Article pubs.acs.org/EF
Process Balances of Vegetable Oil Hydrogenation and Coprocessing Investigations with Middle-Distillates Matthias Endisch,* Thomas Kuchling,* and Jan Roscher* TU Bergakademie Freiberg, Institute for Energy Process Engineering and Chemical Engineering, Fuchsmühlenweg 9/Haus 1, 09599 Freiberg, Germany ABSTRACT: The hydrogenation of vegetable oil is a promising technology for the production of highly valuable diesel components. Suitable production routes are the hydrogenation on sulfided NiMo catalysts with pure vegetable oil as feedstock or their coprocessing with middle distillates. To obtain the yields of all products for a complete process balance, experimental studies with rapeseed oil and jatropha oil were performed in a pilot scale test bench. IR spectroscopy and 13C NMR spectroscopy as well as gas chromatography were used for analyzing the liquid product. On the basis of a simplified reaction scheme, a procedure for the recalculation of specific yields and hydrogen consumption is presented and compared with the experimental values. The coprocessing experiments include the admixture of 5−20 wt % jatropha oil to a straight run middle distillate in the temperature range of 300−350 °C. It could be shown that the optimum reaction temperature for production of alkanes with nearly complete oxygen removal is significantly lower than that for the desulfurization of middle distillates down to 10 ppm sulfur. Concerning the cold flow properties, an admixture up to 15 wt % jatropha oil does not increase the Cold Filtration Plugging Point (CFPP) of the base middle distillate.
1. INTRODUCTION Finding a sustainable energy supplement as well as the need for carbon dioxide reduction leads to the necessity to integrate more and more renewable energy sources into the transportation fuel markets. In many countries (e.g. Germany) this is forced by blending quotas. Currently, the required quota for diesel in Germany is achieved by using FAME (fatty acid methyl ester), especially RME (rape-oil methyl ester). It is produced by transesterification of vegetable oils with methanol. With carbonnumbers of 11−20, these esters are in the diesel range and can be used as blending components. A typical structure is shown in Scheme 1.
diesel components from renewable sources, equal to fossil diesel components, are required for higher quotas without an adaptation of the existing motor systems. Synthetic diesel fuel of the second generation via biomass gasification, the so-called BtL-diesel (Biomass to Liquid) is not commercially available in valuable amounts in the next years. This results in a need for a bridging technology based on vegetable oil, which is easy to handle but limited in raw material supply. An alternative approach should use resources that are available in sufficient amounts and approved technologies for realization. One such concept is the catalytic conversion from vegetable oils in typical refinery processes (cracking, hydrogenation) with the aim to produce components, which are principally not differing in their chemical configuration, from fuels made from petroleum. Considering the triglyceride configuration of vegetable oils (Scheme 2), the process engineering aim is the cleavage of the ester bonds, the removal of the oxygen and a decrease of the
Scheme 1. Typical Fatty-Acid-Methyl-Ester
Scheme 2. Triglyceride Molecule and Reaction Scheme of the Hydrogenation Reaction
In contrast to typical petroleum derived diesel components like alkanes (paraffin’s), these chains contain up to three doublebonds as well as oxygen-containing ester groups which cause some disadvantages in use as transportation fuel: • Lower energy content per mass (85−90% of conventional Diesel) • Minor chemical stability (storage stability, coke formation, deposit) • Affinity to water (microorganism) • High boiling range (dilution of motor oil) • Coke formation in injector nozzles Blending rates more than 7 vol% of FAME can cause technical problems in cars due to the chemical properties of methyl esters and are therefore out of European fuel specification range. The European fuel directive claims an increasing energy-based content of biofuels in the coming years.1 Thus, oxygen free © 2013 American Chemical Society
Received: January 3, 2013 Revised: March 13, 2013 Published: March 19, 2013 2628
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
Energy & Fuels
Article
catalysts.16,17 More detailed information about the influence of the catalyst design and process conditions can be found in ref 18. A possibility for introduction of hydrogenated vegetable oils on the market is the coprocessing in conventional hydrotreater and hydrocracker units in a refinery. First results were published by Balfanz and Baldauf.19 It could be shown that hydrotreating with middle-distillate is to be preferred in opposite to hydrocracking with VGO (vacuum gas oil) because of the lower cracking-gas production from vegetable oil. Furthermore, the higher temperatures (>360 °C) in mild hydrocracking units leads to the formation of higher molecular by-products, which lowers the overall conversion.20 Regarding the influence on the sulfur removal there are only few results available, which can be applied hardly on the present situation of nearly sulfur free diesel19 and the coprocessing with middle distillates.20,21 Concerning the product properties, hydrogenated vegetable oil is a high cetane diesel component (CN ≈ 100). It is free of aromatics and sulfur. Motoric tests showed that adding hydrogenated vegetable oil to conventional diesel can lower the CO and hydrocarbon emissions and reduces also the particle emissions.19 Unfortunately, the blending rates are limited by the poor cold flow properties of alkanes which are the main products (cold filtration plugging point ≈ 18 °C). Also the effect of flow improvers in mixtures with conventional diesel is much lower compared to pure conventional diesel. For this reason, a blending of maximum 30 wt % (summer diesel, CFPP 0 °C) or 10 wt % (intermediate diesel, CFPP −10 °C) is possible with typical hydrogenated vegetable oil.22,23 For winter diesel quality the maximum blending rate was indicated with 1.5 wt %.19 Improved cold flow properties could be reached with higher amounts of isoalkanes,24 which would be possible by isomerization of the nalkanes on a suitable bifunctional catalyst with a high selectivity.25 Hydrocracking at temperatures up to 420 °C and 180 bar results also in improved cold flow properties of the vegetable oil products but leads also to the formation of liquid products with boiling temperatures above diesel range and cracking products, which are not yet quantified.23 Although many experimental works were done, complete yields derived from balancing the process are very rare. Therefore, overall yields of the hydrogenation process for different vegetable oils based on experiments in a bench scale trickle bed reactor are presented in this work and compared with the theoretical amounts of an idealized reaction scheme. Furthermore, the coprocessing with a middle distillate fraction and the influence of the amount of added vegetable oil on product properties and sulfur removal are also in focus of those investigations. The products were analyzed by gas-chromatography as well with IR-spectroscopy and 13C NMR.
