with the formation of enolates intermediates

r10 ¼ k1 3 a 3 ½CH3 O p 3 ½Da

ð22Þ

3Þ CH3 O þ M f G þ E r100 ¼ k1 3 a 3 ½CH3 O p 3 ½Ma Then, the enolates react with methanol, dissolved in the apolar phase, to give methylester and a new diacylglycerolate anion

ð23Þ

being that p and a are respectively referred to polar and apolar phase. As it can be seen, the kinetic constants for these three reactions have been assumed the same as a reasonable first approximation. Apolar phase. 4Þ D þ CH3 OH f M þ E r2 ¼ k2 3 ½D a 3 ½CH3 OHa

With the same mechanism, in successive steps are formed also monoglycerides and monoacylglycerolate anions, glycerol, and glycerolate anion. We have individuated at least 12 different anionic species (considering also enolates species) and 12 different reactions. These ionic species are predictable on the basis of the mechanism described in the literature proposed by Dijkstra et al.;14 in particular, they have not been experimentally observed, but their presence can be reasonable from a mechanistic point of view as intermediate or charged species. Clearly, the system needs to be simplified for the kinetic approach. Therefore, on the basis of the described mechanism, we have developed, as a first approximation, a lumped biphasic kinetic model. 3.3. Description of the Biphasic Kinetic Model and Application to Batch Runs Taken by the Literature. As before mentioned, kinetic data collected in well stirred batch reactors, reported in the literature, are often in disagreement. Moreover, these data are normally interpreted with a pseudomonophasic model. The most reliable and recently published kinetic data are the ones reported by Vicente et al.5,6 Therefore, we considered those data for testing the validity of the biphasic model developed by us and determining the corresponding kinetic parameters. Our biphasic kinetic approach considers the following: (i) all the reactions occurring in each phase, polar or apolar, and (ii) the partition of methanol and glycerol among the two phases. Although all the components are partitioned and we can evaluate the composition of the two phases as a function of the conversion19 by using ChemCAD 6.3 software setting the UNIFAC LLE model, we neglected the partition of the components other than methanol and glycerol in order to simplify the model. It is important to mention that in our previous studies we have verified that the UNIFAC LLE model can adequately simulate the experimental data, collected for a MeOH/glycerol/FAME system, that are present in the literature.20,21 For this purpose, in Figure 4, a parity plot that shows the goodness of the UNIFAC LLE predictions referred to the mentioned experimental data, collected by other authors at 60 °C, is reported. Therefore, UNIFAC LLE can be applied in predicting also the phase equilibria for the chemical species involved in the transesterification reaction. (iii) The methanol and glycerol mass transfer have been determined by adopting the Whitman’s two film theory.22 Let us consider, first of all, the adopted reaction scheme and related kinetic equations Interface. 1Þ CH3 O þ T f D þ E r1 ¼ k1 3 a 3 ½CH3 O p 3 ½Ta

ð24Þ

5Þ M þ CH3 OH f G þ E r3 ¼ k3 3 ½M a 3 ½CH3 OHa

ð25Þ

Cat

6Þ T þ CH3 OH sf D þ E r4 ¼ k4 3 ½D þ M 3 ½Ta 3 ½CH3 OHa

ð26Þ

Cat

7Þ D þ CH3 OH sf M þ E r5 ¼ k5 3 ½D þ M 3 ½Da 3 ½CH3 OHa

ð27Þ

Cat

8Þ M þ CH3 OH sf G þ E r6 ¼ k6 3 ½D þ M 3 ½Ma 3 ½CH3 OHa

ð28Þ

Polar phase. 9Þ CH3 OH þ G / CH3 O þ G ( ) 1 p p p p r7 ¼ k7 3 ½CH3 OH 3 ½G 3 ½CH3 O 3 ½G K eq, 7 3 ð29Þ Here, the dependence between the kinetic constant with the temperature is expressed with a modified Arrhenius, whose expression is reported below Eai 1 1 exp ð30Þ ki ¼ kref i 3 R 3 T ref T with T expressed in Kelvin, and Tref is the reference temperature chosen at 298 K. In particular, we have imposed, as a first approximation, three activation energy values that are a Ea1: Activation energy for the reactions 21-23 b Ea2: Activation energy for the reactions 24 and 25 c Ea3: Activation energy for the reactions 26-28 Reaction 29 is always near to the equilibrium. The equations describing the mass transfer of methanol and glycerol from the two phases are as follows: Polar phase. p

