Process Intensification for Hydroprocessing of Vegetable Oils

Article pubs.acs.org/IECR

Process Intensification for Hydroprocessing of Vegetable Oils: Experimental Study A. K. Sinha,* M. G. Sibi, N. Naidu, S. A. Farooqui, M. Anand, and R. Kumar Hydroprocessing Area, CSIR−Indian Institute of Petroleum, Dehradun 248005, India ABSTRACT: Process intensification for hydroprocessing of vegetable oils was achieved using microchannel and monolithic reactors with lower diffusion lengths, a thin layer of catalyst coating, high surface-to-volume ratios, and better heat/mass transfer. Use of such reactors results in improved reaction selectivities and throughput, with reliable utilization of catalysts. Laboratory scale microchannel and monolithic reactors were washcoated with sol−gel prepared Ni−Mo/γ-Al2O3 catalyst. The catalyst was thoroughly characterized by various techniques. Activity tests were performed on the sulfided catalyst for the hydroprocessing of vegetable oil, a mass and heat transfer limited reaction. Both reactors gave up to 20 times higher kerosene selectivities, up to 30 times higher isomer selectivities, and at least 20 times higher throughput than the conventional trickle-bed reactor. Undesirable oligomeric product yield was 3 times lower in the case of the monolithic reactor than in the case of the microchannel reactor.

1. INTRODUCTION Microstructured (microchannel and monolithic) reactors have smaller process volumes and holdup, larger specific transport rates, and higher surface-to-volume ratios.1 These reactors give several-times-higher heat flux than conventional reactors and enable reduction in size, narrow residence time distributions, reliable utilization of catalysts with reduced quantity, and lower cost of the chemical processing hardware.2,3 Microstructred reactors offer the advantage of thinner walls with thin fluid layers, high geometric surface areas, and low pressure drops. These key features significantly enhance heat and mass transfer rates, improve process control, and accelerate the contacting and mixing of the fluids.4 At the same time the process parameters such as pressure, temperature, residence time, and flow rates are more easily controlled.5 Despite several advantages, there are still challenges and uncertainties in the application of microchannel reactors.6 The major technical challenges are component integration into more complex systems and poor mixing especially with liquid phase systems.7 Another issue is the joining of the microchannel plates, which is inconvenient and costly for large scale manufacturing.8 As the benefits of microchannel reactors are evident, they have been applied in several applications which are mostly mass transfer and heat transfer limited reactions. Carrying out reactions which are highly exothermic or endothermic in these reactors provide better heat transfer coefficients. Monoliths offer lower pressure drop and higher specific external catalyst surface area for better mass transfer, compared to microchannel reactors. When compared to microchannel reactors, they may have higher selectivity due to lower axial dispersion and back mixing, extended life due to reduction of fouling, easy cleanup, and easy scale-up. By contrast, it is difficult to maintain temperature control inside a monolithic reactor with thin wall ceramic monolith supports, due to low radial heat transfer rate, lower heat transfer rate from the monolith to the internal reactor wall, potential nonuniform fluid distribution, higher cost, and lack of experience in large scale processes.3,6 © 2014 American Chemical Society

Since higher catalytic activities and selectivities are expected in microstructured reactors compared to traditional reactors, effective catalyst coating with a long life is necessary. A number of thin layer catalyst deposition methods are available in the literature. They are sol−gel, vapor deposition, electrophoretic deposition, impregnation, flame spray deposition, powder plasma spraying, electrochemical deposition, and electroless plating.9−17 Among these, the sol−gel method has advantage of preparing mixed oxides at lower temperatures, with controlled shape and tailored porosity. Production of liquid fuels from gases18,19 and biomass,20 hydrogenation,21−27 reforming,28−30 and CO selective oxidative processes for proton exchange membrane (PEM) fuel cells31,32 are some of the energy sector applications of these reactors. Hydroprocessing of vegetable oils,33 and coprocessing of vegetable oils and petroleum-derived oils for the production of gasoline, kerosene, and diesel, has been well studied.34−36 Vegetable oil triglycerides are bulky molecules which create considerable mass transfer limitations during reactions in a packed bed reactor. Hydrocracking of vegetable oil is highly exothermic reaction, nearly 10 times more exothermic than conventional gas-oil hydrocracking. The present work aims at looking into process intensification benefits of laboratory scale microchannel and monolithic reactors for hydroprocessing of jatropha oil, using sulfided Ni−Mo/Al2O3 catalysts. The reaction conditions for optimizing the throughputs and product selectivities were studied in detail and correlated with the results obtained.

