Î³-Al2O3 for Hydrodeoxygenation of

Article pubs.acs.org/EF

Evaluation of Presulfided NiMo/γ-Al2O3 for Hydrodeoxygenation of Microalgae Oil To Produce Green Diesel Lin Zhou* and Adeniyi Lawal New Jersey Center for Microchemical Systems, Department of Chemical Engineering and Materials Science, Stevens Institute of Technology, 1 Castle Point on Hudson, Hoboken, New Jersey 07030, United States ABSTRACT: In the present work, reduced presulfided NiMo/γ-Al2O3, the conventional hydrotreating catalyst, was evaluated for green diesel production via hydrodeoxygenation of unrefined microalgae oil in a microreactor, mimicking the single channel of a monolithic reactor. The effect of reactor inner diameter on space-time yield of hydrocarbon and microalgae oil conversion was studied first to confirm the superiority of the microreactor for three-phase reactions. Based on the external and internal mass transfer limitation analyses, a range of process conditions without mass transfer limitation was determined for catalyst evaluation. The results showed that NiMo/γ-Al2O3 is deactivated due to the accumulation of produced oxygenated intermediates in hydrodeoxygenation reaction, and its selectivity to even-numbered carbon hydrocarbon produced from hydrodehydration correlates with the catalyst activity. The catalyst activity and life can be preserved by increasing hydrogen to oil ratio, residence time, reaction temperature, and pressure, which will decrease the adsorption of oxygenates on the catalyst surface. For the reaction condition: 500 psig H2, 360 °C, H2/oil ratio of 1000 SmL/mL, and residence time of 1 s, the initial catalyst activity was maintained without any signs of deactivation for at least 7 h and the obtained C13 to C20 hydrocarbon yield was 56.2%, with a carbon yield of 62.7%, nearly complete conversion (98.7%) of microalgae oil, and HC(2n)/HC(2n − 1) ratio of 6.

1. INTRODUCTION Microalgae are ubiquitous, and although they are primarily found in all the oceans and seas, an area that covers 71% of the Earth’s surface, they also grow in freshwater bodies as well as on and in soil, rocks, ice, snow, plants, and animals.1 They grow extremely rapidly and can double their biomass within 24 h, which is 10−200 times faster than terrestrial oil crops.2,3 Oleaginous phototropic microalgae are sunlight-driven cell factories that convert carbon dioxide to lipids, with an average lipid content varying between 1% and 70%, and even reaching up to 90% of dry weight under certain conditions.2,4,5 The accumulated oil in almost all microalgae is mainly triglycerides (>80%) with a fatty acid profile rich in C16 and C18,2 which is a substantial energy resource for liquid fuel production. Therefore, microalgae oil has been considered as a good candidate for low-net carbon liquid transportation fuel production, which has no direct competition with edible food or oil production. However, some of the physical properties of microalgae oil prohibit its direct use in existing engines, such as low flowability, high viscosity, and low volatility. In order to improve the oil’s physical properties while maintaining its heating value, an ideal upgrading process should only rearrange the oil molecular structure while avoiding or minimizing cracking.6 On the basis of the experience of vegetable oil upgrading, the two currently practiced oil upgrading methods are transesterification and hydrotreating, which produce biodiesel and green diesel, respectively. Biodiesel has a lower viscosity than its parent oil, but still much higher than that of petrodiesel. With the cold-flow related issues unresolved,7 biodiesel cannot be used in its pure form; rather, it is blended with petrodiesel, the most common product, commercially referred to as B20, blends 20 vol % biodiesel with 80 vol % petrodiesel. Green diesel, on the other hand, is essentially a © 2014 American Chemical Society

mixture of hydrocarbons, has the same chemical properties as petrodiesel, and is compatible with all existing engines, pipelines, and infrastructure for application and distribution. Compared with biodiesel, green diesel is a more ideal substitute for petrodiesel. In the production of green diesel from microalgae oil, oxygen needs to be removed from lipids, mainly triglycerides, to produce hydrocarbons. Because the sulfur, nitrogen, and phosphorus contents in microalgae oil are really low, the main heteroatom to be removed is oxygen. For this reason, the hydrotreating of microalgae oil mainly refers to the oxygen removal. Due to the nature of oxygen bonding in triglycerides, removal of oxygen could be achieved in two ways: (1) hydrodehydration (DHYD), in which oxygen is removed in the form of H2O and (2) hydrodecarboxylation (DCO2) or hydrodecarbonylation (DCO), in which oxygen is removed in the form of CO2 and CO, respectively. One triglyceride molecule produces three hydrocarbon chains. As the oxygen is present in crude oil at rather low levels, of the order of 0.5%, deoxygenation in petroleum refining is not of much concern, and no catalysts are specifically formulated for oxygenates hydrotreating. Hence, one of the critical technical challenges to make the hydrodeoxygenation of microalgae oil process economically feasible is related to the research and development of effective catalysts. Many studies have been performed with commercial hydrotreating catalysts, such as CoMo, NiMo, and NiW supported on alumina, for hydrotreating of natural lipids from different sources and have revealed that complete triglycerides conversion could be Received: October 8, 2014 Revised: December 5, 2014 Published: December 5, 2014 262