JCH3 OH ¼ kl, CH3 OH 3 a 3 ð½CH3 OHp HCH3 OH 3 ½CH3 OHa, Þ

ð21Þ

ð31Þ 8782



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Table 3. List of the Experimental Runs Reported in the Literature,5,6 Submitted to Mathematical Regression run

temperature (°C)

stirring rate (rpm)

KOH (wt.%)

MeOH/oil (mol/mol)

oil

1

25

600

1.0

6:1

Brassica Carinata

2

25

600

1.5

6:1

Sunflower

3

25

600

0.5

6:1

Brassica Carinata

4

25

600

1.5

6:1

Brassica Carinata

5

25

600

0.5

6:1

Sunflower

6

25

600

1.0

6:1

Sunflower

7

35

600

0.5

6:1

Sunflower

Table 4. List of the Parameters Obtained by Mathematical Regression Analysis on the Batch Runs Reported in the Literature parameters

values

units

γ1 = k1a

0.427

k2

0.225

k3

0.225

k4

18.16

k5

4.61

k6 Keq,7

1.92 1.69 e-3

L2,a/(mol2 3 min) -

βCH3OH

1.52

min1

p

L /(mol 3 min) La/(mol 3 min) La/(mol 3 min)

L2,a/(mol2 3 min) L2,a/(mol2 3 min)

Ea1

≈0

kcal/mol

Ea2

17.61

kcal/mol

Ea3

26.62

kcal/mol

p

JG ¼ kl, G 3 a 3 ð½Gp HG 3 ½Ga, Þ

ð32Þ

JG ¼ kl, G 3 a 3 ð½G p HG 3 ½G a, Þ p

Figure 5. Comparison between the FAME concentration obtained in the same experimental conditions (25 °C and 1.5 wt.% KOH referred to the oil) but using two different oils, that are Brassica Carinata and Sunflower oil. Taken from Vicente et al.5,6

ð33Þ

model we can determine the evolution of the solubility parameters with the triglycerides conversion (x). The following optimized polynomial equations have been obtained

Apolar phase. a JCH ¼ kl, CH3 OH 3 a 3 ð½CH3 OH 3 OH

a,

a

½CH3 OH Þ

JGa ¼ kl, G 3 a 3 ð½Ga, ½Ga Þ

ð34Þ ð35Þ

JGa ¼ kl, G 3 a 3 ð½G a, ½G a Þ

ð36Þ

The interfacial concentration (*) can be calculated solving the following balance that, considering a steady-state condition, is always valid for the J component p JJ

3V ¼ p

JJa

3V

By solving the previous equation it is possible to obtain

½Ja, ¼

p

v 3 βj 3 ½J þ βj 3 V 3 ½J βj 3 V a þ vp 3 βj 3 HA p

½Jp, ¼ HJ 3 ½Ja,

a

ð40Þ

1=HG ¼ 6304:75 15430:10 3 x 1643:71 3 x2 þ 64912:73 3 x3 92127:03 3 x4 þ 43103:45 3 x5

ð41Þ

It is now possible to write the mass balances for both the polar and the apolar phase: Polar phase.

ð37Þ

a

1=HCH3 OH ¼ 25:85 23:00 3 x 11:21 3 x2 þ 36:41 3 x3

a

ð38Þ ð39Þ

being that βj = kl,ja. For solving the mass transfer equations it is necessary to define also the solubility parameters Hj = [J]a/[J]p for both methanol and glycerol. By using ChemCAD V. 6.3 and the UNIFAC LLE 8783

d½CH3 OHp p ¼ JCH3 OH r7 dt

ð42Þ

d½Gp p ¼ þ JG þ r7 dt

ð43Þ

d½CH3 O p Va ¼ ðr1 þ r10 þ r100 Þ 3 þ r7 dt Vp

ð44Þ

d½G p p ¼ þ JG r7 dt

ð45Þ

½CH3 O p, 0 ¼ ½KOHp

ð46Þ



ARTICLE

Figure 7. Simulations of three experimental runs taken by the literature at different catalyst concentration, that correspond to Runs 1, 3, and 4 of Table 3. Symbols represent the experimental points, while lines represent the simulations.

Figure 6. Simulations of two experimental runs taken by the literature (Vicente et al.5,6): (a) Run 1 of Table 3 and (b) Run 7 of Table 3. Symbols represent the experimental points, while lines represent the simulations.

