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Evaluation of Presulfided NiMo/γ-Al2O3 for Hydrodeoxygenation ofMicroalgae Oil To Produce Green DieselLin Zhou* and Adeniyi Lawal
New Jersey Center for Microchemical Systems, Department of Chemical Engineering and Materials Science, Stevens Institute ofTechnology, 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 evaluatedfor green diesel production via hydrodeoxygenation of unrefined microalgae oil in a microreactor, mimicking the single channelof a monolithic reactor. The effect of reactor inner diameter on space-time yield of hydrocarbon and microalgae oil conversionwas studied first to confirm the superiority of the microreactor for three-phase reactions. Based on the external and internal masstransfer 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 inhydrodeoxygenation reaction, and its selectivity to even-numbered carbon hydrocarbon produced from hydrodehydrationcorrelates with the catalyst activity. The catalyst activity and life can be preserved by increasing hydrogen to oil ratio, residencetime, reaction temperature, and pressure, which will decrease the adsorption of oxygenates on the catalyst surface. For thereaction condition: 500 psig H2, 360 °C, H2/oil ratio of 1000 SmL/mL, and residence time of 1 s, the initial catalyst activity wasmaintained without any signs of deactivation for at least 7 h and the obtained C13 to C20 hydrocarbon yield was 56.2%, with acarbon yield of 62.7%, nearly complete conversion (98.7%) of microalgae oil, and HC(2n)/HC(2n − 1) ratio of 6.
1. INTRODUCTIONMicroalgae are ubiquitous, and although they are primarilyfound in all the oceans and seas, an area that covers 71% of theEarth’s surface, they also grow in freshwater bodies as well as onand in soil, rocks, ice, snow, plants, and animals.1 They growextremely 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 cellfactories that convert carbon dioxide to lipids, with an averagelipid content varying between 1% and 70%, and even reachingup to 90% of dry weight under certain conditions.2,4,5 Theaccumulated oil in almost all microalgae is mainly triglycerides(>80%) with a fatty acid profile rich in C16 and C18,2 which isa substantial energy resource for liquid fuel production.Therefore, microalgae oil has been considered as a goodcandidate for low-net carbon liquid transportation fuelproduction, which has no direct competition with edible foodor oil production.However, some of the physical properties of microalgae oil
prohibit its direct use in existing engines, such as lowflowability, high viscosity, and low volatility. In order toimprove the oil’s physical properties while maintaining itsheating value, an ideal upgrading process should only rearrangethe oil molecular structure while avoiding or minimizingcracking.6 On the basis of the experience of vegetable oilupgrading, the two currently practiced oil upgrading methodsare transesterification and hydrotreating, which producebiodiesel and green diesel, respectively. Biodiesel has a lowerviscosity than its parent oil, but still much higher than that ofpetrodiesel. With the cold-flow related issues unresolved,7
biodiesel cannot be used in its pure form; rather, it is blendedwith petrodiesel, the most common product, commerciallyreferred to as B20, blends 20 vol % biodiesel with 80 vol %petrodiesel. Green diesel, on the other hand, is essentially a
mixture of hydrocarbons, has the same chemical properties aspetrodiesel, and is compatible with all existing engines,pipelines, and infrastructure for application and distribution.Compared with biodiesel, green diesel is a more ideal substitutefor petrodiesel.In the production of green diesel from microalgae oil, oxygen
needs to be removed from lipids, mainly triglycerides, toproduce hydrocarbons. Because the sulfur, nitrogen, andphosphorus contents in microalgae oil are really low, themain heteroatom to be removed is oxygen. For this reason, thehydrotreating of microalgae oil mainly refers to the oxygenremoval. 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 theform of H2O and (2) hydrodecarboxylation (DCO2) orhydrodecarbonylation (DCO), in which oxygen is removed inthe form of CO2 and CO, respectively. One triglyceridemolecule 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 ofmuch concern, and no catalysts are specifically formulated foroxygenates hydrotreating. Hence, one of the critical technicalchallenges to make the hydrodeoxygenation of microalgae oilprocess economically feasible is related to the research anddevelopment of effective catalysts. Many studies have beenperformed with commercial hydrotreating catalysts, such asCoMo, NiMo, and NiW supported on alumina, for hydro-treating of natural lipids from different sources and haverevealed that complete triglycerides conversion could be
Received: October 8, 2014Revised: December 5, 2014Published: December 5, 2014
© 2014 American Chemical Society 262 dx.doi.org/10.1021/ef502258q | Energy Fuels 2015, 29, 262−272
achieved over these catalysts.8−13 These conventional hydro-treating catalysts are more active in sulfided form than in oxideform.14 The sulfidation creates active sites that can play a role inthe rupture of the carbon−heteroatom bond.15 Toba et al.studied the hydrotreating of waste cooking oil in both batchreactor (7 MPa, 250−350 °C) and fixed-bed flow reactor (5MPa, 350 °C) using CoMo, NiMo, and NiW catalysts. In theirstudy, the NiMo and NiW catalysts showed high and stablehydrogenation activity, whereas deactivation was observedwhen using the CoMo catalyst. The NiW catalyst favorsmore hydrodecarboxylation or hydrodecarbonylation thanhydrodehydration.16 Senol also reported that the NiMo catalystshowed a higher activity than CoMo in hydrodehydration andhydrogenation reactions.15 Kubicka and Kaluza studied the Ni,Mo, and NiMo sulfided catalysts in deoxygenation of rapeseedoil at 260−280 °C, 3.5 MPa, and 0.25−4 h−1 in a fixed-bedreactor. They found bimetallic NiMo catalysts showed higheractivity and yields of hydrocarbons than monometallic catalyst;NiMo yielded a mixture of decarboxylation and hydro-dehydration products, whereas Ni yielded only decarboxylationproducts, and Mo yielded almost exclusively hydrodehydrationhydrocarbon products.17 