Article pubs.acs.org/EF
Cite This: Energy Fuels 2017, 31, 10946-10953
Bleaching and Hydroprocessing of Algal Biomass-Derived Lipids to Produce Renewable Diesel Fuel Jacob S. Kruger,* Earl D. Christensen, Tao Dong, Stefanie Van Wychen, Gina M. Fioroni, Philip T. Pienkos, and Robert L. McCormick National Renewable Energy Laboratory, 15013 Denver West Parkway, Golden, Colorado 80401, United States S Supporting Information *
ABSTRACT: Algal lipids represent a promising feedstock for production of renewable diesel, but there is little information available regarding the integration of pretreatment, extraction, and catalytic upgrading steps. In this work, we examined oil bleaching by two methods and the effects of bleaching on oil deoxygenation over Pd/C and hydroisomerization over Pt/SAPO11 catalysts. The raw oil was completely deoxygenated and 90% denitrogenated after dilution to 25 wt % in hexanes. The bleaching operations (using either a polar adsorbent or concentrated H3PO4) removed 85−90% of the nitrogen and led to 95− 99% nitrogen removal after deoxygenation. Oil processability was also improved by bleaching. The bulk chemistry of the deoxygenation and isomerization was not strongly affected by bleaching, as post-isomerization products with cloud points less than −10 °C and boiling ranges within or close to specification for No. 2 diesel fuel were obtained through 10 h time on stream with or without bleaching.
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INTRODUCTION Lipids are perhaps the most energy dense bioproducts, making them an attractive feedstock for fuel production. Indeed, commercial scale production of hydrocarbon renewable diesel (RD) and jet fuel from animal fat and vegetable oil feedstocks that consist largely of triacylglycerols (TAGs) is being realized using catalytic deoxygenation/hydroisomerization (DO/HI) technology.1−3 Lipids extracted from algal biomass have potential to be a major feedstock source for production of biofuels,4 but process-relevant data on refining crude algal lipids to hydrocarbon fuels is scarce. Techno-economic analyses (TEA) have demonstrated that algal biomass cost is the largest single factor affecting process economics, but the cost of hexane for lipid extraction and hydrogen for deoxygenation are the second and third largest operating costs, respectively.5,6 Thus, there is considerable motivation to gain a greater understanding of the unit operations required for refining algal lipids, which have different chemistry than either conventional crude oils or terrestrial animal or plant lipids. Generally, these operations can be categorized into oil cleanup, deoxygenation (DO), and hydroisomerization (HI). Oil Cleanup. Crude extracted algal oils contain a significant fraction of impurities, including sterols, pigments, polar lipids, and hydrophobic proteins. Because of their inherent high degree of unsaturation and heteroatom content, these impurities have the potential to increase hydrogen consumption and decrease catalyst life. Thus, some form of cleanup will likely be necessary prior to catalytic upgrading. An analogous industrial process, production of food-grade vegetable oils, comprises four common cleanup (or refining) steps:7 1. Deodorization. Steam distillation to remove free fatty acids (FFAs), aldehydes, ketones, and peroxides. 2. Neutralization. Saponification using NaOH to convert FFA to soaps, which are separated by centrifugation. © 2017 American Chemical Society
3. Degumming. Treatment with hot water or warm dilute aqueous acid (∼2 wt % H2O, 0.1−0.2 wt % acid in oil) to remove phospholipids. 4. Bleaching. Treatment with acidic, hydrophilic adsorbent (commonly clays) to remove pigments, reactive components, basic nitrogen compounds, residual phospholipids and soaps, trace metals, and other contaminants. Of the four oil cleanup processes, two (deodorization and neutralization) are directed toward removal of free fatty acids from a primarily triglyceride stream. Because both TAGs and FFAs are fuel precursors, these two processes are not relevant for cleaning oils prior to DO and HI for fuel production. Degumming and bleaching, which target removal of heteroatoms such as P and N, are much more useful in the context of fuel production, as these elements are well-known catalyst poisons. Deoxygenation. The three major reactions that occur in TAG and FFA DO are hydrodeoxygenation, decarbonylation, and decarboxylation, as shown in Table 1 for a hypothetical lipid stream containing only stearic acid. Hydrodeoxygenation is highly exothermic and consumes three moles of H2 per carboxyl Table 1. Oxygen Removal Reactions from Stearic Acid (R = C17H35) reaction type
