A simple device to test biodiesel process intensification

Chemical Engineering and Processing 50 (2011) 1085–1094

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Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep

A simple device to test biodiesel process intensification E. Santacesaria ∗ , M. Di Serio, R. Tesser, M. Tortorelli, R. Turco, V. Russo Department of Chemistry of the University of Naples “Federico II”, Complesso Universitario Monte S. Angelo, Via Cintia 4, IT 80126 Naples, Italy

a r t i c l e

i n f o

Article history: Received 25 January 2011 Received in revised form 11 April 2011 Accepted 31 May 2011 Available online 7 July 2011 Written in honour of Prof. Albert Renken to celebrate his 70th birthday. Keywords: Biodiesel Transesterification Process intensification Microreactor Static mixer Kinetics

a b s t r a c t In this paper, we have studied the KOH catalyzed transesterification reaction of vegetable oil with methanol in a tubular reactor filled with small spheres of stainless steel of different sizes. Three different packed bed configurations have been tested corresponding to different fluid dynamic situations. The first configuration corresponds to the tubular reactor filled with spheres of 2.5 mm of diameters; in the second configuration an opportune amount of spheres of 1 mm has been added to the mentioned spheres of 2.5 mm for filling the void volume of the octahedral cavities between the bigger spheres; in the third configuration an opportune amount of spheres of 0.39 mm has been added to spheres of 2.5 mm for filling the void volume of the tetrahedral cavities between the bigger spheres. The three mentioned configurations give place to the formation of micro-channels with an approximated size of respectively 1000 ␮m, 500 ␮m and 300 ␮m. These systems, subjected to fluid dynamic characterization, have shown a very high local turbulence (static mixer), in particular, when a packed bed reactor with dual size packing is used. Then, kinetic transesterification runs have been made by using the three mentioned packed bed reactors and a very high productivity has been obtained as a consequence of the induced local micro-mixing. A simplified kinetic model has been developed, which is able to describe many runs in batch conditions reported by the literature. This model resulted unsuitable to simulate the continuous runs performed in the described packed bed reactors. Our conclusion is that monophasic models, often proposed in the literature, are not able to describe the kinetic behavior of KOH catalysed transesterification in microchannels devices. At last, the described microchannels device represents, as it will be discussed, an ideal connection between a traditional tubular packed bed reactor and the recently appeared microreactors that are very efficient in mass and energy transfer operations for process intensification. © 2011 Elsevier B.V. All rights reserved.

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 and atmospheric pressure. A lot of papers dealing with the kinetics of the transesterification reaction have been published in the last years but data reported are not in agreement and often are clearly affected by mass transfer limitation [1–7]. This occurs, because, the reagents, in the mentioned conditions, are immiscible liquids and the interphase surface area plays an important role in this reaction. Moreover, most of the published works report experimental data obtained in batch reactors. In a recently published paper and patent [8,9] we have shown that a very high productivity (420 kg/day) can surprisingly be obtained by performing the transesterification reaction in a continuous corrugated plates heat exchanger reactor (CP-HEX reactor) in a range of temperatures between 60 and 100 ◦ C. We attributed the high performances of the used CP-HEX reactor to

∗ Corresponding author. Tel.: +39 081674027; fax: +39 081674026. E-mail address: [email protected] (E. Santacesaria). 0255-2701/$ – see front matter © 2011 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2011.05.023

a very active local micro-mixing occurring when high liquid flow rates were fed in such type of reactors. Starting from these findings, we imagined the possibility to enhance the mass-transfer rate in the methanol–oil transesterification reaction also by using other kinds of reactors favoring the development of a high liquid–liquid interphase such as, for example, microreactors or static mixers. It is well known that, the scale of the system strongly affects the gradients and the driving forces for mass, heat and momentum transfer. As the size is decreased the surface/volume ratio increases and consequently the efficiency in transport phenomena increases, too. This concept has been intensively deepened in recent studies and dedicated books and nowadays it is well known that microsystems having micro-channels of size smaller than 300 ␮m strongly improve mass, heat and momentum transfer [10–13]. For the explained reasons the use of microreactors is often proposed for the process intensification (PI) of two-phase reactions occurring at liquid–liquid interface (1-pentene epoxydation [14]). At this purpose, the use of microreactors has successfully been proposed for promoting the transesterification reaction [15]. Considering that microreactors are powerful and economic systems for the intensification of industrial processes, at the scope of better modeling these systems it would be useful to develop ideal reactors systems able