Da Rocha Filho et al. studied thehydrotreating of different vegetable oils in a batch reactor usingNiMo catalyst and reported a 66−76 wt % n-alkanes yield after2 h at 360 °C and 14 MPa.18 Huber et al. conducted thehydrotreating of sunflower oil over NiMo catalyst in a fixed-bedreactor under reaction conditions as follows: temperature 350°C, pressure 50 bar H2, LHSV 5.2 h−1, and H2 to oil ratio of1600 SmL/mL. The maximum carbon yield in C15−C18alkanes they obtained was 71%, which is 75% of the maximumtheoretical yield of 95%.19 Peng et al. studied the hydrotreatingof microalgae oil in batch mode with 10 wt % Ni/HBeta at 260°C and 40 bar and obtained 78 wt % yield of liquid alkanes after8 h of reaction time. They also obtained almost identical resultsfrom their trickle-bed reactor system under identical exper-imental conditions.3 As discussed later, in the present study,presulfided NiMo/Al2O3 catalyst was studied for hydrotreatingof microalgae oil as the baseline for further catalyst screeningwork.Although very high oil conversions and hydrocarbon yields
were obtained in the studies mentioned above, which wereconducted in batch reactors or trickle-bed reactors, it should benoted that the space-time yield (STY) was extremely low dueto the severe heat and mass transfer limitations in theconventional macroreactor system. The low STY implies thesize of the reactor and all associated ancillary equipment will beunrealistically large, driving up both the capital and operatingcost. Therefore, a cost-effective upgrading process demands adifferent and innovative approach to reactor design forhydrotreating. Monolith reactors are being studied as asubstitute for conventional multiphase reactors, such astrickle-bed reactors, slurry reactors, and slurry bubble columnreactors for gas−liquid−solid reactions due to their superiorhydrodynamics.20 Monolith substrate is usually made fromceramic, but metallic monolith has also been developed forhighly exothermic reactions, which enable the coupling of thereactor with a heat exchanger to effectively remove heat intransverse direction and control reaction temperature. Themonolith reactor features a honeycomb structure comprisingthousands of finely divided flow passages that extend throughthe whole reactor. The cross-sectional dimensions of eachchannel 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 between0.006 and 0.05 cm, the monolith reactor has correspondingsubmillimeter and, thus, microchannel dimensions, and suchmicrochannels mimic independent microreactors.Many previous studies21−23 of microreactors applied to
different reactions showed that the mass and heat transfer rateswere greatly enhanced and were often several orders ofmagnitude greater than those achievable in conventionalreactors. This is attributed to a combination of short diffusiondistance and the existence of a flow pattern of alternating gasand liquid slugs, also known as Taylor flow, at certain flowconditions. In Taylor flow, the gas bubbles are separated fromthe catalyst surface by a very thin liquid film, the thickness ofwhich is much smaller than the reactor diameter. The bulkliquid is separated by gas bubbles, thus reducing axial mixingbetween the liquid slugs. Also, the recirculation within theliquid slugs improves the radial-mass transfer and mass transferfrom liquid to catalyst as well as from gas to liquid. Thiscombination of good radial-mass transfer and low axial masstransfer in the liquid makes Taylor flow suitable for multiphaseapplications that involve mass transfer or single-phaseapplications that suffer from significant back-mixing. Moreover,as a result of the enhanced mass transfer, the hydrogenconcentration at the catalyst is significantly higher than that atthe conventional reactor, which makes the catalyst to be utilizedmuch more effectively and reduces catalyst deactivation.Therefore, the microreactor, by virtue of its enhanced massand heat transport abilities, allows much shorter residencetimes and, hence, reduced equipment volume and bettercontrol of operating conditions, which should result in aprocess of lower capital and operating expense compared toconventional technology.Microreactor also has advantages over conventional reactors
in the following aspects. First, microreactors contain noagitators and thus no electrical energy input needed to achievecertain flow patterns for better heat and mass transfer. Also, theimproved heat transfer enables improved energy efficiency.Therefore, compared to the batch and semibatch reactors, theenergy consumption is less in microreactor. Second, processsafety is improved in microreactor because of the reducedworker exposure to huge volumes of hazardous chemicals andflammable gases, like H2. Third, only a small amount of energyand material is needed in a microreactor and its response timeis shorter; hence, more information per space and time makesthe microreactor a faster and cost-saving tool for screening ofprocesses, and materials, like catalysts. Lastly, the numbering upapproach to scale up enables faster and cost-effective transfer ofresearch results into production. Moreover, its excellentperformance is not expected to diminish upon scale-up. Insummary, microreactor is an ideal tool for catalyst evaluationfor the gas−liquid−solid three phase hydrodeoxygenationreactions.Because of the lack of experimental data on the performance
of replacement catalysts on hydrodeoxygenation of actual bio-oil feedstock, the present study aims to evaluate theconventional hydrotreating catalyst, NiMo/γ-Al2O3, for greendiesel production from microalgae oil under different reactionparameters, including reactor inner diameter, hydrogen to oilratio, residence time, reaction temperature, and pressure. Theconditions under which external and internal mass transferlimitations exist will also be explored. In this work, the catalystis evaluated in terms of its activity, selectivity, and life, whichwill be quantitatively expressed as changes of product yield and
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distribution and as reactant conversion with time. Datacollected from microreactor system will guide further catalystscreening and is also needed for process optimization and scale-up to monolith reactor.