overall reaction
hydrodeoxygenation
R−COOH + 3H2 → R−CH3 + 2H2O R−COOH → R−H + CO2 R−COOH + H2 → R−H + CO + H2O
decarboxylation decarbonylation
ΔH573K (kJ/mol)8 −115.0 9.2 48.1
Received: June 28, 2017 Revised: August 21, 2017 Published: August 22, 2017 10946
DOI: 10.1021/acs.energyfuels.7b01867 Energy Fuels 2017, 31, 10946−10953
Article
Energy & Fuels
pore size that allows n-paraffins to enter the pore structure, but not isoparaffins. Pt is also included to provide hydrogenation activity and prevent coke formation on the catalyst surface. The most commonly used catalysts are crystalline silica-aluminophosphates (SAPO), zeolite beta, or other zeolite with similar pore size, containing roughly 0.5 wt % Pt. Reaction conditions are typically 350 °C, 3 MPa H2, H2/feed ratio of 26 mol/mol, and WHSV = 1 h−1.27 This approach is commonly used to isomerize alkanes in the range of n-C7 to n-C20 for improvement of fuel properties. These catalysts are very sensitive to basic nitrogen impurities in the feedstock, and thus it is common to include CoMo or NiMo on alumina hydrotreating catalyst bed at the front end of the HI reactor to protect the HI catalyst.28 In the study by Robota noted above, the deoxygenated product was isomerized over Pt-USY zeolite at 258 °C, 5.5 MPa H2, and LHSV of 3 h−1, which achieved 60% conversion of the feedstock n-paraffins. Overall, at least five post-harvest operations govern the economics of renewable diesel production from algal lipids: biomass pretreatment (e.g., dilute acid pretreatment to lyse cells and release lipids), lipid extraction, lipid cleanup, DO, and HI. To our knowledge, there are no reports in the literature that evaluate the interaction of these process steps or the costs and benefits of emphasizing different steps, e.g., lipid cleanup (to allow less severe DO conditions and longer HI catalyst life) over more severe DO conditions (to reduce the costs associated with a lipid cleanup step). Thus, we were motivated to describe a complete process for pretreating and extracting crude algal lipids and converting them to a hydrocarbon product that could be considered a fungible diesel blendstock. Here, we report details on the production of a crude Scenedesmus algae oil, two possible bleaching routes to clean the raw oil prior to hydrotreating, DO of the crude and cleaned oils over a Pd/C catalyst, and HI of the resulting DO products over a Pt/SAPO-11 catalyst.
group, facts that have significant economic impact in both reactor design and in terms of cost of hydrogen. Decarboxylation is nearly thermally neutral and consumes no hydrogen but results in a lower mass and carbon yield. Decarbonylation is slightly endothermic and consumes 1 mol of hydrogen. For more typical feedstocks with unsaturated fatty acid chains, one additional mole of hydrogen will be consumed to hydrogenate each double bond. Deoxygenation of TAG and FFA has been carried out over conventional petroleum hydrotreating catalysts consisting of CoMo or NiMo sulfides supported on alumina. Marker and coworkers9 processed vegetable oil over a commercial NiMo catalyst at 325 °C, 3.5 MPa H2 and a weight hourly space velocity (WHSV) of 0.8 h−1. They obtained 85% deoxygenation with 88% of the products in the diesel range. Subsequent research has obtained similar results2,10−12 but also demonstrated that without a sulfur source in the feedstock, the catalysts gradually lose activity as they are converted from sulfides to oxides.13,14 These sulfide catalysts have also been used for DO of algal oils with some success,15,16 but the loss of sulfur into the product of a low-sulfur feedstock is likely to become increasingly disadvantageous as allowable sulfur concentrations in diesel fuel continue to decrease. As an alternative, a broad range of supported metal catalysts based on Ni or noble metals have been examined for FFA and TAG deoxygenation.8 Ni-based catalysts would be advantageous from a catalyst cost perspective, but results to date have indicated that while these catalysts can show high DO activity initially for both vegetable and algae oils, they are generally prone to deactivation and can increase H2 consumption by converting produced CO and CO2 to CH4.17−22 Among the supported noble metal catalysts employed for DO of FFA and TAG, Pt/Al2O3, Rh/Al2O3, and Pd/C are some of the most studied. While Pt or Rh/Al2O3 have proven promising with dilute algae oils,15 results over Pd/C have been particularly encouraging. This catalyst has been used for DO of model FFAs, model TAGs, and vegetable oils in both batch and continuous reactions. Selectivity to decarbonylation/decarboxylation