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to give, in perspective, a response about the effectiveness of microchannels for a given two phase liquid–liquid reaction. In a recently published paper [16] we proposed the use of a micro-channel device to test the process intensification effects on a reaction. This device was simply a tubular reactor filled with stainless steel spheres of different sizes in which the smaller spheres (0.39 mm of diameter) were perfectly allocated in the void volume existing in the framework created by the bigger ones (2.5 mm of diameter). In this way, micro-channels with a size of about 300 ␮m were created and the device can be assimilated, for the fluid dynamic behavior and mass transfer performances to a microreactor. The proposed device ideally represents a connection between a traditional tubular packed bed reactor and the recently appeared microreactors having in common the micro-channel structure. However, we have shown, by using a CP-HEX reactor, that for liquid–liquid biphasic reactions, like methanol–soybean oil transesterification, mass transfer rate can greatly be increased by favoring an intense local turbulence. This effect can also be obtained by using a static mixer system. Many different devices of this type have been proposed by the literature, in particular Prof. Albert Renken was a pioneer of this research [17]. However, it is not simple to decide what type of device could be more suitable for a given reaction. Again an ideal system could be useful for orienting the choice and in the present work we propose the use of three different tubular reactors respectively containing spheres of uniform size (2.5 mm of diameter) having micro-channels of about 1000 ␮m, spheres of two different sizes of respectively 2.5 and 1 mm of diameters having microchannels of about 500 ␮m and two different sizes of respectively 2.5 and 0.39 mm of diameters having micro-channels of about 300 ␮m, for studying the methanol–soybean oil transesterification reaction. The fluid dynamic behaviors of all these mentioned reactors have been studied by evaluating the Residence Time Distributions (RTD) in stimulus–response step test. Then, all the three reactors have been used for studying the methanol–soybean oil transesterification reaction to biodiesel for different flow rates. Surprisingly, in some cases, the performances obtained were so high that none of the kinetic models proposed by the literature, derived from runs performed in batch conditions, resulted adequate to describe the observed kinetic behavior. For this reason, all the aspects related to both kinetics and mass transfer for methanol–soybean oil transesterification will be deepened and discussed in this paper. 2. Experimental 2.1. Reagents, methods and apparatus 2.1.1. Reagents Soybean oil, used as vegetable oil, was purchased in a local food-store (the fatty acid composition of the used soybean oil, determined by gas chromatographic analysis, was (%, w/w): 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. 2.1.2. Analytical method The FAME (fatty acid methyl esters) yields were determined by using 1 H NMR technique (GEMINI 200 Mz), measuring the area of the H NMR signal related to methoxylic (A1, single signal at ı = 3.7 ppm) and methylenic groups (A2, triplet signal ı = 2.3 ppm), respectively. The FAME yields can be calculated by using the following equation: yieldFAME =

 2 · A1  3 · A2

· 100

The 1 H NMR spectra were obtained with GEMINI-200 equipment in deuterated chloroform [18]. In all cases the conversion values were confirmed by gas-chromatographic 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 and 0.1 ␮m film). Before the analysis, the samples were derivatizated by BSTFA (N,Obis(trimethylsilyl) trifluoroacetamide with trimethylchlorosilane). As internal standards, methyl-heptadecanoate was used for methyl ester and 1,2,4-butantriol for glycerol. The fluid dynamic tests have been performed analyzing the samples by UV–visible measurements.

2.1.3. Reactor setup The used reactor was a stainless steel AISI 316 cylindrical tube of 20 cm length and 12.7 mm (1/2 ) of diameter with internal diameter 10 mm. The internal void volume of the reactor was 15.7 cm3 . Three different kinds of filling composed by stainless steel AISI 316 spheres were used: (1) a single size packing (RX1) in which the reactor was filled with spheres of 2.5 mm diameter; (2) a dual size packing (RX2) in which spheres of two different diameters (2.5 and 1.0 mm respectively) were used; (3) again, a dual size packing (RX3) in which spheres of two different diameters (2.5 and 0.39 mm respectively) were used. A picture of the three different packing configuration is reported in Fig. 1 together with a photo suggesting different random packing ensembles for the monodisperse system RX1. The first configuration RX1 is composed by 1050 spheres that completely fill the tubular reactor. The hexagonal packing of the bigger spheres RX1 leads to two different kind 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, then 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. Therefore, the RX1 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 (see structure RX3). In the case of the RX2 structure the introduction of spheres of 1.0 mm of diameter forces the bigger spheres to assume a new geometrical cubic ensemble as the corresponding structure reported in Fig. 1. An estimation of the dimensions of the microchannels has been made considering an ideal packing, so obtaining an average value of about 1000 ␮m for the structure RX1, 500 ␮m for the RX2 and 300 ␮m for the RX3. For filling the reactor 1050 spheres of 2.5 mm diameter were used for the RX1 reactor, 888 spheres of 2.5 mm diameter and 4560 of spheres of 1.0 mm for the RX2 reactor, 1050 spheres of 2.5 mm diameter and 101800 spheres of 0.39 mm diameter for the RX3 reactor. A stainless steel wool has been put both at the inlet and the outlet of the reactor to avoid the loss of the smaller spheres drag by the liquid flow. The theory of monodisperse packing sphere is a very old problem faced for the first time by Kepler in 1611. Kepler estimated for spheres of the same diameter a maximum packing density of 0.74 corresponding to a void fraction of 0.26. More later, different approaches have been formulated defining a “Random Close Packing” for mono-modal system, like the RX1 reactor, corresponding to a minimum void fraction of ε = 0.36 (ε = void volume/total volume). However, there are in the literature numerous experiments showing that the loosest way of packing spheres (Random Loose Packing) gives a void fraction of 0.45 [19]. It is then possible to estimate a theoretical minimum value of the void fraction for RX2 corresponding to 0.39. This value is greater than the theoretical one of monodisperse RX1 system,

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Fig. 1. Three different geometrical ensembles of spheres filling the used tubular reactor and a picture of different types of packing.