dx.doi.org/10.1021/ef502258q | Energy Fuels 2015, 29, 262−272

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achieved over these catalysts.8−13 These conventional hydrotreating catalysts are more active in sulfided form than in oxide form.14 The sulfidation creates active sites that can play a role in the rupture of the carbon−heteroatom bond.15 Toba et al. studied the hydrotreating of waste cooking oil in both batch reactor (7 MPa, 250−350 °C) and fixed-bed flow reactor (5 MPa, 350 °C) using CoMo, NiMo, and NiW catalysts. In their study, the NiMo and NiW catalysts showed high and stable hydrogenation activity, whereas deactivation was observed when using the CoMo catalyst. The NiW catalyst favors more hydrodecarboxylation or hydrodecarbonylation than hydrodehydration.16 Senol also reported that the NiMo catalyst showed a higher activity than CoMo in hydrodehydration and hydrogenation reactions.15 Kubička and Kaluža studied the Ni, Mo, and NiMo sulfided catalysts in deoxygenation of rapeseed oil at 260−280 °C, 3.5 MPa, and 0.25−4 h−1 in a fixed-bed reactor. They found bimetallic NiMo catalysts showed higher activity and yields of hydrocarbons than monometallic catalyst; NiMo yielded a mixture of decarboxylation and hydrodehydration products, whereas Ni yielded only decarboxylation products, and Mo yielded almost exclusively hydrodehydration hydrocarbon products.17 Da Rocha Filho et al. studied the hydrotreating of different vegetable oils in a batch reactor using NiMo catalyst and reported a 66−76 wt % n-alkanes yield after 2 h at 360 °C and 14 MPa.18 Huber et al. conducted the hydrotreating of sunflower oil over NiMo catalyst in a fixed-bed reactor under reaction conditions as follows: temperature 350 °C, pressure 50 bar H2, LHSV 5.2 h−1, and H2 to oil ratio of 1600 SmL/mL. The maximum carbon yield in C15−C18 alkanes they obtained was 71%, which is 75% of the maximum theoretical yield of 95%.19 Peng et al. studied the hydrotreating of microalgae oil in batch mode with 10 wt % Ni/HBeta at 260 °C and 40 bar and obtained 78 wt % yield of liquid alkanes after 8 h of reaction time. They also obtained almost identical results from their trickle-bed reactor system under identical experimental conditions.3 As discussed later, in the present study, presulfided NiMo/Al2O3 catalyst was studied for hydrotreating of microalgae oil as the baseline for further catalyst screening work. Although very high oil conversions and hydrocarbon yields were obtained in the studies mentioned above, which were conducted in batch reactors or trickle-bed reactors, it should be noted that the space-time yield (STY) was extremely low due to the severe heat and mass transfer limitations in the conventional macroreactor system. The low STY implies the size of the reactor and all associated ancillary equipment will be unrealistically large, driving up both the capital and operating cost. Therefore, a cost-effective upgrading process demands a different and innovative approach to reactor design for hydrotreating. Monolith reactors are being studied as a substitute for conventional multiphase reactors, such as trickle-bed reactors, slurry reactors, and slurry bubble column reactors for gas−liquid−solid reactions due to their superior hydrodynamics.20 Monolith substrate is usually made from ceramic, but metallic monolith has also been developed for highly exothermic reactions, which enable the coupling of the reactor with a heat exchanger to effectively remove heat in transverse direction and control reaction temperature. The monolith reactor features a honeycomb structure comprising thousands of finely divided flow passages that extend through the whole reactor. The cross-sectional dimensions of each channel are determined by cell density (cells per square inch, cpsi), typically ranging between 100 to 1200 cpsi. Considering

the fact that the typical values for wall thickness range between 0.006 and 0.05 cm, the monolith reactor has corresponding submillimeter and, thus, microchannel dimensions, and such microchannels mimic independent microreactors. Many previous studies21−23 of microreactors applied to different reactions showed that the mass and heat transfer rates were greatly enhanced and were often several orders of magnitude greater than those achievable in conventional reactors. This is attributed to a combination of short diffusion distance and the existence of a flow pattern of alternating gas and liquid slugs, also known as Taylor flow, at certain flow conditions. In Taylor flow, the gas bubbles are separated from the catalyst surface by a very thin liquid film, the thickness of which is much smaller than the reactor diameter. The bulk liquid is separated by gas bubbles, thus reducing axial mixing between the liquid slugs. Also, the recirculation within the liquid slugs improves the radial-mass transfer and mass transfer from liquid to catalyst as well as from gas to liquid. This combination of good radial-mass transfer and low axial mass transfer in the liquid makes Taylor flow suitable for multiphase applications that involve mass transfer or single-phase applications that suffer from significant back-mixing. Moreover, as a result of the enhanced mass transfer, the hydrogen concentration at the catalyst is significantly higher than that at the conventional reactor, which makes the catalyst to be utilized much more effectively and reduces catalyst deactivation. Therefore, the microreactor, by virtue of its enhanced mass and heat transport abilities, allows much shorter residence times and, hence, reduced equipment volume and better control of operating conditions, which should result in a process of lower capital and operating expense compared to conventional technology. Microreactor also has advantages over conventional reactors in the following aspects. First, microreactors contain no agitators and thus no electrical energy input needed to achieve certain flow patterns for better heat and mass transfer. Also, the improved heat transfer enables improved energy efficiency. Therefore, compared to the batch and semibatch reactors, the energy consumption is less in microreactor. Second, process safety is improved in microreactor because of the reduced worker exposure to huge volumes of hazardous chemicals and flammable gases, like H2. Third, only a small amount of energy and material is needed in a microreactor and its response time is shorter; hence, more information per space and time makes the microreactor a faster and cost-saving tool for screening of processes, and materials, like catalysts. Lastly, the numbering up approach to scale up enables faster and cost-effective transfer of research results into production. Moreover, its excellent performance is not expected to diminish upon scale-up. In summary, microreactor is an ideal tool for catalyst evaluation for the gas−liquid−solid three phase hydrodeoxygenation reactions. Because of the lack of experimental data on the performance of replacement catalysts on hydrodeoxygenation of actual biooil feedstock, the present study aims to evaluate the conventional hydrotreating catalyst, NiMo/γ-Al2O3, for green diesel production from microalgae oil under different reaction parameters, including reactor inner diameter, hydrogen to oil ratio, residence time, reaction temperature, and pressure. The conditions under which external and internal mass transfer limitations exist will also be explored. In this work, the catalyst is evaluated in terms of its activity, selectivity, and life, which will be quantitatively expressed as changes of product yield and 263


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distribution and as reactant conversion with time. Data collected from microreactor system will guide further catalyst screening and is also needed for process optimization and scaleup to monolith reactor.