products is generally high, but activity and yields improve markedly when the reaction is run under H2 instead of inert atmosphere and when the feedstock (as for the Al2O3-supported Pt and Rh catalysts) is diluted in a hydrocarbon solvent such as dodecane or mesitylene.23−25 However, Robota and co-workers were able to deoxygenate relatively refined neat algal TAG using a Pd/C catalyst at 5.6 MPa H2 (350 °C, WHSV = 1.5 h−1, H2/TAG = 30 mol/mol).26 The process was operated for over 190 h without activity loss, obtaining 85% conversion with 95% selectivity to alkane products. These results are promising, but in many cases, the processes used to produce the oil, a critical piece of information for technoeconomic analysis, were not described. Hydroisomerization. Hydroisomerization converts n-alkanes into branched or isoalkanes in order to lower the temperature where wax crystals begin to form in a diesel fuel (i.e., the cloud point). The cloud point is a critical performance property for both diesel and jet fuels. For example, the melting point of noctadecane is about 30 °C and isomerization to introduce one methyl branch reduces the melting point to −5 °C.1 This is also known in the petroleum refining industry as dewaxing and can occur by two mechanisms: HI or cracking (molecular weight reduction). For production of diesel, excessive cracking can lead to products that do not boil in the diesel range, representing a yield loss to the process. The most common catalysts are zeolites and similar crystalline molecular sieves that are selected to have a
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METHODS
Feedstock Oils. Oleic acid (90%, Sigma-Aldrich) and n-hexadecane (99%, Sigma-Aldrich) were used as model feedstocks. Hexanes (reagent grade, BDH) were used as a DO solvent. Algal biomass (Scenedesmus acutus LRB-AP 0401) was produced at Arizona State University in 2013. Cultivation and harvest procedures are described in Laurens et al.29 The algal biomass was mid-growth stage, and thus had a relatively high carbohydrate content. The oil extracted from this biomass is abbreviated below as HCSD (high-carbohydrate Scenedesmus). The algal biomass was pretreated in a batch reactor using 2% (w/w) sulfuric acid for 15 min at 155 °C. Pretreated slurry was cooled, and then centrifuged to separate the liquid fraction containing glucose and mannose from the lipid−protein solids fraction. Pretreated lipid− protein solids (∼20% solid in slurry) and equal volume of hexane were fed into a round-bottom flask (5 L). The mixture was stirred for 3 h using a mechanical stirrer for oil extraction and then allowed to stand at least 2 h for phase separation. The upper hexane phase was transferred to a round-bottom flask for rotary evaporation at 40 °C. Hexane recovered from the condensate was fed back to the biomass mixture for a second extraction cycle. The extraction was carried out four times and hexane extracts were combined. This process was described in detail by Laurens et al.30 Extracted oil was centrifuged at 200 g in a PTFE bottle for 30 min to remove solid residue from clean oil. The solid residue was washed with equal amount of hexane and centrifuged at 200 g for 30 min. The liquid phase was transferred to a round-bottom flask for rotary evaporation, and all the obtained oil was combined. The oil was purged with nitrogen and stored at −20 °C. Algae Oil Silica Bleaching. In one bleaching process, the algae oil was bleached by passing over a silica gel column. The oil was diluted to a concentration of 100 mg/mL in a 50:50 mixture of methanol/petroleum 10947
DOI: 10.1021/acs.energyfuels.7b01867 Energy Fuels 2017, 31, 10946−10953
Article
Energy & Fuels ether (35−60 °C) and passed over a bed of high purity silica gel (60 Å, 60−100 mesh, from Sigma-Aldrich) saturated with petroleum ether. The oil was added to silica gel at a mass ratio of 10:1 and eluted using a mixture of 90:10 petroleum ether/diethyl ether. Elution solvent was added at a ratio of 10:1 to the volume of diluted algae oil to fully rinse the oil from the column. After elution through the column, the solvent was removed by rotary evaporation and bleached oil was stored at −20 °C. The oil bleached by this method is referred to as sb-HCSD (sb = silicableached). Algae Oil Acid Bleaching. As an alternative to silica bleaching, some of the crude algae oil was bleached with concentrated phosphoric acid, by a technique that has been used previously to remove chlorophyll from vegetable oils.31,32 Briefly, crude algae oil was mixed with 3.5 vol % of 85% H3PO4 and heated under vacuum in boiling water bath for 30 min to remove water, and