because, the size of the smaller sphere has been chosen in a way to constrain the larger sphere to assume a different type of packing, as it can be seen in Fig. 1. In the case of the bidisperse RX3 system, a minimum theoretical void fraction of 0.25 has been estimated. As it will be seen, these values have been determined also experimentally and the obtained results are in a reasonable agreement for RX1, considering a random loose packing, and in the other cases considering the minimum theoretical void fractions. 2.1.4. Methanol–soybean oil transesterification runs The described reactors have then been used to perform transesterification reaction in a laboratory plant schemed in Fig. 2. Kinetics runs have been made by feeding to the reactor independently, methanol containing the dissolved catalyst (KOH, normally 1 wt% with respect to the oil) and a refined soybean oil. The oil was preheated at the reaction temperature, whereas methanol was fed at room temperature. 2.1.5. Fluid dynamic characterization tests All the 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 a flowing

medium and cobalt nitrate (II) hexahydrate, having a characteristic pink color, has been used in the test as a tracer. 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 response to the step function. Then, the samples have been analyzed by an UV–visible spectrophotometer, monitoring the absorbance of cobalt nitrate at 511 nm. The tests have been repeated for each reactor at different flow rates. An HPLC pump was used for regulating the cobalt solution flow rate.

3. Results and discussion 3.1. Results of the fluid dynamic tests In order to evaluate accurately the void volume of the systems under examination and to study the fluid dynamics of the three described reactors, some step tests have been performed. By reading the flex of the sigmoidal response function that best fit the experimental data it was possible to determine the mean residence time and from these values, to estimate the empty volumes

Fig. 2. Scheme of the laboratory plants used for performing transesterification runs. (1)–(2) are HPLC pumps for respectively methanol and oil feeding, (3) tubular reactor with the electric resistance, (4) check valve for preventing the back flow of oil towards the methanol pump, (5) T mixer, (6) condenser filled with Rashig rings, (7) oil pre-heater, (8) collecting tank, (9) withdrawing samples. The reactor and the pre-heater are heated with an electric resistance and insulated with rock wool.

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D/uL=0

1.0

D/uL=0.033

D/uL=0.062

0.8

0.8

D/uL=0.08

D/uL=0.1 0.6

F(t)=C/C 0

F(t)=C/C0

D/uL=0

1.0

0.4

1.4 cm3/min 3 5.6 cm /min

0.2

0.6

0.4

1.4 cm3/min 5.6 cm3/min

0.2

0.0 0.0

0.5

1.0

1.5

2.0

2.5

3.0

8.3 cm3/min

3.5

0.0

θ = t/tm

0.0

Fig. 3. F function plot against the dimensionless time for the RX1 fluid dynamic tests.

of the three packing filling the reactors. The values obtained were 7.15 cm3 for the RX1 packing, 6.05 cm3 for the RX2 and 3.5 cm3 for the RX3. The corresponding void fractions experimentally determined were: 0.45 for the RX1 packing, 0.38 for the RX2 and 0.22 for the RX3. The dead volume at the exit of the reactor was 0.32 cm3 that is negligible with respect to the reactor void volume. The results of this characterization for the reactor RX1 are shown in Fig. 3, where the dimensionless tracer concentrations F (C/C0 ), are plotted as a function of the dimensionless time  (t/tm , where tm is the mean residence time) for two different volumetric flow rates of respectively 1.4 and 5.6 cm3 /min. C0 is the inlet solution concentration and C is the outlet measured concentration. The results of fluid dynamic characterization for the reactor RX2 are shown in Fig. 4 for two different flow rates that are respectively: 5.6 and 8.3 cm3 /min. The results of fluid dynamic characterization for the reactor RX3 are shown in Fig. 5 for three different flow rates that are respectively: 1.4, 5.6 and 8.3 cm3 /min. Then, the fluid dynamic experimental data can be elaborated by mathematical regression analysis using a single parameter function; the parameter is known as the axial dispersion number (ND ),

D/uL=0

1.0

D/uL=0.0051

D/uL=0.036

F(t)=C/C 0

0.8

0.6

0.4

5.6 cm3/min 8.3 cm3/min

0.2

0.0 0.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

4.0

θ =t/tm Fig. 4. F function plot against the dimensionless time for the RX2 fluid dynamic tests.

0.5

1.0

1.5

2.0

2.5

3.0

3.5

4.0

4.5

5.0

θ =t/tm Fig. 5. F function plot against the dimensionless time for the RX3 fluid dynamic tests.

normally indicated as the reciprocal of the Peclet number (Pe), defined by the following expression. ND =

1 D = Pe u·L

(1)

where u is the fluid linear velocity (m/s) calculated from the volumetric flow rate and the section of the empty reactor, D is the axial dispersion coefficient (m2 /s) and L the length of reactor (m). It is possible to estimate the ND value from the slope of the curve F(t) for  = 1, according to the following relationship [20].

 dF(t)  d(t/tm )

1 = · 2

t/tm =1



uL D

(2)

The integrated function F(t) is reported below.