Table 1. Chemical and Physical Characterization of Microalgaea Oil density (298 K) (g/mL) elemental analysis C (wt %) H (wt %) N (wt %) O (wt %) S (ppm) P (ppm) ash (wt %) water content (KF titration) (wt %) oil compound in microalgae oil monoacylglycerides (wt %) diacylglycerides (wt %) triacylglycerides (wt %) free fatty acids (wt %) total fatty acids (wt %)

2. EXPERIMENTAL SECTION 2.1. Catalyst. A commercial catalyst NiO/MoO3/γ-Al2O3 provided by Albemarle (presulfided and supplied by Eurecat, U.S.A.) Houston, TX was used in this study. A typical formulation of the commercial hydrotreating catalyst is 4.03 wt % NiO and 13.20 wt % MoO3 on dry basis.24 The catalyst particles were provided in the quadralobe form, and were ground and sieved to obtain particles of 75−150 μm diameter range. Previous studies in our research group22,23,25−27 have shown that the influence of internal diffusion on the reaction rate was negligible for this catalyst particle size range. The average BET-surface area of the sieved catalyst was 164 m2/g and the average pore diameter was 106 Å.25 After being packed into the reactor, in situ reduction with pure H2 was done before each experimental run at 310 °C and 500 psig for 2 h. 2.2. Liquid Feed. Nannochloropsis salina was chosen to start with because it is the workhorse of the industry, and it also has a high oil content in comparison with most algae strains. Its extracted oil was kindly provided by Valicor Renewables, LLC Inc. Microalgae oils are completely different from vegetable oils. Vegetable oils are predominantly (∼90−98 wt %) neutral lipids, mostly triglycerides, with a small amount of polar lipids that could be removed by simple aqueous degumming. In contrast, in addition to the neutral lipids (>30 wt %), crude algae oils also contain a significant amount of polar lipids and undetermined natural substances. The polar lipids are usually in the form of phospholipids (10−40 wt %) and glycolipids (10− 40 wt %), depending on the algae strain. Due to this high content of polar lipids, aqueous degumming has been proven to be unviable in removing polar lipids in algae oil. It should be noted that there are heteroatoms (P, N, and S) carried in the polar heads of the polar lipids, which are highly effective in deactivating catalysts and destroying their longevity. Therefore, Valicor’s patented wet extraction technology was applied to drive catalyst poisons into the aqueous phase, hence separating them from the oil to make the algae oil better suited for catalytic conversion to fuels. For the algae oil used in this study, the neutral lipids could be identified and quantified, and an oil composition was determined by Valicor (Table 1). The analysis of undetermined natural substances is still ongoing and the elemental composition analysis could shed some light on it. Calculations were made based on the elemental analysis by Robertson Microlit Laboratories, Ledgewood, NJ (Table 1), and the results are as indicated in Table 2. It can be concluded that the unidentified portion has nearly the same elemental composition as the fatty acids, but slightly higher degree of unsaturation (Table 2). This will help to narrow the search range in future identification work. Although the unidentified compounds account for 52.55%, for this hydrotreating process, the objective is to lower the oxygen content of liquid product to levels that can be processed by the conventional hydrotreating unit of the refinery, if necessary. Therefore, the complete composition profile is of minor importance in product evaluation. In this study, because diluted feed solution (1.3 wt %) was processed, the oxygen content of the liquid product is below the detection limit of

a

0.807 76.29 11.22 0.43 12.06 2033 246 0.34 0.86 4.41 4.56 29.23 9.25 47.45

Nannochloropsis salina.

Table 2. Carbon, Oxygen, and Hydrogen Elemental Compositions in Identified and Unidentified Oil

fatty acids

unidentified

C O H C O H

based on whole microalgae oil

based on respective elementary composition of microalgae oil

37.50% 5.86% 5.90% 38.79% 6.20% 5.32%

49.16% 48.56% 52.57% 50.84% 51.44% 47.43%

direct oxygen content analysis. Therefore, the hydrocarbons (straight chain alkanes) produced upon oxygen removal were quantified and used as an indication of product quality. The ultimate goal of this project is to coprocess microalgae oil with gas oil, which is the intermediate to produce diesel in refinery. Gas oil is a complex hydrocarbon mixture, which cannot be distinguished from the hydrocarbons produced from algae oil and, hence, make it impossible to quantify the produced hydrocarbons. Moreover, gas oil has significant sulfur content, which may affect the hydrodeoxygenation process. In order to reduce the complexity and uncertainty in experiment and analysis, dodecane is selected as a substitute for gas oil, which is a liquid alkane with the chemical formula C12H26. Limited by the catalyst loading (milligrams) in microreactor, if high oil concentration was used, the concentration of unconverted triglycerides would be high in liquid product, which would lead to the inaccuracy in measuring hydrocarbons. Therefore, 1 g of algae oil, which is in semisolid form at room temperature with low flowability, is dissolved in 100 mL of dodecane to make 1.3 wt % solution. After 1 h of sonication for better mixing, this solution was vacuum-filtered using 410 filter paper (VWR, 1 μm retention) to remove any residual particles to avoid blockage of the microreactor. 2.3. Experimental Setup and Procedure. A schematic of the setup for the hydrotreating of microalgae oil is shown in Scheme 1. The reaction was conducted in a continuous flow SS316L microreactor with an inner diameter of 0.762 mm. The ground NiMo catalyst was packed in the microreactor, and the reactor length was varied from 3 to 19.2 cm depending on the residence time required for the experiments. Fresh presulfided NiMo catalyst was used for all runs and was first reduced using 264