Apolar phase. d½Ta ¼ r1 r4 dt

ð47Þ

d½Da ¼ þ r4 r5 r10 dt

ð48Þ

d½Ma ¼ þ r5 r6 r100 dt

ð49Þ

d½Ga ¼ þ r6 JGa dt

ð50Þ

d½Ea ¼ þ r1 þ r10 þ r100 þ r2 þ r3 þ r4 þ r5 þ r6 dt

ð51Þ d½CH3 OHa a ¼ r2 r3 r4 r5 r6 þ JCH 3 OH dt

ð52Þ

d½D a ¼ þ r1 r2 dt

ð53Þ

d½M a ¼ þ r2 r3 þ r10 dt

ð54Þ

d½G a ¼ þ r3 JGa þ r100 ð55Þ dt This model has been applied, first of all, to interpret the kinetic runs performed by Vicente et al.5,6 in a batch reactor. The considered experimental data are the ones of the runs listed in Table 3. These data have been submitted to mathematical regression analysis, and the kinetic parameters of the model giving the best fitting are reported in Table 4. By observing the list of the runs, reported in Table 3, some of these runs have been performed with Sunflower oil, some others with Brassica Carinata oil. We have observed that the two oils have exactly the same kinetic behavior as it can be seen in Figure 5. So, we have considered all the runs together not distinguishing one oil from the other. In Figure 6a,b two examples of simulations made with the biphasic model are reported. As it can be seen, the obtained agreement is very good, but the most important agreement is the one obtained for the runs performed with different amounts of catalysts (Figure 7). In this case the biphasic model is able to describe not only the difference in the initial reaction rates but also the final plateau, while the pseudomonophasic model cannot reproduce this behavior giving the same plateau, corresponding to the chemical equilibrium that is the same in all the cases because the runs have been performed at the same temperature of 25 °C. The difference observed for the final plateau by using a different amount of catalyst are explained by the dynamic behavior of the biphasic system: the initial rate is much higher and the concentration of the catalyst formed in the oil is greater. The plateau, therefore, is the consequence of the reaction rate leveling and not of the chemical equilibrium. The activation energy near to zero found for Ea1 can be accepted being reaction 7 a fast ionic exchange reaction, while reactions 9 and 10 are only apparently ionic exchange reactions, because we have lumped the kinetic model. In those cases the reactions pass through the formation of enolates and reaction 8784



ARTICLE

Table 5. List of the Parameters Used for Simulating the Continuous Runs parameter γ10 βCH3OH

0

RP1

RP21

RP22

RWTR

TRR

4.73

25.84

22.84

72.00

13.11

39.93

2.09

0.29

3.57

6.44

Table 7. Experimental Runs Performed in the Described Continuous Reactors, at 60 °C, by Using 2.0 wt.% KOH Catalyst

unit cm /(mol 3 min) min1

Table 6. Experimental Runs Performed in the Described Continuous Reactors, at 60 °C, by Using 1.0 wt.% KOH Catalyst Qoil run

reactor

QMeOH ‑1

3

(cm min )

3

‑1

(cm min )

Qoil

QMeOH

τ

YEXP

YSIM

(cm3min‑1)

(cm3min‑1)

(min)

(%)

(%)

RP21

28.62

7.382

0.17

93.60

91.30

RP21

19.24

4.962

0.25

93.80

90.70

29

RP21

9.62

2.480

0.50

94.70

90.90

30 31

RP21 RP21

4.85 2.38

1.251 0.615

0.99 2.02

98.20 97.10

92.10 90.10

32

RWTR

3.70

0.950

3.00

97.00

88.80

33

RWTR

2.77

0.710

4.00

93.70

87.00

34

RWTR

2.21

0.570

5.01

85.20

85.20

3,p

τ

YEXP

YSIM

(min)

(%)

(%)