1 F(t) = · 2



1 − erf

1 · 2





uL · D

1−



t tm

 (3)

t/tm

where erf is the error function defined as following: 2 erf (x) = √ · 



x

2

e−y dy

(4)

0

For the first configuration RX1 the ND numbers resulted respectively 0.062 and 0.033, in correspondence of 1.4 and 5.6 cm3 /min, meanwhile, for RX2 the values are respectively 0.051 and 0.036 in correspondence of 5.6 and 8.3 cm3 /min, while, for RX3 are 0.08, 0.1 and 0.1 in correspondence of 1.4, 5.6 and 8.3 cm3 /min. The literature findings confirm that the ND values estimated for the three reactors are very close to the ones expected for the static mixers that are 1/ND = Pe = 70 L [21] (where L stands for the length of the reactor expressed in meters): in this way, it is possible to calculate a ND = 0.07 for a reactor of 0.2 m length, value that is very similar to the ones obtained by mathematical regression on the experimental data collected in this work. It is then possible to calculate the Reynolds numbers for the three different examined systems (values summarized in Table 1). The relations of the Reynolds numbers for, respectively, an empty tube and the same containing spherical particles are reported below: Re =

·u·d , 

Rep =

 · up · dp 

(5)

where  is the water density at 25 ◦ C (995 kg/m3 ), u the fluid velocity (m/s) calculated from the volumetric flow-rate and the

E. Santacesaria et al. / Chemical Engineering and Processing 50 (2011) 1085–1094 Table 1 Reynolds numbers calculated starting from the volumetric flow-rates imposed during the experimental procedure by using RX1, RX2 and RX3 reactors. The Reynolds numbers calculated for the empty tube have been reported for comparison purpose. Re

Rep RX1

Rep RX2

Rep RX3

1.4 5.6 8.3

3.21 12.82 19.01

1.78 7.18 –

0.55 2.20 3.26

– 4.15 6.14

reactor’s section without the packing, up is fluid velocity (m/s) calculated from the volumetric flow-rate and the empty section of the packed bed reactor, d is the reactor’s diameter (m),  the water viscosity at 25 ◦ C (9.225 × 10−4 Pa s) and dp is the particle’s diameter calculated as a weighted average value in the case of RX2 and RX3. As it is possible to observe from the data reported in Table 1, with the used flow rates the particles Reynolds numbers are quite low notwithstanding the axial dispersion numbers (ND ) approaches that of a plug flow behavior. On the other hand, the values reported in Table 1 are in good agreement with the ones reported in the literature related to the behavior of the static mixers, that are in a range of Re < 100. For this reason, it is important to underline that in the described systems the local turbulence is not given mainly by the flow-rate, but by the static elements that warrants an intimate mixing between the two immiscible phases because the formation of eddies. 3.2. Methanol–soybean oil transesterification runs performed in the described packed bed tubular reactors. The typical experimental conditions for many of the runs performed, are: molar ratio methanol/soybean oil = 6:1 considering for soybean oil an average molecular weight of 885, a KOH catalyst concentration of 1% b.w. referred to the oil and a temperature of 60 ◦ C. 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 8.33 cm3 /min. The esters yield obtained in this case was only 3%. This blank value can be compared with the much higher values obtained in the presence of a packed bed. 3.2.1. Methanol–soybean oil transesterification runs performed with the RX1 reactor Some experimental runs were performed fixing the temperature (60 ◦ C) and the catalyst concentration (KOH 1% b.w. referred to the oil) in order to study the effect of the residence time on the methylester yield, while another test was performed varying the catalyst concentration at a fixed residence time. A list of the runs performed and related results are reported in Table 2. It must be observed, first of all, that the presence of the packed bed gives place to a very strong increase of the conversion with respect to the empty reactor due to the static mixing effect. Then, it must be pointed out that, by increasing the residence time, the yield in methylesters increases until a maximum is reached, a further increase of the residence time gives place to a decrease of the yield. This behavior can be better appreciated by plotting the yields as a function of the resi-

100 95

RX1 RX3 Kinetic model prediction

90 85 80

Yield (%)

Flow rates (cm3 /min)

1089

75 70 65 60 55 50 45 40 0

1

2

3

4

5

6

7

8

9

10

Residence time (minutes) Fig. 6. Methylesters yields plot as a function of the residence time obtained using the RX1 and RX3 reactors. In the same plot is reported for comparison also the simulation obtained with a kinetic model derived from kinetic runs performed in well-mixed batch reactors [5,6].

dence time, as reported in Fig. 6. The lowest yield of 44%, obtained at about 10 min of residence time, is very probably due to a less active micromixing giving place to a reduction of the interphase area and to the intervention of mass transfer limitation on the reaction rate. On the contrary, the decrease of the yield observed, at the lowest residence time, would be the normal behavior of a plug flow reactor working in a chemical regime. Another observation is that by decreasing the catalyst concentration the activity consistently decreases (see entry 8). 3.2.2. Runs performed with the RX2 reactor Another set of experimental runs was performed in the reactor with the second typology of filling (RX2) by imposing the following conditions: temperature of 60 ◦ C and catalyst concentration of 1% or 2% b.w. of KOH. The list of the run performed is shown in the Table 3. In this case, the observed behaviors are very singular and can be better appreciated in Fig. 7. The yields decrease very slightly by increasing the residence time for the runs performed with 1% of catalyst and the values, in particular at low residence time are greater than the ones obtained with the RX1 reactor. On the contrary, for the runs performed with 2% of catalyst a small increase of the yield with the residence time is observed until reaching a maximum yield of 98.17% in correspondence of one minute of residence time. This probably is the yield equilibrium value. It is very interesting to note that the role of the catalyst concentration is determinant in this case. As a matter of fact, the yield goes from 82% to 98% by using respectively 1% b.w. or 2% b.w. of catalyst. 3.2.3. Runs performed with the RX3 reactor Another set of experimental runs was performed in the already described RX3 reactor, always by fixing the temperature, at 60 ◦ C and KOH catalyst concentration, at 1% b.w. of oil. The list of the runs performed with this reactor is reported in Table 4 together with the