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Scheme 1. Schematic of Microalgae Oil Hydrodeoxygenation Setup

pure hydrogen with a reactor temperature of 310 °C and pressure of 500 psig for 2 h before making any reaction run. Two pieces of 1/8 in. outer diameter tubes were packed with glass beads (90 μm) and connected to the entrance and exit of the microreactor. Hastelloy micron filter-cloth (200 × 1150 meshes, Unique Wire Weaving Co., Hillside, NJ) was placed at the end of each tube section to retain the packed particles. The reactor was placed in a temperature-controlled furnace, and the temperature indicated by the controller was confirmed by measuring the temperature inside the furnace using a hand-held temperature meter. Pressures at the inlet and outlet of the microreactor were measured and were adjusted by a back pressure regulator (BPR) installed downstream. The difference of the two measured pressure values was due to the pressure drop through the catalyst bed and was used as an indication of extent of reactor blockage. Compressed hydrogen was regulated by a mass flow controller (sccm), and an Agilent ADM 2000 Universal gas flowmeter was also used to verify gas flow rate before and after the reactor. Studies have suggested that introduction of gas and liquid flows head to head could generate a short slug length and, thus, improve the mixing in two-phase flows.28 Therefore, liquid feed was pumped by a HPLC pump at different setting rates (mL/min) and then contacted with hydrogen at the T-mixer before entering the reactor. A sample loop was used due to low product flow rate. Liquid product samples were collected for analysis at ambient conditions from the sampling port as shown in the schematic of the setup. The samples were collected at 1 h intervals for offline analysis, and seven samples were obtained from one experimental run. The remaining product stream flowed to the gas−liquid separator, and liquid would be condensed in the separator, whereas gases, including unconverted hydrogen and gas product would be vented to the hood. After each experimental run, nitrogen would be used as purge gas to clean the system, and the BPR was also cleaned by compressed air to remove any liquid or solid residues. 2.4. Product Analysis. A preliminary analysis showed that the C13 to C20 hydrocarbons were the primary products of microalgae oil hydrodeoxygenation, a result that is consistent

with the observation that the C14, C16, C18, C20 fatty acids are the main components of natural lipids in algae oil. Therefore, the product analysis focused on the identification and quantification of C13 to C20 hydrocarbons by a Varian 450 Gas Chromatograph (GC) equipped with CP-8400 Autosampler. The liquid sample eluted by helium was separated on a ZB-1HT nonpolar capillary column (30 m × 0.25 mm × 0.25 μm) and detected using a flame ionization detector (FID). The temperature program was: initially 50 °C for 1 min, then ramped at 15 °C/min to 240 °C, and finally held for 5 min. Before the analysis of each batch of samples, four GC standards which contained C13 to C20 hydrocarbons at four different levels were run first for GC external calibration. The response factor, which is the ratio of peak area to hydrocarbon concentration, could be used for quantification. The retention time was used for identification of the hydrocarbon. The identification was further confirmed by adding standard hydrocarbon to samples to observe the peak area increase. The C13 to C20 hydrocarbon yield is calculated based on the quantified total fatty acids (47.45%, Table 1) in the liquid feed, which is defined as C13 to C20 hydrocarbon yield total mass of C13 − C20 hydrocarbon in product = total mass of fatty acids in liquid feed (1)

Space−time yield (STY) of hydrocarbon (rate of hydrocarbon formation) is defined as STYhydrocarbon =

mass of hydrocarbon produced/time mass of catalyst

(2)

Even-numbered carbon hydrocarbon to odd-numbered carbon hydrocarbon ratio is defined as HC(2n)/HC(2n − 1) mass of even‐numbered carbon hydrocarbon in product = mass of odd‐numbered carbon hydrocarbon in product (3) 265


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in fatty acids to produce an oxygenated intermediate, an aldehyde which is finally converted either to an even-numbered carbon hydrocarbon through hydrodehydration (with an alcohol intermediate) or an odd number hydrocarbon via hydrodecarboxylation (hydrodecarbonylation). Because the conversion of the fatty acid to the aldehyde is the rate-limiting step3 effort should always be focused on the fatty acids hydrogenation to improve the overall hydrocarbon yield. It should also be noted that the water, CO, and CO2 produced could also react with H2 in methanation and water−gas shift reactions. The formation of methane will increase the hydrogen consumption significantly, which will make the process less economically effective. Therefore, an ideal catalyst should be effective in suppressing the methanation reaction. According to this reaction pathway, based on triglycerides, calculations show that the hydrodehydration route produces the theoretical maximum yield, the values of which are approximately 85% and 95% for hydrocarbon and carbon respectively, whereas the minimum theoretical values for the yield are approximately 80% for hydrocarbon and 90% for carbon corresponding to the hydrodecarboxylation or hydrodecarbonylation route.

Identification and quantification of unconverted fatty acids in liquid product is required for the calculation of lipids conversion. Polar carboxyl functional groups in fatty acids need to be neutralized before being analyzed in GC to avoid adsorption to the column and to obtain better separation. A modified AOCS method was used for sample derivatization. After sample preparation, all fatty acids in the sample were converted to fatty acid methyl ester (FAME). The FAME sample was also analyzed in GC-FID under the conditions shown below: Column: SP2560 column (100 m × 0.25 mm × 0.20 μm) Oven: 140 °C (5 min), 4 °C/min to 240 °C (30 min.) Injector: 250 °C Detector: FID, 280 °C Carrier gas: Helium, constant flow = 1.8 mL/min Injection: 1 μL, 1:20 split In this FAME analysis method, external standards were used for peak identification, and a similar method was applied for FAME quantification as for hydrocarbon described above. Because the hydrogenolysis of saturated triglycerides is slower than double bond saturation, the conversion of lipids could be approximately calculated based on measured fatty acids, and all reported conversion data is calculated according to the definition

3. RESULTS AND DISCUSSION 3.1. Role of Reactor Inner Diameter (ID). It has been discussed in the literature29 that the flow characteristics in microreactor (ID < 10−3 m) are different from those in conventional reactors and, therefore, are expected to affect the catalyst activity. In order to better understand the advantage of microreactor in this three-phase reaction, reactors with ID of 0.762 mm, 1.753 mm, and 4.572 mm were compared by evaluating the effect of reactor ID on space-time yield (STY) of C13 to C20 hydrocarbons, and microalgae oil conversion based on measured fatty acids. Reaction temperature, pressure, and catalyst loading were kept the same for all runs. In addition, superficial flow velocity and gas hourly space velocity (GHSV) were kept constant by varying gas flow rate. For the larger ID reactors with increased volume, the same amount of catalyst was packed into the reactor but the catalyst was premixed thoroughly with glass beads (90 μm) to fill the extra volume. The results summarized in Figure 1 show that microalgae oil conversion and STY increase as the reactor ID reduces. This increase could be attributed to the enhancement of mass

conversion mass of fatty acids in feed − mass of fatty acids in product = mass of fatty acids in feed (4)