2. EXPERIMENTAL SECTION2.1. 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 typicalformulation of the commercial hydrotreating catalyst is 4.03 wt% NiO and 13.20 wt % MoO3 on dry basis.24 The catalystparticles were provided in the quadralobe form, and wereground and sieved to obtain particles of 75−150 μm diameterrange. Previous studies in our research group22,23,25−27 haveshown that the influence of internal diffusion on the reactionrate 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 beingpacked into the reactor, in situ reduction with pure H2 wasdone before each experimental run at 310 °C and 500 psig for 2h.2.2. Liquid Feed. Nannochloropsis salina was chosen to start
with because it is the workhorse of the industry, and it also hasa high oil content in comparison with most algae strains. Itsextracted oil was kindly provided by Valicor Renewables, LLCInc. Microalgae oils are completely different from vegetable oils.Vegetable oils are predominantly (∼90−98 wt %) neutrallipids, mostly triglycerides, with a small amount of polar lipidsthat could be removed by simple aqueous degumming. Incontrast, in addition to the neutral lipids (>30 wt %), crudealgae oils also contain a significant amount of polar lipids andundetermined natural substances. The polar lipids are usually inthe form of phospholipids (10−40 wt %) and glycolipids (10−40 wt %), depending on the algae strain. Due to this highcontent of polar lipids, aqueous degumming has been proven tobe unviable in removing polar lipids in algae oil. It should benoted that there are heteroatoms (P, N, and S) carried in thepolar heads of the polar lipids, which are highly effective indeactivating catalysts and destroying their longevity. Therefore,Valicor’s patented wet extraction technology was applied todrive catalyst poisons into the aqueous phase, hence separatingthem from the oil to make the algae oil better suited forcatalytic conversion to fuels.For the algae oil used in this study, the neutral lipids could be
identified and quantified, and an oil composition wasdetermined by Valicor (Table 1). The analysis of undeterminednatural substances is still ongoing and the elementalcomposition analysis could shed some light on it. Calculationswere made based on the elemental analysis by RobertsonMicrolit Laboratories, Ledgewood, NJ (Table 1), and theresults are as indicated in Table 2. It can be concluded that theunidentified portion has nearly the same elemental compositionas the fatty acids, but slightly higher degree of unsaturation(Table 2). This will help to narrow the search range in futureidentification work.Although the unidentified compounds account for 52.55%,
for this hydrotreating process, the objective is to lower theoxygen content of liquid product to levels that can be processedby the conventional hydrotreating unit of the refinery, ifnecessary. Therefore, the complete composition profile is ofminor importance in product evaluation. In this study, becausediluted feed solution (1.3 wt %) was processed, the oxygencontent of the liquid product is below the detection limit of
direct oxygen content analysis. Therefore, the hydrocarbons(straight chain alkanes) produced upon oxygen removal werequantified 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 inrefinery. Gas oil is a complex hydrocarbon mixture, whichcannot be distinguished from the hydrocarbons produced fromalgae oil and, hence, make it impossible to quantify theproduced hydrocarbons. Moreover, gas oil has significant sulfurcontent, which may affect the hydrodeoxygenation process. Inorder to reduce the complexity and uncertainty in experimentand 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 ofunconverted 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 roomtemperature with low flowability, is dissolved in 100 mL ofdodecane to make 1.3 wt % solution. After 1 h of sonication forbetter mixing, this solution was vacuum-filtered using 410 filterpaper (VWR, 1 μm retention) to remove any residual particlesto avoid blockage of the microreactor.