Table 2 Transesterification runs performed using the RX1 configuration. Run

Qoil (cm3 /min)

QMeOH (cm3 /min)

Qtot (cm3 /min)

T (◦ C)

 (min)

KOH (%b.w.)

Yield (%)

1 2 3 4 5 6 7 8

9.80 6.60 5.50 4.40 1.1 2.4 0.58 6.60

2.53 1.73 1.40 1.15 0.287 0.615 0.15 1.73

12.33 8.33 6.90 5.55 1.39 3.02 0.73 8.33

60 60 60 60 60 60 60 60

0.58 0.86 1.04 1.29 5.16 2.37 9.79 0.86

1 1 1 1 1 1 1 0.5

62.5 79.3 80.7 81.8 73.3 73.3 43.6 55.4

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Table 3 Transesterification runs performed using the RX2 configuration. Run

Qoil (cm3 /min)

QMeOH (cm3 /min)

Qtot (cm3 /min)

T (◦ C)

 (min)

KOH (%wt.)

Yield (%)

9 10 11 12 13 14 15 16 17 18

28.62 19.24 9.62 4.85 2.38 28.62 19.24 9.62 4.85 2.38

7.382 4.962 2.480 1.251 0.615 7.382 4.962 2.480 1.251 0.615

36.00 24.20 12.10 6.10 3.00 36.00 24.20 12.10 6.10 3.00

60 60 60 60 60 60 60 60 60 60

0.17 0.25 0.50 0.99 2.02 0.17 0.25 0.50 0.99 2.02

1 1 1 1 1 2 2 2 2 2

83.6 81.1 80.3 81.9 78.7 93.6 93.8 94.7 98.2 97.1

Table 4 Transesterification runs performed using the RX3 reactor configuration. Run

Qoil (cm3 /min)

QMeOH (cm3 /min)

Qtot (cm3 /min)

T (◦ C)

 (min)

KOH (%wt.)

Yield (%)

20 21 22 23 24 25 26 27 28 29

9.80 7.90 6.60 5.50 4.40 3.20 2.4 1.10 0.58 4.40

2.53 2.04 1.73 1.40 1.15 0.83 0.615 0.29 0.15 1.15

12.33 9.94 8.33 6.90 5.55 4.03 3.02 1.39 0.73 5.55

60 60 60 60 60 60 60 60 60 100

0.31 0.38 0.46 0.55 0.68 0.94 1.26 2.73 5.21 0.68

1 1 1 1 1 1 1 1 1 1

80.5 81.5 78.6 77.2 81.5 80.6 82.3 91.6 80.5 85.8

obtained results in terms of yields to methyl esters. From the data reported in Table 4 we can observe, first of all, that the obtained yields are normally higher than the ones obtained with RX1 reactor and comparable with those of RX2 reactor, at the lowest residence time, but slightly higher at 2.02 min of residence time. This can tentatively be attributed to the development of a greater surface area at the liquid–liquid interphase due to both the effect of the intense local micromixing and of the presence of narrower microchannels (300 ␮m). Again, a maximum in methylesters yields is reached by increasing the residence time to a value of 2.73 min, as it can be appreciated in Fig. 6. Also in this case, as observed for the RX2 reactor, the trend of the yield with the residence time is almost flat in the examined range. These behaviors, as before mentioned, are very singular and difficult to be described with a simple kinetic model based on a pseudo-monophasic assumption like the ones proposed by the literature. In practice, in the range of residence time of 0.5–6 min the yields does not change very much.

100 90

This means that operating at high volumetric flow-rates the productivity of the system becomes very high. At last, the effect of the temperature (see entry 24 and 29 of Table 4 for the comparison) in the described conditions, is consistent but not so important as the effects of both micromixing and catalyst concentration (see runs made with RX2 reactor). Some kinetic runs on RX1, RX2 and RX3 have been repeated different times, in the same conditions, in order to evaluate the reproducibility and an average experimental error of less than 1–2% has been found. 3.3. Kinetic interpretation of the runs The experimental data collected has been interpreted by adopting a kinetic model already published in the literature. All the kinetic approaches, proposed by different authors, have been elaborated on the basis of runs performed in batch laboratory reactors by considering, as a first approximation, the reaction system pseudo-homogeneous and in kinetic regime [2,5,6]. Therefore, the concentration of each reagent and catalyst has been calculated considering the total volume of the two phases. The reaction occurs in three different consecutive steps, according to the scheme (6): k1

T + AD + E

80

Yield (%)

k2

D + AM + E

70

k3

M + AG + E

60 50

1 % KOH Kinetic model prediction 1% KOH 2 % KOH Kinetic model prediction 2% KOH

40 30 20 0

20

40

60

80

100

120

140

160

180

200

Residence time (s) Fig. 7. Methylesters yields plot as a function of the residence time obtained using the RX2 reactor by using different amounts of catalysts (1 and 2% b.w). In the same plot are also reported for comparison the simulations obtained with a kinetic model derived from kinetic runs performed in well-mixed batch reactors [5,6].