2.5. Reaction Mechanism. The mechanism of hydrodeoxygenation of triglycerides was established using a wide range of model compounds by Peng et al.,3 and a detailed reaction pathway is summarized as shown in Scheme 2. The first step is the saturation of double bonds in the alkyl chain, which occurs very fast; then the second step is the cleavage of C−O bond in the hydrogenolysis of saturated triglycerides to produce fatty acids and propane, which is slower than the first step. The next step is the hydrogenation of the carboxylic group Scheme 2. Reaction Pathways Involved in Conversion of Triglycerides into Hydrocarbons.3

Figure 1. Effect of reactor ID on STY of C13 to C20 hydrocarbon and microalgae oil conversion based on measured fatty acids (46 mg of presulfided NiMo, 1.32 wt % Nannochloropsis salina oil in dodecane, temperature = 310 °C, H2 pressure = 500 psig, superficial flow velocity = 0.117 m/s, GHSV = 55 000 h−1). 266


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Scheme 3. External and Internal Mass Transfer Analysis for Taylor Flow

transfer in Taylor flow regime in microreactor. Because the flow velocity was kept the same in all three reactors, convective mass transfer enhancement in the microreactor can only be attributed to the existence of Taylor flow, which is not present in the large diameter reactors. It has been reported that the effect of shear stress caused by friction is amplified in microreactor when compared to a macroreactor, and this generates internal convective recirculation within the liquid slugs (Scheme 3).30 This has been confirmed through flow visualization experiments, and transitional slug flow was observed with liquid slug boundaries broken up by the catalyst particles.31 The recirculation within the liquid slug will decrease the diffusional path length down to the tube radius, which is generally much shorter than the slug half length.29 According to equation of continuity, the diffusion rate is inversely proportional to the diffusion pathway length, and hence, the diffusion rate would be enhanced greatly in microreactor than in macroreactor. Therefore, for multiphase reactions that are usually limited by mass transfer, the microreactor is recommended both for performance study and kinetics modeling due to its enhanced mass transfer compared with conventional reactors. In this work, all the other catalyst evaluation runs were conducted in the microreactor of 0.762 mm ID. 3.2. Mass Transfer Analysis. 3.2.1. Analysis of External Mass Transfer Limitation. According to fluid flow characterization results in the literature,22 all the experiments we conducted for catalyst evaluation were performed in Taylor flow regime, comprising of alternating liquid slugs (blank) and gas bubbles (blue), as indicated in Scheme 3. Because the reaction occurs only at the active sites on the catalyst surface, the hydrodeoxygenation of microalgae oil reaction on NiMo catalyst involves a three-step external mass transfer as shown in Scheme 3: first, the transfer of hydrogen from the bulk gas phase to gas−liquid interface, the cap of the hydrogen bubble; then the diffusion of hydrogen into the bulk liquid phase; finally, the transfer of dissolved hydrogen and microalgae oil to the external surface of the catalyst through the solid−liquid boundary layer. Because the gas phase is almost pure hydrogen and the gas− liquid interface is saturated with this gaseous reactant, the main mass transfer resistances occur at the gas−liquid and liquid− solid interfaces, with the former being predominant. As the transfer rates through interfaces depend greatly on flow velocity, the overall external mass transfer limitation can be studied experimentally by examining the change of hydrocarbon yield rate at various superficial flow velocities by increasing gas and liquid flow rate but keeping the gas to liquid ratio unchanged, and the residence time constant. The latter is achieved by varying the catalyst loading. The STY of C13 to C20 hydrocarbon at different superficial flow velocities is depicted in Figure 2. The result shows that the reaction rate is independent of the flow velocity within the studied range. This indicates that the reaction is not limited by external mass transfer even at the lowest superficial flow velocity of 0.117 m/

Figure 2. Effect of superficial flow velocity on space-time yield of C13 to C20 hydrocarbon. (Presulfided NiMo, 1.32 wt % Nannochloropsis salina oil in dodecane, temperature = 310 °C, H2 pressure = 500 psig, residence time = 1 s, H2/oil = 1000 SmL/mL).

s. This velocity was chosen as the baseline for all subsequent the following experimental studies. 3.2.2. Analysis of Internal Mass Transfer Limitation. As shown in Scheme 3, internal mass transfer refers to the diffusion of reactants from the catalyst pellet pore entrance into the pore. The reaction is considered as internal mass transfer limited when the reactants concentration inside the pores is much lower than that at the entrance. This variation in concentration throughout the catalyst pellet can be evaluated by calculating the Weisz-Prater parameter,32 CWP C WP

′ −rA(obs) ρp R2 actual reaction rate = = a diffusion rate DeCAs

(5)

where −rA(obs) ′ = 1.11 × 10−4 g of hydrocarbon/g of cat./s is the observed reaction rate. ρP = 0.29 g of cat./cm3 is the catalyst pellet density. R = 112.5 × 10−4 cm is the characteristic diameter of the catalyst pellets. CAS = 360 g/m3 is the hydrogen concentration at the external surface of the catalyst pellet, which can be obtained from the hydrogen solubility data in dodecane, approximately. De = the effective diffusivity of hydrogen in the catalyst pellets, which can be calculated using the equation: De = (DABφPσC)/τ, where DAB = 1.12 × 10−7 m2/s is the binary diffusivity of H2 in dodecane;33 φP = 0.4−0.6 is the porosity of the catalyst pellets; σC= 0.7−0.8 is the catalyst constriction factor; τ = 2−8 is the catalyst tortuosity.32 Therefore, De is 3.92 × 10−9 to 2.69 × 10−8 m2/s. The Weisz−Prater parameter is calculated to be between 0.042 and 0.289 depending on effective diffusivity value. By the definition, CWP ≪ 1 means reaction rate is much smaller than 267