1

RP1

9.80

2.530

0.58

62.50

84.87

2

RP1

6.60

1.730

0.86

79.30

81.88

3

RP1

5.50

1.400

1.04

80.70

80.61

4

RP1

4.40

1.150

1.29

81.80

78.61

5 6

RP1 RP1

1.10 2.40

0.287 0.615

5.16 2.37

73.30 73.30

64.73 73.21

7

RP1

0.58

0.150

9.79

43.60

56.96

8

RP21

28.62

7.382

0.17

83.60

83.50

9

RP21

19.24

4.962

0.25

81.10

81.00

10

RP21

9.62

2.480

0.50

80.30

80.20

11

RP21

4.85

1.251

0.99

81.90

81.80

12

RP21

2.38

0.615

2.02

78.70

78.60

13 14

RP22 RP22

9.80 7.90

2.530 2.040

0.31 0.38

80.50 81.50

85.10 85.20

15

RP22

6.60

1.730

0.46

78.60

84.08

16

RP22

5.50

1.400

0.55

77.20

83.22

17

RP22

4.40

1.150

0.68

81.50

81.60

18

RP22

3.20

0.830

0.94

80.60

80.08

19

RP22

2.40

0.615

1.26

82.30

78.55

20

RP22

0.58

0.150

5.21

80.50

71.80

21 22

RWTR RWTR

3.69 2.77

0.950 0.710

3.00 4.00

76.90 74.90

77.20 74.90

23

RWTR

2.21

0.570

5.01

73.60

73.20

24

TRR

5.50

1.400

1.09

70.44

70.44

25

TRR

6.60

1.730

0.91

75.59

75.59

26

TRR

9.80

2.526

0.61

72.87

72.87

run

reactor

27 28

Figure 8. Runs performed with the reactors RP1 and RP22 by using 1% of KOH catalyst.

p

d½Gp þJ þ r7 ¼ G dV Q p

d½CH3 O ¼ dV

with methanol (reactions 19 and 20), and the rate determining step is probably an activated intermediate reaction of the overall mechanism. 3.4. Biphasic Kinetic Model Applied to Continuous Runs. By operating with continuous packed bed tubular reactors we observed two important phenomena: (i) the final conversion is strongly affected by the initial interface area and (ii) the conversion is strongly affected by the initial catalyst concentration. In some cases the conversion increases by decreasing the residence time until a maximum value is reached; in other cases the conversion remains relatively constant with the residence time. In order to simulate the continuous runs the following mass balance equations must be solved: Polar phase.

ðr1 þ r10 þ r100 Þ 3

d½G p þJ r7 ¼ G dV Q

Q

Qa r7 Qp

ð58Þ

p

ð59Þ

Apolar phase.

p

JCH3 OH r7 d½CH3 OHp ¼ dV Q

ð57Þ

ð56Þ 8785

d½Ta r1 r4 ¼ dV Q

ð60Þ

d½Da þr4 r5 r10 ¼ dV Q

ð61Þ

d½Ma þr5 r6 r100 ¼ dV Q

ð62Þ



ARTICLE

Figure 9. Parity plot obtained for the runs performed with the reactors RP21, RWTR, and TRR.

d½Ga þr6 JGa ¼ dV Q

ð63Þ

d½Ea þr1 þ r10 þ r100 þ r2 þ r3 þ r4 þ r5 þ r6 ¼ dV Q ð64Þ a r2 r3 r4 r5 r6 þ JCH d½CH3 OHa 3 OH ¼ dV Q

ð65Þ

d½D a þr1 r2 ¼ dV Q

ð66Þ

d½M a þr2 r3 þ r10 ¼ dV Q

ð67Þ

d½G a þr3 JGa þ r100 ¼ dV Q

ð68Þ

In order to simulate simultaneously experimental runs performed at different flow rates, we have assumed that both γ1 = k1a and βCH3OH = kl,CH3OHa values depend linearly on the overall volumetric flow rate, being that these parameters are both dependent on the interfacial area γ1 ¼ γ1 0 3 Q =Q ref

ð69Þ

βCH3 OH ¼ β0CH3 OH 3 Q =Q ref

ð70Þ

with Qref = 1.0 cm3/min. All the other equations, of mass transfer, partition, reaction rate laws, and the equations of congruence are the same as reported in section 3.3. The agreements in simulating the continuous runs have been obtained by using the same kinetic parameters, derived from batch runs, reported in Table 4. For simulating all the continuous runs, it is necessary to adjust only two parameters, that are as follows: γ10, correlated with 0 , the interface area (relation 69), and eventually βCH 3 OH correlated with the mass transfer coefficient for methanol

and again with the interface area (relation 70). In Table 5 are reported γ10 and β0CH3OH. In Table 6, all the experimental conditions adopted and the comparison between the experimental and calculated yields are reported for the runs performed in the different reactors, at 60 °C, by using 1% of catalyst. On the contrary, in Table 7 are reported all the experimental conditions adopted and the comparison between the experimental and calculated yields for the runs performed in the different reactors, at 60 °C by using 2% of catalyst. As it can be seen the agreement found is satisfactory considering all the simplifications introduced in the model. In the continuous runs the difference in the conversions observed with respectively 1 or 2% of catalyst are much more pronounced than in the batch conditions, but the reason for this difference is the same. The obtained agreement can be better appreciated in Figure 8 reporting the trends of the yield with the residence time for the reactors RP1 and RP22 corresponding to the runs reported in Table 6. At last, in Figure 9 a parity plot related to the experimental runs performed with the reactors RP21, RWTR, and TRR (see conditions in Tables 6 and 7). As it can be seen, the parity plot shows an error inferior to 10%. Moreover, it seems that a systematic error affects the simulations at the highest level of yield. This means that the model needs to be further improved.