Keq Keq

(6)

Keq

where T represent the triglycerides, D the diglycerides, M the monoglycerides, G glycerol, A methanol and E the methylesters. In order to describe the evolution with time of reagents and products, in a batch reactor, three kinetic expressions, containing also the equilibrium terms, are necessary, that is, we need to know three kinetic parameters and related activation energies, three equilibrium constants and related enthalpy change parameters. The kinetic laws suggested in the literature, for the three mentioned reactions, are pseudo-second order laws with equilibrium. This approach works reasonably well to describe the runs performed in batch conditions, for which many experimental points as a function of time are available, but the extrapolation of the same kinetics and parameters to the description of few experimental data, obtained in a contin-

E. Santacesaria et al. / Chemical Engineering and Processing 50 (2011) 1085–1094

uous reactor, is problematic for the large number of parameters that must be introduced. For this reason, in this paper, a simplified mathematical model has been developed with the aim to reduce significantly the number of kinetic parameters but describing the experimental data, reported in literature, with a satisfactory accuracy. This model has then been applied to the experimental data collected using the reactors RX1, RX2 and RX3 described in the previous parts of this paper. The mathematical model proposed takes into account the simplified reaction scheme (6), where ki (i = 1, 2 and 3) and Keq are respectively the kinetic and equilibrium constants for the three reactions in series. The mathematical model is based on the following three assumptions:

1091

Table 5 Kinetic and equilibrium parameters obtained by mathematical regression on the experimental data, performed in batch conditions, reported in the literature [5,6]. Run

T (◦ C)

KOH (%wt.)

Oil

k1 (L2 /mol2 min)

Keq

1 2 3

25 35 65

1 0.5 1

Brassica Sunflower Sunflower

1.02 2.69 9.5

3.86 3.31 3.50

Methanol Methyl Esters Monoglycerides Diglycerides Triglycerides Glycerol

5.5 5.0 4.5 4.0

Ci (mol/L)

3.5

- The system is considered as a pseudo-homogeneous (single phase). - The equilibrium constant is assumed the same for all the three reactive steps.

3.0 2.5 2.0 1.5

A linear correlation has been assumed between the three different kinetic constants, that is we assumed that: k2 = 2k1 and k3 = 3k1 . This assumption is reasonable, considering the reaction mechanism, that is, when the catalyst CH3 O− attacks an ester group of the triglyceride, its relative concentration, referred to the number of available esters, is half than in the presence of a diglyceride molecule and one third in the presence of a monoglyceride. It is now possible to write the reaction rate expressions for each step, expressed in [mol L−1 min−1 ], as it follows:



[Cat][T ][A] −

 r2 = 2k1

1 [Cat][D][E] Keq

0.0 0 10 20 30 40 50 60 70 80 90 100 110 120 130

Time (minutes) Fig. 8. Concentration profiles for each component obtained by simulation of the run 1 of Table 5 by using a simplified kinetic model with reduced number of parameters. Symbols represent the experimental data extracted by the literature, lines represent the calculated profiles.

(7)



[Cat][D][A] −

1 [Cat][M][E] Keq

(8)

[Cat][M][A] −

1 [Cat][G][E] Keq

(9)

 r3 = 3k1

0.5



The mass balance equations for each component, expressed in [mol L−1 min−1 ], for a batch reaction system, can be written as follows: d[T ] = −r1 dt

(10)

d[D] = r1 − r2 dt

(11)

d[M] = r2 − r3 dt

(12)

d[G] = r3 dt

(13)

d[E] = r1 + r2 + r3 dt

(14)

d[A] = −r1 − r2 − r3 dt

(15)

batch runs considered. In conclusion the developed model is suitable to describe the experimental runs performed in discontinuous conditions in the temperature range 25–65 ◦ C. Another important aspect concerns the dependence on the temperature of both the equilibrium and the kinetic parameters obtained by mathematical regression analysis on the experimental runs reported by the literature. In Fig. 9, the fitting of the logarithm of both the kinetic and the equilibrium constants against the reverse of the temperature is reported. From this plot, it is possible to calculate the activation energy and the pre-exponential factor for the kinetic constant k1 . The obtained values are reported in Table 6 together with the standard errors. The equilibrium constant resulted independent of the temperature and equal to an average value of 3.52. The kinetic parameters seem independent of the type of used oil, as we have 2.5

Kinetic constants runs 1 - 2 - 3 Equilibrium constants runs 1 - 2 - 3

2.0

In order to test the goodness of the model described above, some experimental data reported by the literature have been used for determining by regression the kinetic parameters [5,6]. The kinetic runs considered, according to the authors, would not be limited by mass-transfer phenomena. The best kinetic parameters, obtained by regression on the data reported in the literature [5,6], are reported in Table 5. An example of the obtained agreement between the experimental and calculated data for the run 1 of Table 5 is reported in Fig. 8. As it can be observed, there is a good agreement between experimental data and calculated profile. The same occurs for the other

ln (k1), ln (Keq)



r1 = k1

1.0

1.5

1.0

0.5

0.0 0.0029

0.0030

0.0031

0.0032

0.0033

0.0034

-1

1/T (K ) Fig. 9. Fitting of the logarithm of both the kinetic and the equilibrium constants against the reverse of the temperature.