2.3. Experimental Setup and Procedure. A schematic ofthe setup for the hydrotreating of microalgae oil is shown inScheme 1. The reaction was conducted in a continuous flowSS316L microreactor with an inner diameter of 0.762 mm. Theground NiMo catalyst was packed in the microreactor, and thereactor length was varied from 3 to 19.2 cm depending on theresidence time required for the experiments. Fresh presulfidedNiMo catalyst was used for all runs and was first reduced using
Table 1. Chemical and Physical Characterization ofMicroalgaea Oil
density (298 K) (g/mL) 0.807elemental analysis
C (wt %) 76.29H (wt %) 11.22N (wt %) 0.43O (wt %) 12.06S (ppm) 2033P (ppm) 246ash (wt %) 0.34water content (KF titration) (wt %) 0.86
oil compound in microalgae oilmonoacylglycerides (wt %) 4.41diacylglycerides (wt %) 4.56triacylglycerides (wt %) 29.23free fatty acids (wt %) 9.25total fatty acids (wt %) 47.45
Table 2. Carbon, Oxygen, and Hydrogen ElementalCompositions in Identified and Unidentified Oil
based on wholemicroalgae oil
based on respective elementarycomposition of microalgae oil
fatty acids C 37.50% 49.16%O 5.86% 48.56%H 5.90% 52.57%
unidentified C 38.79% 50.84%O 6.20% 51.44%H 5.32% 47.43%
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pure hydrogen with a reactor temperature of 310 °C andpressure of 500 psig for 2 h before making any reaction run.Two pieces of 1/8 in. outer diameter tubes were packed withglass beads (90 μm) and connected to the entrance and exit ofthe microreactor. Hastelloy micron filter-cloth (200 × 1150meshes, Unique Wire Weaving Co., Hillside, NJ) was placed atthe end of each tube section to retain the packed particles. Thereactor was placed in a temperature-controlled furnace, and thetemperature indicated by the controller was confirmed bymeasuring the temperature inside the furnace using a hand-heldtemperature meter. Pressures at the inlet and outlet of themicroreactor were measured and were adjusted by a backpressure regulator (BPR) installed downstream. The differenceof the two measured pressure values was due to the pressuredrop through the catalyst bed and was used as an indication ofextent of reactor blockage. Compressed hydrogen was regulatedby a mass flow controller (sccm), and an Agilent ADM 2000Universal gas flowmeter was also used to verify gas flow ratebefore and after the reactor. Studies have suggested thatintroduction of gas and liquid flows head to head couldgenerate a short slug length and, thus, improve the mixing intwo-phase flows.28 Therefore, liquid feed was pumped by aHPLC pump at different setting rates (mL/min) and thencontacted with hydrogen at the T-mixer before entering thereactor.A sample loop was used due to low product flow rate. Liquid
product samples were collected for analysis at ambientconditions from the sampling port as shown in the schematicof the setup. The samples were collected at 1 h intervals for off-line analysis, and seven samples were obtained from oneexperimental run. The remaining product stream flowed to thegas−liquid separator, and liquid would be condensed in theseparator, whereas gases, including unconverted hydrogen andgas product would be vented to the hood. After eachexperimental run, nitrogen would be used as purge gas toclean the system, and the BPR was also cleaned by compressedair 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 ofmicroalgae oil hydrodeoxygenation, a result that is consistent
with the observation that the C14, C16, C18, C20 fatty acidsare the main components of natural lipids in algae oil.Therefore, the product analysis focused on the identificationand quantification of C13 to C20 hydrocarbons by a Varian 450Gas Chromatograph (GC) equipped with CP-8400 Autosam-pler. The liquid sample eluted by helium was separated on aZB-1HT nonpolar capillary column (30 m × 0.25 mm × 0.25μm) and detected using a flame ionization detector (FID). Thetemperature program was: initially 50 °C for 1 min, thenramped at 15 °C/min to 240 °C, and finally held for 5 min.Before the analysis of each batch of samples, four GC standardswhich contained C13 to C20 hydrocarbons at four differentlevels were run first for GC external calibration. The responsefactor, which is the ratio of peak area to hydrocarbonconcentration, could be used for quantification. The retentiontime was used for identification of the hydrocarbon. Theidentification was further confirmed by adding standardhydrocarbon 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 yieldtotal mass of C13 C20 hydrocarbon in product
total mass of fatty acids in liquid feed(1)
Space−time yield (STY) of hydrocarbon (rate of hydrocarbonformation) is defined as
=STYmass of hydrocarbon produced/time
mass of catalysthydrocarbon(2)
Even-numbered carbon hydrocarbon to odd-numbered carbonhydrocarbon ratio is defined as
n nHC(2 )/HC(2 1)mass of even numbered carbon hydrocarbon in productmass of odd numbered carbon hydrocarbon in product
Scheme 1. Schematic of Microalgae Oil Hydrodeoxygenation Setup
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Identification and quantification of unconverted fatty acids inliquid product is required for the calculation of lipidsconversion. Polar carboxyl functional groups in fatty acidsneed to be neutralized before being analyzed in GC to avoidadsorption to the column and to obtain better separation. Amodified AOCS method was used for sample derivatization.After sample preparation, all fatty acids in the sample wereconverted to fatty acid methyl ester (FAME). The FAMEsample was also analyzed in GC-FID under the conditionsshown 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 °CDetector: FID, 280 °CCarrier gas: Helium, constant flow = 1.8 mL/minInjection: 1 μL, 1:20 splitIn this FAME analysis method, external standards were used
for peak identification, and a similar method was applied forFAME quantification as for hydrocarbon described above.Because the hydrogenolysis of saturated triglycerides is slowerthan double bond saturation, the conversion of lipids could beapproximately calculated based on measured fatty acids, and allreported conversion data is calculated according to thedefinition