then heated to 120 °C and held for 30 min. Afterward, the oil was separated from the precipitate by decanting. To the decanted oil, equal volume of H2O was added, and the mixture was stirred at 80 °C for 1 h to remove residual polar lipids. The aqueous and oil phases were then separated by centrifuging and decanting, and the oil phase was washed three more times with H2O to remove residual H3PO4. The oil bleached by this method is referred to as ab-HCSD (ab = acid-bleached). Algae Oil Characterization. Algae oils were analyzed by several techniques to provide insight into the composition of the crude oils and the changes in composition effected by silica and acid bleaching. As a preliminary measure, the oils were characterized by solid phase extraction (SPE) on an amino propyl cartridge to determine total neutral, FFA and total polar fractions. Agilent Bond Elut NH2 SPE cartridges (500 mg) were used for oil fractionation. Details of the protocol are given in the Supporting Information. After recording weights of each lipid fraction, the lipids were dissolved in methanol/ chloroform (1:2) and transferred to GC vials for FAME derivatization and GC analysis.33 Oils were also analyzed by gas chromatography with flame ionization detection (GC-FID) to measure mono-, di-, and triacylglycerols and fatty acid content. The methodology employed for this analysis is a modified version of ASTM method D6584 for the analysis of residual glycerides in biodiesel. The method was modified for the measurement of FFAs in conjunction with TAG by addition of standards of stearic acid, oleic acid, and palmitic acid to the calibration curve, and sample size was reduced to ensure component concentrations were within the instrument calibration range. All other method parameters were followed as written. Elemental composition, including C, H, and O was measured by combustion analysis (O calculated by difference). Nitrogen content was quantified using ASTM method D4629. Phosphorus, sulfur, calcium, magnesium, iron, potassium, sodium, and silicon were determined by inductively coupled plasma (ICP). Sterols and phytol were analyzed as TMS derivatives by GC-FID after an alkaline saponification using a one-third-scale version of the AOCS Ca 6b-53 method for unsaponifiable matter. Details of the analysis are given in the Supporting Information. Pigments were measured qualitatively by extracting the oils with methanol, followed by UV−vis spectroscopy. The “Other Neutrals” fraction was calculated as the difference between the “Total Neutral” determined by SPE and the sum of the glyceride fraction determined by the modified ASTM D6584 described above, the sterols, and the phytol quantified in the unsaponifiable matter. The “Unknown” fraction was calculated as the mass balance. Reactor System. Reactions were run in a continuous flow fixed bed tubular reactor, 14 in. (length) × 0.375 in. (i.d.). The reactor tube contained the catalyst bed in the center, with quartz chips upstream and downstream to hold the catalyst bed in place and facilitate heat transfer to the feed. The reactor tube was positioned inside a furnace, and temperature was measured by a thermowell in the axial and radial center of the reactor tube. For both the DO and HI steps, catalyst particles were sieved to 60−120 mesh and diluted to 20 wt % in 60−80 mesh SiC. Liquids were fed from an Eldex Optos HPLC pump and mixed with a gas flow of 95% H2/5% Ar before entering the reactor. A tube-in-shell heat exchanger downstream of the reactor was used to condense liquid
products, which were collected in a knockout vessel and recovered for analysis. Uncondensed products flowed to an online GC for analysis. The setup is shown schematically in Figure S1. Catalysts. The DO catalyst was 5 wt % Pd/C, which was used asreceived from Johnson Matthey, except for crushing and sieving to obtain particles in the 125−250 μm (60−120 mesh) range. The catalyst was reduced for at least 1 h in flowing hydrogen at pressures over 0.70 MPa and temperatures over 100 °C in the process of bringing the reactor up to operating conditions. The HI catalyst was a 1 wt % Pt/SAPO-11 catalyst. The SAPO-11 base material was synthesized as previously described.34,35 A detailed description of the synthesis is given in the Supporting Information. Pt was loaded onto the sieved particles as an aqueous solution of H2PtCl6· 6H2O by incipient wetness impregnation. The catalyst was dried overnight at 40 °C in a vacuum oven prior to loading into the reactor. The catalyst was reduced in situ at 350 °C and 3.45 MPa H2 at 250 sccm for 2 h before the reaction was started. Catalysts were characterized by N2 physisorption, NH3 temperature programmed desorption (TPD), and transmission electron microscopy (TEM), as described in the Supporting Information. The catalysts had surface area, acid site content, and metal particle diameter as shown in Table 2. The Pd/C catalyst showed roughly 40% of the Pd metal