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E. Santacesaria et al. / Chemical Engineering and Processing 50 (2011) 1085–1094

Table 6 Activation energy, pre-exponential factor and equilibrium constant obtained from Fig. 9.

RX1 (KOH 1.0 % b.w.) RX1 (KOH 0.5 % b.w.) RX2 (KOH 1.0 % b.w.) RX2 (KOH 2.0 % b.w.) RX3 (KOH 1.0 % b.w.)

5.5 5.0

Value

2

10.67 18.2 3.52

1.94 3.1 0.045

observed for, respectively, Brassica and Sunflower oils. We can assume, therefore, that the kinetic model and related parameters can be applied also for soybean oil in the elaboration of the experimental data obtained in the described continuous systems. For these systems it is possible to write the mass balance equations for each component fixing, as the integration variable, the reactor volume.

4.5 4.0

kapp1/k1

Ea (Kcal/mol) ln A (L2 /mol2 min) Keq

6.0

3.5 3.0 2.5 2.0 1.5 1.0 0.5 0.0 0

−r1 d[T ] = Q dV

(16)

r1 − r2 d[D] = Q dV

(17)

r2 − r3 d[M] = Q dV

(18)

r3 d[G] = Q dV

(19)

r1 + r2 + r3 d[E] = Q dV

(20)

−r1 − r2 − r3 d[A] = Q dV

(21)

Q is the overall volumetric flow rate (expressed in [cm3 /min]) given by the sum of the volumetric flow rates of both oil and methanol sent to the reactor. Setting the values of the reactor inlet concentrations, the kinetic runs performed in the reactors RX1, RX2 and RX3 have been simulated with the developed kinetic model and related parameters by using the software Berkeley–Madonna v.8.0 setting a fourth-order Runge–Kutta algorithm in order to integrate the differential equations ((16)–(21)). The simulation results are reported in Figs. 6 and 7 and compared with the corresponding experimental kinetic data. By considering the runs of Fig. 6, related to the reactors RX1 and RX3, as all the runs have been performed at the same temperature of 60 ◦ C, by using the same methanol/oil ratio, the first observation is that the yields for RX1 and RX3 would be the same for a given residence time, but this does not occurs, as it can be seen in the figure. By comparing, then, the curve related to the simulation made with the performances of each considered reactor configuration we can note that the runs performed in RX1, at lower residence times, are well described by the model, while, the experimental data obtained at higher residence times are lower, probably, for the occurrence of the mass transfer limitation. In the case of RX3 the yields obtained, at lower residence times, are much higher than the simulated ones. This behavior can be explained by assuming that for very active local micromixing or in the presence of narrow micro-channels, as it occurs in the reactor RX3, the one-phase homogeneous kinetic model is not adequate to describe the kinetic behavior of the transesterification reaction. The simulation curves obtained for the runs performed in the RX2 reactor are reported in Fig. 7. As it can be seen, in both cases, that is for the runs performed with respectively 1% or 2% of the catalyst, the kinetic model does not reproduce again the runs at lower residence times. A satisfactory agreement is obtained only for the runs performed with 2% of catalyst at higher residence time, because, in these cases we are near to the equilibrium conditions. By interpreting all the kinetic runs performed in the three mentioned reactors with the described model the agreement could be obtained only

10

100

Residence time (seconds) Fig. 10. Apparent kinetic constants for all the performed runs.

by considering a value for k1 = kapp 1 greater than 6, the value corresponding to the kinetic constant found at 60 ◦ C for a well stirred batch reactor, or alternatively lower than this value for a reaction limited by mass transfer. In Fig. 10 the ratio kapp 1 /k1 is reported as a function of residence time for all the performed runs. A narrow plateau can be observed in correspondence of kapp 1 /k1 = 1. While, a value inferior to 1 can easily be interpreted as a consequence of mass transfer limitation, the greater values probably are the consequence of the complex equilibria occurring in the studied system. By using ChemCAD 6.2 software with UNIFAC LLE model [22] it is possible to estimate the composition of the biphasic system as a function of the conversion (see Fig. 11). As it can be seen, we have a polar phase that initially contains mainly methanol–diglyceride and monoglyceride (0–60%), while, at conversion greater than 60% methanol + glycerol predominates in the polar phase. The catalyst CH3 O− , formed as a consequence of the reaction, KOH + CH3 OH → KOCH3 + H2 O