conversionmass 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 hydro-deoxygenation of triglycerides was established using a widerange of model compounds by Peng et al.,3 and a detailedreaction pathway is summarized as shown in Scheme 2. Thefirst step is the saturation of double bonds in the alkyl chain,which occurs very fast; then the second step is the cleavage ofC−O bond in the hydrogenolysis of saturated triglycerides toproduce fatty acids and propane, which is slower than the firststep. The next step is the hydrogenation of the carboxylic group
in fatty acids to produce an oxygenated intermediate, analdehyde which is finally converted either to an even-numberedcarbon hydrocarbon through hydrodehydration (with analcohol intermediate) or an odd number hydrocarbon viahydrodecarboxylation (hydrodecarbonylation). Because theconversion of the fatty acid to the aldehyde is the rate-limitingstep3 effort should always be focused on the fatty acidshydrogenation to improve the overall hydrocarbon yield. Itshould also be noted that the water, CO, and CO2 producedcould also react with H2 in methanation and water−gas shiftreactions. The formation of methane will increase the hydrogenconsumption significantly, which will make the process lesseconomically effective. Therefore, an ideal catalyst should beeffective in suppressing the methanation reaction. According tothis reaction pathway, based on triglycerides, calculations showthat the hydrodehydration route produces the theoreticalmaximum yield, the values of which are approximately 85% and95% for hydrocarbon and carbon respectively, whereas theminimum theoretical values for the yield are approximately 80%for hydrocarbon and 90% for carbon corresponding to thehydrodecarboxylation or hydrodecarbonylation route.
3. RESULTS AND DISCUSSION
3.1. Role of Reactor Inner Diameter (ID). It has beendiscussed in the literature29 that the flow characteristics inmicroreactor (ID < 10−3 m) are different from those inconventional reactors and, therefore, are expected to affect thecatalyst activity. In order to better understand the advantage ofmicroreactor in this three-phase reaction, reactors with ID of0.762 mm, 1.753 mm, and 4.572 mm were compared byevaluating the effect of reactor ID on space-time yield (STY) ofC13 to C20 hydrocarbons, and microalgae oil conversion basedon measured fatty acids. Reaction temperature, pressure, andcatalyst 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 IDreactors with increased volume, the same amount of catalystwas packed into the reactor but the catalyst was premixedthoroughly 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. Thisincrease could be attributed to the enhancement of mass
Scheme 2. Reaction Pathways Involved in Conversion ofTriglycerides into Hydrocarbons.3
Figure 1. Effect of reactor ID on STY of C13 to C20 hydrocarbon andmicroalgae oil conversion based on measured fatty acids (46 mg ofpresulfided 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).
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transfer in Taylor flow regime in microreactor. Because the flowvelocity was kept the same in all three reactors, convective masstransfer enhancement in the microreactor can only beattributed to the existence of Taylor flow, which is not presentin the large diameter reactors. It has been reported that theeffect of shear stress caused by friction is amplified inmicroreactor when compared to a macroreactor, and thisgenerates internal convective recirculation within the liquidslugs (Scheme 3).30 This has been confirmed through flowvisualization experiments, and transitional slug flow wasobserved with liquid slug boundaries broken up by the catalystparticles.31 The recirculation within the liquid slug will decreasethe diffusional path length down to the tube radius, which isgenerally much shorter than the slug half length.29 According toequation of continuity, the diffusion rate is inversely propor-tional to the diffusion pathway length, and hence, the diffusionrate would be enhanced greatly in microreactor than inmacroreactor. Therefore, for multiphase reactions that areusually limited by mass transfer, the microreactor isrecommended both for performance study and kineticsmodeling due to its enhanced mass transfer compared withconventional reactors. In this work, all the other catalystevaluation runs were conducted in the microreactor of 0.762mm ID.3.2. Mass Transfer Analysis. 3.2.1. Analysis of External
Mass Transfer Limitation. According to fluid flow character-ization results in the literature,22 all the experiments weconducted for catalyst evaluation were performed in Taylorflow regime, comprising of alternating liquid slugs (blank) andgas bubbles (blue), as indicated in Scheme 3. Because thereaction occurs only at the active sites on the catalyst surface,the hydrodeoxygenation of microalgae oil reaction on NiMocatalyst involves a three-step external mass transfer as shown inScheme 3: first, the transfer of hydrogen from the bulk gasphase 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 tothe external surface of the catalyst through the solid−liquidboundary layer.Because the gas phase is almost pure hydrogen and the gas−
liquid interface is saturated with this gaseous reactant, the mainmass transfer resistances occur at the gas−liquid and liquid−solid interfaces, with the former being predominant. As thetransfer rates through interfaces depend greatly on flowvelocity, the overall external mass transfer limitation can bestudied experimentally by examining the change of hydro-carbon yield rate at various superficial flow velocities byincreasing gas and liquid flow rate but keeping the gas to liquidratio unchanged, and the residence time constant. The latter isachieved by varying the catalyst loading. The STY of C13 toC20 hydrocarbon at different superficial flow velocities isdepicted in Figure 2. The result shows that the reaction rate isindependent of the flow velocity within the studied range. Thisindicates that the reaction is not limited by external masstransfer even at the lowest superficial flow velocity of 0.117 m/
s. This velocity was chosen as the baseline for all subsequentthe following experimental studies.