2. EXPERIMENTAL SECTION 2.1. Materials and Methods. Aluminum isopropoxide (AIP), acetyl acetone (AcAc), aluminum acetyl acetone (AlAcAc), polyvinyl alcohol (PVA), nitric acid (concentrated), Received: Revised: Accepted: Published: 19062

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Figure 1. Microchannel plate design with dimensions.

abraded with alumina powder for removal of unwanted loose particles, and again degreased. Then, acid treatment (H2SO4 or HCl) followed by thermal oxidation (800 °C) was performed. Finally, primer coating (alumina washcoat) was applied. A commercial ceramic monolith (Mg2Al4Si5O18; 13% MgO, 34.86% Al2O3, 51.36% SiO2) was dried in an oven to remove moisture followed by primer treatment (alumina washcoating) to improve the adherence of the catalyst layer on the channels. 2.5. Deposition of Washcoat. Washcoats were deposited on the bare metallic microchannel substrate. The microchannel plate was weighed before each step of coating and calcinations, to get the accurate amount of catalyst deposited inside the channels. The channels were filled with washcoat solution and drained vertically, removing excess solution from the channels. In the case of monolithic reactors, washcoats were deposited in the ceramic monolith channels using the dip-coating technique. The substrates were vertically immersed in the washcoat solution for 1 min. The immersed substrates were gradually withdrawn from the washcoat solution (at a velocity of 10 cm/min) and vertically drained at room temperature. The excess solution was blown out of the monolith channels. All materials were dried at room temperature for 5 h followed by drying in an oven at 100 °C overnight. Finally, all the materials were calcined at 600 °C. The adhesion test of the catalyst was carried out following two test methods. In the first test, the substrate was exposed to ultrasonication for 30 min, and in the second test (drop test), it was dropped from a height of 50 cm onto a wooden table (catalyst layer facing the table). The percentage of weight loss was 3.4%, which proved that the washcoat layer was stable (microchannel reactor). A part of the suspension was dried and calcined separately under the same conditions for characterization. After catalyst coating, adhesion tests were carried out in similar way for all the reactors, to ensure proper catalyst adherence to the metallic/ceramic surface. For good adherence the weight loss should be less than 3.5%; otherwise, catalyst was recoated on the channel.

deionized water, ethyl alcohol, dimethyl disulfide, nickel nitrate hexahydrate, and ammonium heptamolybdate tetrahydrate were purchased from Aldrich and used as received. The microchannel reactor was made using SS-316 material (Fe, C18 oligomers was only marginally suppressed. In the case of a trickle-bed reactor we had reported that the yield of >C18 oligomers did not change significantly at higher space velocities.37 However, for the MR (Table 2, MR-6, MR-7), at very high space velocities (132, 159 h−1) significantly higher amounts (nearly 20−30% of product) of undesirable >C18 oligomeric products were observed. Similar patterns (more >C18 oligomeric products at higher space velocities) were observed for the MCR, too (Table 2, MCR-3− 5). This indicates that, though it was possible to have high conversion of jatropha oil at very high space velocities, the increased formation of undesirable >C18 oligomers limits the operation at such high throughputs. The formation of these oligomers (>C18) was less at lower pressure (Table 2, MCR-3) as compared to higher pressures (Table 2, MCR-4). There is a nonmonotonic dependence of hydrocarbon composition with WHSV as expected and as observed in our previous studies.37 Hence there is a optimum space velocity for both microchannel and monolith reactors. At the optimum space velocity of 53 h−1 the product distribution is similar in the reaction condition range of 400−420 °C, 1700−3400 H2/HC ratio for MR (Table 2, MR-1, MR-2, MR-3). Distributions of desired (C9−C18) hydrocarbon products were different over the two types of reactors (Table 2). The maximum yield of pure C9−C18 hydrocarbons over the MR was about 83%, of which about 64% was the direct deoxygenation heavy product (C16− C18) and 20% was the cracked middle range (C9−C15) hydrocarbons (see MR-4 in Table 2). The maximum desired pure (C9−C18) hydrocarbons over the MCR was about 50%, of which about 33% was the direct deoxygenation heavy product (C16−C18) and 21% was the cracked middle range hydrocarbons (Table 2, MCR-6). It was also observed that the MR could be operated at very high space velocities (up to 159 h−1) with negligible pressure drop across the reactor, whereas in the case of the microchannel reactor a high pressure drop was observed at very high space velocities (>46 h−1). This difference could be due to the reactor configuration. The MR has 126 square channels with straight and vertically opposite inlet and outlet,

Figure 5. Conversion of triglycerides as a function of temperature in microchannel reactor (MCR, 46 h−1) and monolith reactor (MR, 53 h−1) at 80 bar. Inset shows the conversions per unit area of the reactor.

Table 2. Product Distribution for Both Microchannel (MCR) and Monolith (MR) Reactors (420 °C, 80 bar, >99% Conversion) exptl code

H2/HC

WHSV

C18

MCR-3a MCR-4b MCR-5c MCR-6 MR-1c MR-2 MR-3 MR-4 MR-5 MR-6 MR-7d

2500 2500 2500 5000 1700 1700 3400 3400 3400 3400 3400

46 46 46 23 53 53 53 26 106 132 159

23.77 10.95 10.93 14.08 11.89 13.04 11.57 7.4 7.05 8.37 11.9

20.71 10.36 21.35 21.41 24.34 17.69 20.74 19.25 22.06 28.13 25.49

14.17 11.60 16.52 33.25 56.7 61.78 57.61 63.64 60.71 35.09 43.18

41.34 64.22 51.20 31.27 7.5 7.52 10.07 9.69 9.6 28.39 19.4

a

At 380 °C, 38 bar. bAt 380 °C. cAt 400 °C. dAt 93% conversion.