4. CONCLUSIONS In the present work, the effect of different types of static mixers on the intensification of the biodiesel synthesis has been deeply investigated. On the basis of the experimental results it is possible to conclude that an intense local micromixing, favoring the development of a great interface area, gives place to very high yield in mild conditions (60 °C) and very short residence times (less than 1 min). Another important experimental observation is that by increasing the catalyst concentration the yield increases very much in the same conditions of temperature and residence time. This means that the maximum yields obtained cannot correspond to equilibrium values depending on the catalyst concentration. On the other hand, in the literature are reported many runs, performed in batch conditions, in which the final conversion increases with the catalyst concentration. It is possible to conclude that the monophasic kinetic models, normally adopted in the literature, considering the transesterification steps as equilibrium reactions, are not adequate to describe in any detail biodiesel production. To do this requires the development and adoption of a new biphasic model based on a more reliable reaction mechanism. A model of this type has been described in this work and successfully applied to both batch runs, taken from the literature, and continuous runs in tubular reactor filled with different static mixer elements. Further improvements about the change in the interfacial area as conversion proceeds and about the changes in physical properties as the reaction proceeds could be the challenge for a future investigation in order to further improve the biphasic model. ’ AUTHOR INFORMATION Corresponding Author

*E-mail: [email protected]. 8786



’ ACKNOWLEDGMENT Thanks are due to EC VII Framework Programme CP-IP 228853-2 COPIRIDE and to MIPAF (Italian Ministry of Agricultural, Food and Forest Policies) “Project AGROPROM New technologies for the production of biodiesel from waste oil and fats sources” D.M. 246/2007 (23/10/2007) and 16912/ 7303/10 (23/7/2010) for the financial support. ’ LIST OF SYMBOLS A1 methoxylic groups single signal at δ = 3.7 ppm A2 methylenic groups triplet signal δ = 2.3 ppm dispersion number ND Pe Peclet number, defined as the reverse of the dispersion number D axial dispersion coefficient (m2/s) u axial velocity (m/s) L reactor’s length (m) F dimensionless tracer concentration (C/C0) where C and C0 are respectively the actual and the initial concentrations Θ dimensionless time defined as t/tm, with tm being the residence time TRR reactor packed with 8 stainless steel threaded rods: length = 20 cm, diameter = 3 mm, thread = 0.5 mm RP1 single size packing filled with spheres of 2.5 mm diameter RP21 dual size packing filled with spheres of two different diameters: 2.5 and 1.0 mm RP22 dual size packing filled with spheres of two different diameters: 2.5 and 0.39 mm RWTR reactor packed with stainless steel wool T triglycerides D diglycerides M monoglycerides G glycerol E methylesters diacylglycerolate anion D monoacylglycerolate anion M glycerolate anion G reaction rate (mol L1 min1) ri kinetic constant (units depend on the reaction) ki Cat catalyst a interfacial area (cm2/cm3) kinetic parameter of reaction defined as k1 3 a γ1 activation energy (kcal/mol) Eai equilibrium constant Keq,i mass transfer rate for the j component in the phase Jji i (mol L1 min1) mass transfer coefficient of the j component (cm/min) kl,j mass transfer parameter of the j component defined as βj kl,j 3 a (min1) partition coefficient of the j component ([j]a/[j]p) Hj x conversion degree (-) Q overall volumetric flow rate (cm3/min) τ residence time (minutes) V reactor’s volume (cm3)