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the diffusion rate, which indicates that hydrogen concentration gradient is negligible inside the catalyst pores and the reaction is not internal-diffusion-limited. 3.3. Evaluation of Catalyst. Based on the evaluation of the effect of reactor ID and the mass transfer analysis results, the catalyst was evaluated in a nonmass transfer limited range in which the measured product yield and reactant conversion is a reflection of catalyst activity, with the product distribution reflecting catalyst selectivity. The effect of industrially used operating parameters, namely hydrogen to oil ratio, residence time, reaction temperature, and pressure will be discussed. 3.3.1. Effect of Hydrogen to Oil Ratio. The hydrogen to oil ratio is an important operating parameter in hydrotreating processes. The conventional H2/oil ratio refers to the ratio of hydrogen feed (standard temperature and pressure, STP) to the total liquid feed (solvent included).34 In the present study, experiments were conducted to investigate the effect of H2/oil ratio between 1000 SmL/mL to 385 SmL/mL by adjusting liquid flow rate. Reaction temperature, pressure, GHSV, and catalyst loading were kept constant. Because the actual hydrogen flow rate was at least 24 times higher than the liquid flow rate, the effect of liquid flow rate variation on superficial flow velocity was negligible; hence, the residence time remained the same at all studied H2/oil ratios. As shown in Figure 3a, during the 7 h-long reaction, hydrocarbon yield increased with the increase in H2/oil ratio. For H2/oil ratio less than 1000 SmL/mL, the yields decreased rapidly after 1 h reaction and were less than 10% after 3 h. However, there was no observable change in yield for H2/oil ratio of 1000 SmL/mL for at least the next 4 h. The same trend is observed in the microalgae oil conversion results (Figure 3b). During the experiment, it was observed that the liquid slug length decreased with the increase of H2/oil ratio. Because the gas to liquid mass transfer is a strong function of the liquid slug length,35 the increase of H2/oil ratio will enhance the gas− liquid convective mass transfer, therefore resulting in high yield of product. In contrast, more intermediates were formed at lower H2/oil ratio because of insufficient amount of hydrogen. As is stated above, because the conversion of fatty acids is the rate-limiting step, the formed intermediates were mainly fatty acids, which had low solubility in dodecane and would precipitate on the catalyst surface. Due to the accumulation of oxygenated intermediates on the active sites of the catalyst, the catalyst was, in general, deactivated. However, at elevated H2/oil ratio, catalyst activity was high and seemed to decrease with time at a moderate rate. These results show that hydrogen acts not only as reactant but it also serves to protect the hydrotreating catalyst by preventing the adsorption of the oxygenates on the catalyst surface. Catalyst was combusted after 7 h on stream using dilute air at 650 °C for 1 h, and gas product from combustion was analyzed. Because no CO or CO2 peak was observed in GC-TCD, it can be concluded that coke formation was not significant during the reaction. The ratio of even-numbered carbon hydrocarbon to oddnumbered carbon hydrocarbon (Figure 3c) changes with onstream time and follows the same trend as the product yield and reactant conversion, which indicates that the catalyst selectivity correlates with catalyst activity. Therefore, to keep this ratio constant for product quality control, catalyst deactivation needs to be prevented. 3.3.2. Effect of Residence Time. The definition of residence time in this study is the volume of the reactor divided by the superficial volumetric flow rate of the reactants, which is

Figure 3. Effect of H2/oil ratio on (a) C13 to C20 hydrocarbon yield, (b) microalgae oil conversion based on measured fatty acids, and (c) ratio of even-numbered carbon hydrocarbon to odd-numbered carbon hydrocarbon (46 mg of presulfided NiMo, 1.32 wt % Nannochloropsis salina oil in dodecane, temperature = 310 °C, H2 pressure = 500 psig, residence time = 1s, GHSV = 55 000 h−1).

calculated based on the empty bed assumption. The residence time was changed by varying reactor length from 3 to 18 cm, as well as the catalyst loading, while keeping all fluid flow rates unchanged. Other experimental conditions, such as temperature, pressure, superficial flow velocity were held constant. The H2/oil ratio of 1000 SmL/mL was held fixed in all runs to eliminate the convective mass transfer influence. Therefore, weight hourly space velocity (WHSV) was decreased with the increase of residence time. The results in Figure 4a−c show that hydrocarbon yield, microalgae oil conversion, and ratio of even-numbered carbon hydrocarbon to odd-numbered carbon hydrocarbon increases with the increase of residence time (between 0.25 to 1.5 s) because the catalyst loading increases. 268


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280−360 °C. Reaction pressure, fluid flow rates and WHSV were kept the same for all runs with the H2/oil ratio of 1000 SmL/mL. Figure 5 shows the performance of the catalyst under

Figure 4. Effect of residence time on (a) C13 to C20 hydrocarbon yield, (b) microalgae oil conversion based on measured fatty acids, and (c) ratio of even-numbered carbon hydrocarbon to odd-numbered carbon hydrocarbon (presulfided NiMo catalyst, 1.32 wt % Nannochloropsis salina oil in dodecane, temperature = 310 °C, H2 pressure = 500 psig, superficial flow velocity = 0.117 m/s).

Figure 5. Effect of reaction temperature on (a) C13 to C20 hydrocarbon yield, (b) microalgae oil conversion based on measured fatty acids, and (c) ratio of even-numbered carbon hydrocarbon to odd-numbered carbon hydrocarbon (46 mg of presulfided NiMo catalyst, 1.32 wt % Nannochloropsis salina oil in dodecane, H2 pressure = 500 psig, WHSV = 0.65 g of oil/g of cat./h).