(22)

dissolves mainly in the methanolic polar phase and promotes initially the reaction. Then CH3 O− is involved in ion exchange equilibrium reactions with the hydroxyls of di- and mono-glycerides and after 60% of conversion directly with glycerol, giving place to glyceroxide anions whose activities are not known. Moreover, mono, di- and triglycerides are very few soluble in the methanol–glycerol polar phase, where the catalyst (mixture of methoxide and glyceroxide anions) is dissolved at the end of the reaction. For this reason, after 60–70% of conversion, the reaction rate rapidly slows down. The increase of the catalyst concentration from 1 to 2% of KOH affects the equilibrium of the anions population and the conversion dramatically increased as a consequence. Another important aspect is that the volume ratio between polar and apolar phase decreases by increasing the conversion (see Fig. 12), from about 0.17 to a minimum of 0.13, at 0.7 of conversion, then increases coming back approximately to the original value but changing completely the nature of both the polar and apolar phase (see compositions in Fig. 11). As a consequence, the catalyst concentration initially increases reaching a maximum and then decreases. All these changes occur inside the described continuous tubular reactors from the inlet to the outlet. Maybe that all these aspects are not sufficient to explain all the experimentally observed phenomena such as: the strong increase of the conversion for the presence of an inert filling material of opportune shape and size (activity more than 5 times the one observed in well stirred batch reactors), the strong increase of conversion simply by doubling the amount of used catalyst and the observed invariance of the con-

Component mass fraction (X i)

E. Santacesaria et al. / Chemical Engineering and Processing 50 (2011) 1085–1094

0.0 1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0 1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0 0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1093

1.0 0.20 Left y axis Methanol 0.15 Glycerol

Polar Phase

Right y axis Triglycerides Methylesters 0.05 Monoglycerides Diglycerides 0.00 0.08 0.07 Left y axis Triglycerides 0.06 Methylesters 0.05 Diglycerides 0.04 0.10

Apolar Phase

0.03 Right y axis 0.02 Glycerol 0.01 Methanol 0.00 Monoglycerides 0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

Conversion Degree Fig. 11. Mass fraction plot against the conversion degree related to the triglycerides for each component involved in the transesterification reaction. The upper section of the diagram is related to the polar phase, while, the lower to the apolar one. The right axis of both the sections is related to: methanol and glycerol for the polar phase, triglycerides and methylesters for the apolar phase. The simulation has been performed by using ChemCAD 6.2 software, setting the UNIFAC LLE model.

obtained performances with a pseudo-homogeneous kinetic model derived from the literature completely failed. In the presence of an intense local micromixing or very narrow microchannels the transesterification reaction rates are much higher than the ones predicted by the mentioned kinetic model. It is clear that a more sophisticated model is required to describe the behavior of continuous reactors characterised by the presence of local micromixing or microchannels and this will be our next research challenge.

Polar/Apolar (vol./vol.)

0.18

0.16

0.14

0.12

Acknowledgements

0.10

0.08 0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

Conversion Degree Fig. 12. Polar on apolar phase ratio as a function of the conversion degree related to the triglycerides setting a reagent molar ratio feeding of MeOH/oil 6:1. The calculation has been performed by using ChemCAD 6.2 software, setting the UNIFAC LLE model.

version with the residence time. However, a more reliable biphasic kinetic model would be useful to have a more detailed description of what effectively occurs in the studied system. The development of such a biphasic model will be the subject of our next work in the field.

Thanks are due to EC VII Framework Programme CP-IP 2288532 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. Appendix A. List of symbols

RX1 RX2 RX3

4. Conclusions We have shown that a tubular reactor filled with small spheres of different sizes can usefully be used to simulate the behavior of both a “static mixer” and a “microreactor”. As a matter of fact, with this type of device it is possible to change opportunely the size of the microchannels, so simulating microreactors, or the intensity of local micromixing by changing feeding flow rates, so simulating static mixers. A tubular reactor filled with spheres of different sizes has been tested in the transesterification of vegetable oils with methanol, promoted by KOH, for producing biodiesel. Very high conversions in a very short residence time (less than 1 min) at 60 ◦ C with a methanol/oil ratio equal 6, have been obtained confirming the performances already observed in the literature by using static mixers or microreactors. This reaction is characterised by the presence of two immiscible liquid reactants and the attempt to describe the

ε F

 ND Pe D ui L d Rep  

tubular packed bed reactor filled with 2.5 mm diameter AISI 316 spheres tubular packed bed reactor filled with 2.5 and 1.0 mm diameter AISI 316 spheres tubular packed bed reactor filled with 2.5 and 0.39 mm diameter AISI 316 spheres void degree dimensionless tracer concentration (C/C0 ) where C and C0 are respectively the punctual and the initial concentrations dimensionless time (t/tm ) where t and tm are respectively the punctual and the average residence time dispersion number Peclet number, defined as the reverse of the dispersion number axial dispersion coefficient (m2 /s) axial velocity (m/s) reactor’s length (m) diameter (m) Reynolds particles number viscosity (Pa s) density (kg/m3 )

1094

 Q A1 A2 r R Keq ki ki app T D M E G A Cat ri

E. Santacesaria et al. / Chemical Engineering and Processing 50 (2011) 1085–1094

residence time (min) overall volumetric flow rate (cm3 /min) methoxylic groups single signal at ı = 3.7 ppm methylenic groups triplet signal ı = 2.3 ppm radius of the spheres (mm) radius of the spheres (mm) equilibrium constant kinetic constant apparent kinetic constant triglycerides diglycerides monoglycerides methylesters glycerol methanol catalyst reaction rate (mol L−1 min−1 )

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