ARTICLE

’ REFERENCES (1) Freedman, B.; Butterfield, R. O.; Pryde, E. H. Transesterification Kinetics of Soybean Oil. J. Am. Oil Chem. Soc. 1986, 63, 1375–1380. (2) Noureddini, H.; Zhu, D. Kinetics of Transesterification of Soybean Oil. J. Am. Oil Chem. Soc. 1997, 74, 1457–1463. (3) Darnoko, D.; Cheryan, M. Kinetics of Palm Oil Transesterification in a Batch Reactor. J. Am. Oil Chem. Soc. 2000, 77, 1263–1267. (4) Komers, K.; Skopal, F.; Stloukal, R.; Machek, J. Kinetics and mechanism of the KOH-catalyzed methanolysis of rapeseed oil for biodiesel production. Eur. J. Lipid Sci. Technol. 2002, 104, 728–737. (5) Vicente, G.; Martinez, M.; Aracil, J.; Esteban, A. Kinetics of sunflower oil methanolysis. Ind. Eng. Chem. Res. 2005, 44, 5447–5454. (6) Vicente, G.; Martinez, M.; Aracil, J. Kinetics of Brassica carinata oil methanolysis. Energy Fuels 2006, 20, 1722–1726. (7) Stamenkovic, O. S.; Todorovic, Z. B.; Lazic, M. L.; Veljkovic, V. B.; Skala, D. U. Kinetics of sunflower oil methanolysis at low temperatures. Bioresour. Technol. 2008, 99, 1131–1140. (8) Santacesaria, E.; Di Serio, M.; Tesser, R.; Casale, L.; Verde, D.; Turco, R.; Bertola, A. Use of a Corrugated Plates Heat Exchanger Reactor for Obtaining Biodiesel with Very High Productivity. Energy Fuels 2009, 23, 5206–5212. (9) Santacesaria, E.; Di Serio, M.; Tesser, R.; Turco, R.; Verde, D.; Casale, L. PCT Application assigned to MYTHEN SpA Co WO 2010/ 073216. (10) Lowe, H.; Ehrfeld, W. State-of-the-art in microreaction technology: concepts, manufacturing and applications. Electrochim. Acta 1999, 44, 3679–3689. (11) Hessel, V.; Hardt, S.; Lowe, H. Chemical Micro Process Engineering: Wiley-VCH: Weinheim, Germany; 2004. (12) Hessel, V.; Renken, A.; Shouten, J. T.; Micro, J. Process Engineering; Wiley-VCH: Weinheim, Germany; 2009. (13) Qiu, Z.; Zhao, L.; Weatherley, L. Process intensification technologies in continuous biodiesel production. Chem. Eng. Process. 2010, 49, 323–330. 00 00 (14) Dijkstra, A. J.; Toke, E. R.; Kolonits, P.; Recseg, K.; Kov ari, K.; Poppe, L. The base-catalyzed, low- temperature interesterification mechanism revisited. Eur. J. Lipid Sci. Technol. 2005, 107, 912–921. (15) Santacesaria, E.; Di Serio, M.; Tesser, R.; Tortorelli, M.; Turco, R.; Russo, V.; A simple device to test biodiesel process intensification. Chemical Engineering and Processing: Process Intensification; doi:10.1016/ j.cep.2011.05.023 13. (16) Gelbard, G.; Bres, O.; Vargas, R. M.; Vielfaure, F.; Schuchardt, U. F. 1H nuclear magnetic resonance determination of the yield of the transesterification of rapeseed oil with methanol. J. Am. Oil Chem. Soc. 1995, 72, 1239–1241. (17) Nauman, E. B. Chemical reactor design, optimization, and scaleup; McGraw-Hill: 2001. (18) Kouzu, M.; Hidaka, J. S.; Wakabayashi, K.; Tsunomori, M. Solid base catalysis of calcium glyceroxide for a reaction to convert vegetable oil into its methyl esters. Appl. Catal., A 2010, 390, 11–18. (19) Chemical Process Simulator ChemCAD 6.2 by Chemstations, 2009. (20) Zhou, H.; Lu, H.; Liang, B. Solubility of Multicomponent Systems in the Biodiesel Production by Transesterification of Jatropha curcas L. Oil with Methanol. J. Chem. Eng. Data 2006, 51, 1130–1135. (21) Oliveira, M. B.; Teles, A. R. R.; Queimada, A. J.; Coutinho, J. A. P. Phase equilibria of glycerol containing systems and their description with the Cubic-Plus-Association (CPA) Equation of State. Fluid Phase Equilib. 2009, 280, 22–29. (22) Whitman, W. G. Preliminary Experimental Confirmation of the Two-Film Theory of Gas Absorption. Chem. Metall. Eng. 1923, 29, 146–148.

Subscripts and Superscripts

oil p a

oil phase polar phase apolar phase 8787


Biodiesel Process Intensification by Using Static Mixers Tubular

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