Also, all studied performance parameters decline more slowly at longer residence time, which indicates that per unit mass of catalyst, less oxygenated intermediates were produced and accumulated on catalyst surface. It can be concluded that when same amount of reactant is fed into the reactor, the catalyst activity and its stability are a function of catalyst loading. Therefore, decreasing the WHSV is a way to maintain the overall catalyst activity, but productivity will be sacrificed at the same time. 3.3.3. Effect of Reaction Temperature. Temperature is an important operating parameter when evaluating catalysts. Experiments were carried out in the temperature range of

four different reaction temperatures; it is evident that 280 °C is too low for this hydrodeoxygenation process. When the temperature increases from 310 to 360 °C the resulting effect is negligible during the first 4 h, but subsequently, the stability of the catalyst activity correlates with reaction temperature. At 360 °C, there is no sign of reduction in the hydrocarbon yield for 7 h. It seems that at elevated temperatures, it may be difficult for the oxygenates to adsorb on the catalyst surface. As 269


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shown in Figure 5a and b, a high reaction temperature can help to maintain catalyst life and activity. From Figure 5c, it can be observed that the HC(2n)/HC(2n − 1) ratio increases as temperature increases from 280 to 310 °C. This indicates that for the overall hydrodeoxygenation reaction, hydrodehydration appears to be favored at higher temperatures for presulfided NiMo/γ-Al2O3 catalyst. Also, the ratio HC(2n)/HC(2n − 1) becomes less sensitive to temperatures at 310 °C and above, which indicates that the catalyst selectivity is not temperaturedependent in this range. 3.3.4. Effect of Reaction Pressure. The reaction pressure was controlled by adjusting the back pressure regulator to maintain a constant total reaction pressure within the reactor system. In this study, reactions were run at 300 psig, 400 psig and 500 psig, and each run was of a duration of 7 h. The reaction temperature, fluid flow rates and WHSV were kept the same for all runs at H2/oil ratio of 1000 SmL/mL. It should be noted that because the gas flow rates were measured under standard conditions, the actual gas flow rates were pressure dependent; hence, the residence time increased with increase in pressure. As shown in Figure 6a and b, the reaction pressure has negligible effect on the initial catalyst activity. However, the hydrocarbon yield and microalgae oil conversion decrease in the following order during the subsequent on-stream hours: 500 psig > 300 psig > 400 psig. Similarly, Figure 6c shows that the HC(2n)/HC(2n − 1) ratio also decreases following the same order. This could be explained by the effect of hydrogen partial pressure. During hydrotreating, Yang et al. found that higher hydrogen partial pressure was favorable to the DHYD reaction because more adsorbed hydrogen on the surface active sites promoted DHYD reaction but impeded the DCO/DCO2 reaction.34 A similar result was also obtained by Guzman et al.36 and Krár et al.37 At 500 psig, hydrogen partial pressure is elevated and residence time is longer than that under the lower pressure conditions at which enhanced catalyst activity was measured as expected. At 400 psig, the HC(2n)/HC(2n − 1) ratio is the lowest, which indicates relatively higher preference for the DCO/DCO2 reaction route than at the other two reaction pressures. Although pure hydrogen was fed into the reactor, non-negligible amounts of propane, water, carbon monoxide, carbon dioxide, and methane were produced.38 These gases will decrease the hydrogen partial pressure and inhibit the catalyst activity; more gases will be formed as a result. At 300 psig, the production of gases is reduced compared to that at 400 psig and, therefore, will have less effect on hydrogen partial pressure. However, at lower hydrogen partial pressure, residence time is reduced and catalyst deactivated faster due to the accumulation of oxygenated intermediates that are produced even in greater amount at the low pressure. In conclusion, hydrogen partial pressure is an important factor for maintaining catalyst life and activity. Therefore, it is not surprising that in the petroleum industry, hydrotreating is conducted at 700 psig to prolong catalyst life. Gaseous products need to be handled properly if treated gas is to be recycled.38

Figure 6. Effect of reaction pressure on (a) C13 to C20 hydrocarbon yield, (b) microalgae oil conversion based on measured fatty acids, and (c) ratio of even-numbered carbon hydrocarbon to odd-numbered carbon hydrocarbon (46 mg of presulfided NiMo catalyst, 1.32 wt % Nannochloropsis salina oil in dodecane, temperature = 310 °C, WHSV = 0.65 g of oil/g of cat./h).

of the microreactor compared to a macroreactor for this threephase reaction. Due to the enhanced mass transfer characteristics in microreactor, the STY of hydrocarbon was significantly higher in a microreactor (ID < 1 mm). On the basis of the external and internal mass transfer limitation analysis, the catalyst was evaluated by studying changes in the microalgae oil conversion and the product yield and composition as a function of operating conditions: hydrogen to oil ratio, residence time, reaction temperature, and pressure. The results revealed that the conventional hydrotreating catalyst, presulfided NiMo/γAl2O3, is deactivated due to the accumulation of produced oxygenated intermediates in hydrodeoxygenation reaction, and its selectivity to even-numbered carbon hydrocarbon produced from hydrodehydration correlates with the catalyst activity. The

4. CONCLUSIONS Reduced presulfided NiMo/γ-Al2O3 catalyst has been evaluated for green diesel production from microalgae (Nannochloropsis salina) oil via hydrodeoxygenation. In the present study, packed-bed reactors of three different inner diameters were compared to confirm the superior mass transfer characteristics 270