molecular size. Long chained n-alkanes with boiling points ranges of diesel fuels are to be expected as main products. In catalytic cracking processes, oxygen is separating in the form of carbon oxides (decarboxylation). Aliphatic decomposition products occur, which are principally suitable as diesel components. Laboratory investigations with mostly acidic catalysts revealed numerous side reactions, which direct the conversion to undesirable components such as aromatic compounds and unsaturated and light decomposition products, thus leading to a reduction of the selectivity of the overall process.2−4 Under given conditions as hydrogen atmosphere and in the presence of a hydrogenation active catalyst, decarboxylation, oxygen removal accompanied by production of water and saturation of double bonds and arising free valences are observed. That should result in considerably improved diesel yields and qualities. The positive influence of high hydrogen partial pressures on the yield of long chained alkanes was shown in autoclave tests.5 Further studies dealt with the mechanism of the process,6 which is quite complex due to the simultaneous appearance of hydrogenation and catalytic, respectively, thermal decarboxylation and which is discussed, to some extent controversially, in the literature. Studies for the creation of cetane number improving additives on vegetable oil hydrogenation basis were carried out in a continuous pilot plant.7,8 Furthermore, a process with hydrogenation and a subsequent alkane isomerization for the improvement of the low temperature resistance is patent-registered.9 A first industrial installation (170.000 t/a) has been put on stream for the Neste Oil Corp. in 2007. Results of studies in hydrotreating vegetable oils with complete hydrogenation of three different vegetable oils were published in refs 10 and 11. The product depended strongly on the fatty acid distribution of the plant oils, thus, the final chain distribution of the alkanes could be easily predicted, if the hydrogenation/decarboxylation-ratio is known. The typical reaction temperature for complete conversion with sulfide catalysts is about 310−340 °C, depending on catalyst, space velocity, and pressure.10,12 The influence of active metal sulfide (Ni, Mo, NiMo-mixtures) was investigated by Kubicka et al.13 It could be shown that combining Ni and Mo sulfides results in the highest activity. The hydrogenation/decarboxylation-ratio depended on metal composition, where Ni favors the decarboxylation and Mo the hydrogenation pathway. Also sulfided NiW showed a comparable activity. With a more acid support resulting in an increase of isomerization reactions, the yields of iso-paraffins and therewith the product quality could be improved.14 Long-term stability for different vegetable oil qualities and the effect of adding a hydrogen sulfide source were studied in ref 15 with a CoMo-catalyst. Phosphor-containing components, like phospholipids and alkalis, both have higher amounts in lower grade vegetable oils, leads to a larger decrease of the activity with time on stream. Because of the sulfur-free feedstock, the presulfided catalyst was leached out. This was prevented by adding a sulfur component. It was possible to shift the reaction to higher decarboxylation ratios by higher sulfur concentration in the feed. However, the sulfur in the process necessitates a sulfur separation from the gaseous products. Thus, an upcoming challenge is the development of a sulfur free catalyst, suitable for stand-alone plants. Such plants could be operated without a sulfur removal unit. Promising results were presented in experimental studies based on Pd/SAPO-31 and Ni−Cu Al2O3
2. EXPERIMENTAL SECTION 2.1. Pilot Plant Configuration and Catalyst. Experiments were performed in a bench scale continuous trickle-bed reactor at 60 bar. Scheme 3 illustrates the plant configuration. The reactor has a maximum liquid flow of approximately 250 mL/h. The total volume of the 1.2 m long tube reactor is 325 mL. The catalyst bed, a mixture of silicon carbide (d = 100−200 μm) and catalyst (volume proportion 1:1 referred to the bulk density), was placed between a flowin and a flow-off layer. The dilution was done to prevent wall effects, provide plug-flow and ensure isothermal conditions. The catalyst bulk volume was about 90 mL. A typical catalyst for gasoil hydrogenation (sulfided NiMo on alumina) is used for all experiments. Therefore, for pure vegetable oil treating, a sulfur additive has to be added to the feed to prevent leaching and to guarantee the catalyst stability. Additin S400 from the Rhein Chemie Corporation was used for this purpose. The 2629
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