3.2.2. Analysis of Internal Mass Transfer Limitation. Asshown in Scheme 3, internal mass transfer refers to thediffusion of reactants from the catalyst pellet pore entrance intothe pore. The reaction is considered as internal mass transferlimited when the reactants concentration inside the pores ismuch lower than that at the entrance. This variation inconcentration throughout the catalyst pellet can be evaluatedby calculating the Weisz-Prater parameter,32 CWP
D Cactual reaction rate
a diffusion rateWPA(obs) p
e As (5)
−rA(obs)′ = 1.11 × 10−4 g of hydrocarbon/g of cat./s is theobserved 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 thecatalyst pellets.CAS = 360 g/m3 is the hydrogen concentration at theexternal surface of the catalyst pellet, which can beobtained from the hydrogen solubility data in dodecane,approximately.De = the effective diffusivity of hydrogen in the catalystpellets, which can be calculated using the equation: De =(DABφPσC)/τ, where DAB = 1.12 × 10−7 m2/s is thebinary diffusivity of H2 in dodecane;33 φP = 0.4−0.6 isthe porosity of the catalyst pellets; σC= 0.7−0.8 is thecatalyst constriction factor; τ = 2−8 is the catalysttortuosity.32 Therefore, De is 3.92 × 10−9 to 2.69 × 10−8
The Weisz−Prater parameter is calculated to be between0.042 and 0.289 depending on effective diffusivity value. By thedefinition, CWP ≪ 1 means reaction rate is much smaller than
Scheme 3. External and Internal Mass Transfer Analysis for Taylor Flow
Figure 2. Effect of superficial flow velocity on space-time yield of C13to C20 hydrocarbon. (Presulfided NiMo, 1.32 wt % Nannochloropsissalina oil in dodecane, temperature = 310 °C, H2 pressure = 500 psig,residence time = 1 s, H2/oil = 1000 SmL/mL).
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the diffusion rate, which indicates that hydrogen concentrationgradient is negligible inside the catalyst pores and the reactionis 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, thecatalyst was evaluated in a nonmass transfer limited range inwhich the measured product yield and reactant conversion is areflection of catalyst activity, with the product distributionreflecting catalyst selectivity. The effect of industrially usedoperating parameters, namely hydrogen to oil ratio, residencetime, 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 hydrotreatingprocesses. The conventional H2/oil ratio refers to the ratio ofhydrogen feed (standard temperature and pressure, STP) to thetotal liquid feed (solvent included).34 In the present study,experiments were conducted to investigate the effect of H2/oilratio between 1000 SmL/mL to 385 SmL/mL by adjustingliquid flow rate. Reaction temperature, pressure, GHSV, andcatalyst loading were kept constant. Because the actualhydrogen flow rate was at least 24 times higher than the liquidflow rate, the effect of liquid flow rate variation on superficialflow velocity was negligible; hence, the residence time remainedthe 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 decreasedrapidly after 1 h reaction and were less than 10% after 3 h.However, there was no observable change in yield for H2/oilratio of 1000 SmL/mL for at least the next 4 h. The same trendis observed in the microalgae oil conversion results (Figure 3b).During the experiment, it was observed that the liquid sluglength decreased with the increase of H2/oil ratio. Because thegas to liquid mass transfer is a strong function of the liquid sluglength,35 the increase of H2/oil ratio will enhance the gas−liquid convective mass transfer, therefore resulting in high yieldof product. In contrast, more intermediates were formed atlower H2/oil ratio because of insufficient amount of hydrogen.As is stated above, because the conversion of fatty acids is therate-limiting step, the formed intermediates were mainly fattyacids, which had low solubility in dodecane and wouldprecipitate on the catalyst surface. Due to the accumulationof oxygenated intermediates on the active sites of the