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(6) Vam, A. Kinetics of the Hydro-Deoxygenation of Stearic Acid over Palladium on Carbon Catalyst in Fixed-Bed Reactor for the Production of Renewable Diesel; University of Dayton: Dayton, OH, 2013. (7) James, O. O.; Maity, S.; Mesubi, M. A.; Usman, L. A.; Ajanaku, K. O.; Siyanbola, T. O.; Sahu, S.; Chaubey, R. A review on conversion of triglycerides to on-specification diesel fuels without additional inputs. Int. J. Energy Res. 2012, 36 (6), 691−702. (8) Donnis, B.; Egeberg, R. G.; Blom, P.; Knudsen, K. G. Hydroprocessing of bio-oils and oxygenates to hydrocarbons. Understanding the reaction routes. Top. Catal. 2009, 52 (3), 229−240. (9) Holmgren, J.; Gosling, C.; Couch, K.; Kalnes, T.; Marker, T.; McCall, M.; Marinangeli, R. Refining biofeedstock innovations. Pet. Technol. Q. 2007, 12 (4), 119. (10) Holmgren, J.; Gosling, C.; Marinangeli, R.; Marker, T.; Faraci, G.; Perego, C. New developments in renewable fuels offer more choices. Hydrocarbon Process. 2007, 86 (9), 67−72. (11) Laurent, E.; Delmon, B. Study of the hydrodeoxygenation of carbonyl, carboxylic and guaiacyl groups over sulfided CoMo/γ-Al2O3 and NiMo/γ-Al2O3 catalysts. I: Catalytic reaction schemes. Appl. Catal., A 1994, 109 (1), 77−96. (12) Laurent, E.; Delmon, B. Study of the hydrodeoxygenation of carbonyl, carboxylic and guaiacyl groups over sulfided CoMo/γ-Al2O3 and NiMo/γ-Al2O3 catalyst. II: Influence of water, ammonia and hydrogen sulfide. Appl. Catal., A 1994, 109 (1), 97−115. (13) Šimácě k, P.; Kubička, D.; Šebor, G.; Pospíšil, M. Hydroprocessed rapeseed oil as a source of hydrocarbon-based biodiesel. Fuel 2009, 88 (3), 456−460. (14) Şenol, O.; Ryymin, E.-M.; Viljava, T.-R.; Krause, A. Effect of hydrogen sulphide on the hydrodeoxygenation of aromatic and aliphatic oxygenates on sulphided catalysts. J. Mol. Catal. A: Chem. 2007, 277 (1), 107−112. (15) Şenol, O. I.̇ Hydrodeoxygenation of aliphatic and aromatic oxygenates on sulphided catalysts for production of second generation biofuels. D.S. Dissertation, Helsinki University of Technology, Espoo, Finland, 2007. (16) Toba, M.; Abe, Y.; Kuramochi, H.; Osako, M.; Mochizuki, T.; Yoshimura, Y. Hydrodeoxygenation of waste vegetable oil over sulfide catalysts. Catal. Today 2011, 164 (1), 533−537. (17) Kubička, D.; Kaluža, L. Deoxygenation of vegetable oils over sulfided Ni, Mo and NiMo catalysts. Appl. Catal., A 2010, 372 (2), 199−208. (18) da Rocha Filho, G.; Brodzki, D.; Djéga-Mariadassou, G. Formation of alkanes, alkylcycloalkanes and alkylbenzenes during the catalytic hydrocracking of vegetable oils. Fuel 1993, 72 (4), 543−549. (19) Huber, G. W.; O’Connor, P.; Corma, A. Processing biomass in conventional oil refineries: production of high quality diesel by hydrotreating vegetable oils in heavy vacuum oil mixtures. Appl. Catal., A 2007, 329, 120−129. (20) Roy, S.; Bauer, T.; Al-Dahhan, M.; Lehner, P.; Turek, T. Monoliths as multiphase reactors: a review. AIChE J. 2004, 50 (11), 2918−2938. (21) Okafor, O. C.; Tadepalli, S.; Tampy, G.; Lawal, A. Microreactor Performance Studies of the Cycloaddition of Isoamylene and αMethylstyrene. Ind. Eng. Chem. Res. 2010, 49 (12), 5549−5560. (22) Tadepalli, S.; Halder, R.; Lawal, A. Catalytic hydrogenation of onitroanisole in a microreactor: Reactor performance and kinetic studies. Chem. Eng. Sci. 2007, 62 (10), 2663−2678. (23) Tadepalli, S.; Qian, D.; Lawal, A. Comparison of performance of microreactor and semi-batch reactor for catalytic hydrogenation of onitroanisole. Catal. Today 2007, 125 (1), 64−73. (24) Absi Halabi, M.; Stanislaus, A.; Qamra, A.; Chopra, S. Effect of presulfiding on the activity and deactivation of hydrotreating catalysts in processing Kuwait vacuum residue. Stud. Surf. Sci. Catal. 1996, 100, 243−251. (25) Joshi, N.; Lawal, A. Hydrodeoxygenation of pyrolysis oil in a microreactor. Chem. Eng. Sci. 2012, 74, 1−8. (26) Joshi, N.; Lawal, A. Hydrodeoxygenation of 4-Propylguaiacol (2Methoxy-4-propylphenol) in a Microreactor: Performance and Kinetic Studies. Ind. Eng. Chem. Res. 2013, 52 (11), 4049−4058.

catalyst activity and life can be preserved by elevating hydrogen to oil ratio, residence time, reaction temperature, and pressure, which will decrease the adsorption of oxygenates on the catalyst surface. Reaction temperature (in the range of 310 to 360 °C) and pressure between 300 and 500 psig have negligible effect on initial catalyst activity. Hydrogen acts not only as a reactant, but it also serves to protect the hydrotreating catalyst. Weight hourly space velocity (WHSV) is reduced when residence time is increased by increasing catalyst loading; therefore, the packed bed is exposed to a decreasing amount of reactant per gram catalyst. It should be noted that the formation of gaseous product, such as CO, CO2, CH4, propane, and water, decreases the hydrogen partial pressure which leads to poor catalyst activity. The optimum conditions for hydrotreating of microalgae oil were determined to be 500 psig H2, 360 °C, H2/oil ratio of 1000 SmL/mL, and residence time of 1 s. For this set of conditions the initial catalyst activity was maintained without any signs of deactivation for at least 7 h and the obtained C13 to C20 hydrocarbon yield was 56.2%, with a carbon yield of 62.7%, nearly complete conversion (98.7%) of microalgae oil, and HC(2n)/HC(2n − 1) ratio of 6. It is expected that this work will provide baseline data which will guide future work on formulation of catalysts for the hydrotreating of microalgae oil.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Phone: 201-216-8314. Fax: 201216-8306. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS

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REFERENCES

This research was funded by the DOE-Bioenergy Technologies Office through grant DE-EE0006063. The authors are thankful for their support. We would also like to acknowledge the support of our principal industrial partner Valicor Renewables LLC, and we are especially grateful to Dr. Brian Goodall for his technical contributions to the study. Discussions on product analysis with Dr. Peter Domaille, formerly of Synthetic Genomics, were very helpful. Special thanks to Dr. James Manganaro for his inspiring guidance, and the painstaking review of the experimental results as well as the manuscript. We would also like to acknowledge Dr. Moinuddin Malik for his review of the manuscript. The authors, of course, bear the sole responsibility for any errors.

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