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monotonic dependence of product composition with WHSV is observed, particularly at higher space velocities. Since the MCR forms more oligomers, we checked the effect of temperature on the product profile to find the reaction temperature at which oligomer formation was minimized. Figure 7 shows that with increasing reaction temperature the

while the MCR has 16 channels with nearly zigzag inlet and outlets, which leads to a greater pressure drop. For the monolithic reactor (MR), since no pressure drop was observed at various space velocities, we did a detailed study of this reactor at various space velocities. Figure 6 shows that with

Figure 6. Product compositions in reactor effluent at various space velocities (at 420 °C) for monolithic reactor (MR).

Figure 7. Product distributions at various temperatures (at 46 h−1) for microchannel reactor (MCR).

yield of diesel range products (C15−C18) increased from 8 to 28%, those of undesirable (>C18) oligomers decreased from 37 to 21%, those of light naphtha products (C18) product formation. 3.2.2. Isomerization Selectivity. The presence of isomers is important in the reaction product to achieve the desired properties of the fuel. The isomer/normal (i/n) hydrocarbon ratios obtained over both microchannel and monolithic reactors with various reaction parameters are presented in Table 3. The i/n ratios for the C9−C15 fraction are about 1−2 for the

increasing space velocity the yield of diesel range products (C15−C18) decreased from 70 to 45%, those of undesirable (>C18) oligomers increased from 10 to 20%, those of light naphtha products (106 h−1) the product pattern changes significantly. Space velocities of 106 h−1 and lower were best for avoiding the formation of undesirable oligomers, which are coke precursors in the case of the MR. There is a nonmonotonic dependence of product yields with WHSV. The product yields for jatropha oil conversion go through a maximum as earlier reported.37 In the case of the monolithic reactor at lower space velocities (26−106 h−1), i.e., at high residence time, there was nearly constant formation of all the products. Primary direct deoxygenation diesel range products (65−70%), undesired oligomeric product (>C18, ∼10%), as well as secondary cracked products, 420 °C) to minimize oligomer formation. Higher selectivities and throughputs were obtained over these reactors than over a conventional trickle-bed reactor: jet fuel (kerosene) selectivities were 6−20 times higher, isomer selectivities were 20−30 times higher, and feed throughput was at least 20 times higher.

Figure 9. Decarbonation selectivities for monolithic reactor (MR).

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Table 4. Comparison of Product Yields for Microstructured and Trickle-Bed Reactors (at >99% Conversion)35 reactor

catalyst

WHSV (h−1)

C6−C14

C15−C18

i/nd

trickle beda MCRb MRc

Ni−Mo/Al2O3 Ni−Mo/Al2O3 Ni−Mo/Al2O3

1 46 53

4.2 44.5 36.2

95.8 14.2 61.7

0.1 2.3 0.8

Corresponding Author

*E-mail: [email protected]. Tel.: +91-1352525842. Notes

The authors declare no competing financial interest.

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At 380 °C, 50 bar, H2/feed = 1700. Microchannel reactor, 380 °C, 38 bar, H2/feed = 2500. cMonolithic reactor, 400 °C, 80 bar, H2/feed = 1700. dIsomerized to normal hydrocarbon ratio.

a

AUTHOR INFORMATION

b

ACKNOWLEDGMENTS The financial support for this research from DST, India, and CSIR (Project No. CSC-0115) is gratefully acknowledged. M.G.S. acknowledges UGC, India, for a fellowship. We thank the reviewers for their comments which helped us to improve the manuscript significantly.

cracking of the deoxygenation products into lower hydrocarbons. In comparison, there was considerable secondary cracking over the microstructured reactors. At various reaction conditions the cracked products yield is 20−45% (Table 4). For example, 45% C6−C14 cracked products are obtained using the microchannel reactor (MCR) (at 380 °C), while 36% C6−C14 cracked products are obtained using the monolithic reactor (at 400 °C) (Table 4). The isomerization selectivity is also much higher (8−23 times higher) over the microstructured reactors. From these results it can be concluded that with the same catalytic system microstructured reactors give higher cracked and isomerized product yields, compared to fixed trickle-bed reactors. The space velocities for the microstructured reactors are several times higher (20−150 times higher) than those for the trickle-bed reactor. These differences result in different reaction pathways over the two types of reactors. Greater amounts of acidic intermediates were formed using microstructured reactors which catalyze the cracking and isomerization reactions. TAN (total acidity number) analyses of the reaction products showed that the products from microstructured reactors had 2−3 times higher acidity than the products from the trickle-bed reactor. Our earlier works also have indicated the role of acidic intermediates in catalyzing cracking and isomerization reactions during hydroprocessing of vegetable oils.35,37

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4. CONCLUSIONS In conclusion, process intensification using microstructured (microchannel and monolithic) reactors for hydroprocessing of vegetable oils was achieved. Improved reaction selectivity and 19069

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