catalyst,the catalyst was, in general, deactivated. However, at elevatedH2/oil ratio, catalyst activity was high and seemed to decreasewith time at a moderate rate. These results show that hydrogenacts not only as reactant but it also serves to protect thehydrotreating catalyst by preventing the adsorption of theoxygenates on the catalyst surface. Catalyst was combusted after7 h on stream using dilute air at 650 °C for 1 h, and gas productfrom combustion was analyzed. Because no CO or CO2 peakwas observed in GC-TCD, it can be concluded that cokeformation was not significant during the reaction.The ratio of even-numbered carbon hydrocarbon to odd-
numbered carbon hydrocarbon (Figure 3c) changes with on-stream time and follows the same trend as the product yieldand reactant conversion, which indicates that the catalystselectivity correlates with catalyst activity. Therefore, to keepthis ratio constant for product quality control, catalystdeactivation 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 thesuperficial volumetric flow rate of the reactants, which is
calculated based on the empty bed assumption. The residencetime was changed by varying reactor length from 3 to 18 cm, aswell as the catalyst loading, while keeping all fluid flow ratesunchanged. Other experimental conditions, such as temper-ature, pressure, superficial flow velocity were held constant. TheH2/oil ratio of 1000 SmL/mL was held fixed in all runs toeliminate the convective mass transfer influence. Therefore,weight hourly space velocity (WHSV) was decreased with theincrease of residence time. The results in Figure 4a−c showthat hydrocarbon yield, microalgae oil conversion, and ratio ofeven-numbered carbon hydrocarbon to odd-numbered carbonhydrocarbon increases with the increase of residence time(between 0.25 to 1.5 s) because the catalyst loading increases.
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 carbonhydrocarbon (46 mg of presulfided NiMo, 1.32 wt % Nannochloropsissalina oil in dodecane, temperature = 310 °C, H2 pressure = 500 psig,residence time = 1s, GHSV = 55 000 h−1).
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Also, all studied performance parameters decline more slowly atlonger residence time, which indicates that per unit mass ofcatalyst, less oxygenated intermediates were produced andaccumulated on catalyst surface. It can be concluded that whensame amount of reactant is fed into the reactor, the catalystactivity and its stability are a function of catalyst loading.Therefore, decreasing the WHSV is a way to maintain theoverall catalyst activity, but productivity will be sacrificed at thesame 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
280−360 °C. Reaction pressure, fluid flow rates and WHSVwere kept the same for all runs with the H2/oil ratio of 1000SmL/mL. Figure 5 shows the performance of the catalyst under
four different reaction temperatures; it is evident that 280 °C istoo low for this hydrodeoxygenation process. When thetemperature increases from 310 to 360 °C the resulting effectis negligible during the first 4 h, but subsequently, the stabilityof the catalyst activity correlates with reaction temperature. At360 °C, there is no sign of reduction in the hydrocarbon yieldfor 7 h. It seems that at elevated temperatures, it may bedifficult for the oxygenates to adsorb on the catalyst surface. As
Figure 4. Effect of residence time on (a) C13 to C20 hydrocarbonyield, (b) microalgae oil conversion based on measured fatty acids, and(c) ratio of even-numbered carbon hydrocarbon to odd-numberedcarbon hydrocarbon (presulfided NiMo catalyst, 1.32 wt %Nannochloropsis salina oil in dodecane, temperature = 310 °C, H2pressure = 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 measuredfatty acids, and (c) ratio of even-numbered carbon hydrocarbon toodd-numbered carbon hydrocarbon (46 mg of presulfided NiMocatalyst, 1.32 wt % Nannochloropsis salina oil in dodecane, H2 pressure= 500 psig, WHSV = 0.65 g of oil/g of cat./h).