Table 2. Characterization of Fresh DO and HI Catalysts Pd/C Pt/SAPO-11
surface area (m2/g)
μmol NH3/g
metal dp (nm)
1235.6 187.9
439.1 1080.1
23.5 ± 25.0 1.7 ± 0.5
particles at dp < 10 nm, 50% at 10 nm < dp < 50 nm, and 10% at dp > 50, with a range extending up to nearly 200 nm, resulting in the large relative standard deviation in metal dp for that catalyst. Reaction Conditions. Reaction conditions were selected based on preliminary work with model lipids (i.e., soybean oil and oleic acid), model impurities (i.e., indole), and the available quantities of algae oils. For the DO step, the catalyst bed volume was set to 6 cm3, the liquid flow rate to 0.1 mL/min (LHSV = 1 h−1), and the H2 flow rate to 92 sccm (H2/feed ratio = 1000 N m3/m3). Temperature, pressure, and oil dilution with hexane were 450 °C, 8.96 MPa, and 25 wt % oil in hexanes, respectively, unless otherwise specified. For the HI step, the bed volume was 5.2 cm3, the liquid flow rate was set to 0.043 mL/min (LHSV = 0.5 h−1) and the H2 flow rate to 92 sccm (H2/feed ratio to 2325 N m3/m3). Temperature and pressure were 350 °C and 3.45 MPa H2, respectively. DO product was fractionated into an organic phase, containing hexane and deoxygenated oil, an aqueous phase, containing product water, and a waxy solid phase that suspended between the aqueous and organic phases and may have been comprised of wax esters or ketones.36 The raw product was centrifuged at 3500 rpm for 5 min in a Thermo Electron Corporation IEC CentraCL2 centrifuge, and the hexane/ deoxygenated oil fraction was separated by decanting. The hexane was removed by rotary evaporation, and the deoxygenated oil was reacted over the HI catalyst undiluted. Product Characterization. DO and HI liquid-phase products were characterized by gas chromatography with mass spectrometry (GC/ MS) to identify components and qualitatively assess contents of component classes. The GC configuration and analysis program is described in the Supporting Information. Cloud points of HI products were determined by ASTM method D5773. Distillation temperatures and n-alkane compositions were measured by GC-FID simulated distillation following ASTM method D2887. These methods are accepted by ASTM for determining cloud point and distillation, respectively, in diesel fuels.37 Nitrogen contents of DO and HI products were measured by ASTM method D4629. Residual acid in oleic acid DO products was measured following ASTM method D664 and quantified as oleic acid. DO and HI gas-phase products were characterized by online GC equipped with TCD and FID detectors. The GC configuration and analysis program is described in the Supporting Information. 10948
DOI: 10.1021/acs.energyfuels.7b01867 Energy Fuels 2017, 31, 10946−10953
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Energy & Fuels
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RESULTS AND DISCUSSION Feedstock Characterization and Bleaching Effects. Compositions of the high-carb Scenedesmus (HCSD), SiO2bleached HCSD (sb-HCSD), and acid-bleached HCSD (abHCSD) oils are provided in Figure 1 and Table 3. The oils were
Mass yields for the silica and acid bleaching were 89% and 86%, respectively, indicating a minor loss of fuel precursors in addition to the impurity removal. Although the silica bleaching resulted in higher mass yields in the current experiments, the acid bleaching using H3PO4 may be more scalable, as the solid−liquid separation required for either silica or clay bleaching is avoided. While bleaching with H3PO4 is effective, bleaching with concentrated H2SO4 may be even more economical, as it is reported to remove chlorophyll even at room temperature.31 Further insight into the phytol trends can be gleaned from the UV−vis spectra of the oils, which are shown in Figure 2. There
Figure 1. Composition of HCSD algae oil used in DO experiments. XAG = triacylglycerides + diacylglycerides + monoacylglycerides. Other neutral = total SPE neutral fraction minus (XAG + sterols + phytol). FFA = free fatty acids. Unk. = unknown.
Table 3. Elemental Composition of Crude and Treated HCSD Oils element
crude HCSD
sb-HCSD
ab-HCSD
carbon, wt % hydrogen, wt % oxygen, wt % nitrogen, ppm sulfur, ppm phosphorus, ppm calcium, ppm iron, ppm magnesium, ppm potassium, ppm sodium, ppm silicon, ppm
77.2 12.0 10.8 1644 167 18 17 6 5 26 17 36
75.3 11.8 12.7 254 69 4 372 1 50