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shown in Figure 5a and b, a high reaction temperature can helpto maintain catalyst life and activity. From Figure 5c, it can beobserved that the HC(2n)/HC(2n − 1) ratio increases astemperature increases from 280 to 310 °C. This indicates thatfor the overall hydrodeoxygenation reaction, hydrodehydrationappears to be favored at higher temperatures for presulfidedNiMo/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 temperature-dependent in this range.3.3.4. Effect of Reaction Pressure. The reaction pressure
was controlled by adjusting the back pressure regulator tomaintain a constant total reaction pressure within the reactorsystem. In this study, reactions were run at 300 psig, 400 psigand 500 psig, and each run was of a duration of 7 h. Thereaction temperature, fluid flow rates and WHSV were kept thesame for all runs at H2/oil ratio of 1000 SmL/mL. It should benoted that because the gas flow rates were measured understandard conditions, the actual gas flow rates were pressuredependent; hence, the residence time increased with increase inpressure.As shown in Figure 6a and b, the reaction pressure has
negligible effect on the initial catalyst activity. However, thehydrocarbon yield and microalgae oil conversion decrease inthe following order during the subsequent on-stream hours:500 psig > 300 psig > 400 psig. Similarly, Figure 6c shows thatthe HC(2n)/HC(2n − 1) ratio also decreases following thesame order. This could be explained by the effect of hydrogenpartial pressure. During hydrotreating, Yang et al. found thathigher hydrogen partial pressure was favorable to the DHYDreaction because more adsorbed hydrogen on the surface activesites promoted DHYD reaction but impeded the DCO/DCO2reaction.34 A similar result was also obtained by Guzman et al.36
and Krar et al.37 At 500 psig, hydrogen partial pressure iselevated and residence time is longer than that under the lowerpressure conditions at which enhanced catalyst activity wasmeasured as expected. At 400 psig, the HC(2n)/HC(2n − 1)ratio is the lowest, which indicates relatively higher preferencefor the DCO/DCO2 reaction route than at the other tworeaction pressures. Although pure hydrogen was fed into thereactor, non-negligible amounts of propane, water, carbonmonoxide, carbon dioxide, and methane were produced.38
These gases will decrease the hydrogen partial pressure andinhibit the catalyst activity; more gases will be formed as aresult. At 300 psig, the production of gases is reducedcompared to that at 400 psig and, therefore, will have lesseffect on hydrogen partial pressure. However, at lowerhydrogen partial pressure, residence time is reduced andcatalyst deactivated faster due to the accumulation ofoxygenated intermediates that are produced even in greateramount at the low pressure. In conclusion, hydrogen partialpressure is an important factor for maintaining catalyst life andactivity. Therefore, it is not surprising that in the petroleumindustry, hydrotreating is conducted at 700 psig to prolongcatalyst life. Gaseous products need to be handled properly iftreated gas is to be recycled.38
4. CONCLUSIONSReduced presulfided NiMo/γ-Al2O3 catalyst has been evaluatedfor green diesel production from microalgae (Nannochloropsissalina) oil via hydrodeoxygenation. In the present study,packed-bed reactors of three different inner diameters werecompared to confirm the superior mass transfer characteristics
of the microreactor compared to a macroreactor for this three-phase reaction. Due to the enhanced mass transfer character-istics in microreactor, the STY of hydrocarbon was significantlyhigher in a microreactor (ID < 1 mm). On the basis of theexternal and internal mass transfer limitation analysis, thecatalyst was evaluated by studying changes in the microalgae oilconversion and the product yield and composition as a functionof operating conditions: hydrogen to oil ratio, residence time,reaction temperature, and pressure. The results revealed thatthe conventional hydrotreating catalyst, presulfided NiMo/γ-Al2O3, is deactivated due to the accumulation of producedoxygenated intermediates in hydrodeoxygenation reaction, andits selectivity to even-numbered carbon hydrocarbon producedfrom hydrodehydration correlates with the catalyst activity. The
Figure 6. Effect of reaction pressure on (a) C13 to C20 hydrocarbonyield, (b) microalgae oil conversion based on measured fatty acids, and(c) ratio of even-numbered carbon hydrocarbon to odd-numberedcarbon 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).
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catalyst activity and life can be preserved by elevating hydrogento oil ratio, residence time, reaction temperature, and pressure,which will decrease the adsorption of oxygenates on the catalystsurface. Reaction temperature (in the range of 310 to 360 °C)and pressure between 300 and 500 psig have negligible effecton initial catalyst activity. Hydrogen acts not only as a reactant,but it also serves to protect the hydrotreating catalyst. Weighthourly space velocity (WHSV) is reduced when residence timeis increased by increasing catalyst loading; therefore, the packedbed is exposed to a decreasing amount of reactant per gramcatalyst. It should be noted that the formation of gaseousproduct, such as CO, CO2, CH4, propane, and water, decreasesthe hydrogen partial pressure which leads to poor catalystactivity. The optimum conditions for hydrotreating of micro-algae oil were determined to be 500 psig H2, 360 °C, H2/oilratio of 1000 SmL/mL, and residence time of 1 s. For this set ofconditions the initial catalyst activity was maintained withoutany signs of deactivation for at least 7 h and the obtained C13to C20 hydrocarbon yield was 56.2%, with a carbon yield of62.7%, 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 thehydrotreating of microalgae oil.
■ AUTHOR INFORMATION
Corresponding Author*E-mail: [email protected]. Phone: 201-216-8314. Fax: 201-216-8306.
NotesThe authors declare no competing financial interest.
This research was funded by the DOE-Bioenergy TechnologiesOffice through grant DE-EE0006063. The authors are thankfulfor their support. We would also like to acknowledge thesupport of our principal industrial partner Valicor RenewablesLLC, and we are especially grateful to Dr. Brian Goodall for histechnical contributions to the study. Discussions on productanalysis with Dr. Peter Domaille, formerly of SyntheticGenomics, were very helpful. Special thanks to Dr. JamesManganaro for his inspiring guidance, and the painstakingreview of the experimental results as well as the manuscript. Wewould also like to acknowledge Dr. Moinuddin Malik for hisreview of the manuscript. The authors, of course, bear the soleresponsibility for any errors.
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