Energy & Fuels
Article
Scheme 3. Flow Sheet of the Pilot Plant
Table 2. Petroleum Derived Feedstock straight run gas-oil sulfur (ppm) aromatics (wt %) T5 (°C) T65 (°C) T95 (°C)
H2, CO, CO2, and H2S. The water in the liquid fraction is separated in a settler and weighted. The organic liquid fraction is analyzed by several methods (Table 3). 2.4. Process Balance. The whole pilot plant is balanced when it was stabilized after 4−5 h on stream. Before each stabilization time, the pump is calibrated to adjust the correct feed stream. During the sampling time (usually 2−10 h) after stabilization, the gas volume which passed the outlet is summed up by a gas volume measuring device and corrected for standard temperature and pressure. At the end of sampling, the liquid product in the low pressure separator is weighted. Due to the small amount of sulfur additive (10 ppm)
hydrogenation shows only bands of the CH2 and CH3 groups (Figure 2). The results of the IR and the 13C-MNR analysis revealed the complete hydrogenation of the triglycerides of the vegetable oils under the chosen reaction conditions. The oxygen of the ester groups has been removed completely and the products contain no double bonds. The conversion of organic bonded oxygen in water and carbon oxides in dependence of the temperature is investigated for jatropha oil by IR analyses and elemental analysis of the organic liquid product (Figure 3). It has to be taken into account the low
signals appears. These signals can be associated to different structural groups. Figure 1 shows the recorded NMR spectrum of
Figure 1. 13C NMR spectrum of rapeseed oil and hydrogenated rapeseed oil.
the utilized rapeseed oil and the hydrogenation product. Apart from the signals of the carbon atoms of the fatty acids, the signals of the glyceridic bounded C atoms are being registered. The spectrum of the product shows signals of terminal methylcarbon atoms and C atoms of the CH2 groups. The triglycerides and partial glycerides have been transformed and the carbon double bonds have been saturated with hydrogen. 3.1.2. IR Spectroscopy. In Figure 2, the IR spectra of the feedstock material and the hydrogenation product of the rapeseed oil hydrogenation are compared. The significant differences between the two spectra are the absorption bands of the oxygen compounds, which can be measured in the charge material only. The IR spectrum of the product of the rapeseed oil
Figure 3. Oxygen removal detected by IR and elemental analysis (trend curve based on experimental data).
sensitivity of an IR-analysis to small concentrations. With regard to this fact, a complete oxygen removal is achieved at 340 °C (LHSV 2 h−1). 3.1.3. Mass Yields and Hydrogen Consumption. The process is supposed to be run with complete oxygen removal, thus, the yields for the different oil feedstocks where calculated according to 3.4 based on experimental runs at a temperature of 340 °C and LHSV of 2 h−1, where the oxygen is almost out of the liquid organic fraction (Figure 4). All liquid organic products are counted as diesel components. The errors in the balances for these two experimental runs were 6% for rapeseed oil and only 1.4% for jatropha oil. There is only a small difference between the product distributions of the two feedstocks especially for diesel components, due to the very similar fatty acid composition. Oxygen is removed as carbon oxides or water. Other hydrocarbon gases, mainly methane, ethane, and butane, are products from side reactions like hydrogenolysis, methane formation from CO, and hydrocracking. Overall, the diesel product loss in side reactions is very low. The hydrogen consumption can also be calculated as a relative consumption related to the feedstock (eq 6) and gives 294 L/kg (rapeseed oil) and 280 L/kg (jatropha oil).
Figure 2. IR spectra of the rapeseed oil hydrogenation. 2631
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
Energy & Fuels
Article
As shown in Figure 6, the reaction pathway is influenced by temperature where the decarboxylation is favored by higher temperatures resulting in more uneven numbered chains.
Figure 4. Yields for complete oxygen removal for rapeseed oil (left) and jatropha oil (right) at a temperature of 340 °C and LHSV of 2 h−1.
The hydrogen consumption is affected by the oxygen removal pathway (decarboxylation or hydrogenation) where the water formation consumes more hydrogen. The temperature dependence is shown in Figure 5. Figure 6. HC chain distribution for hydrogenated jatropha oil.
3.2. Calculation of Theoretical Yields and Comparison to Experimental Results. On the basis of the measured yields an idealized global reaction scheme is proposed (Scheme 4). All Scheme 4. Global Reaction Pathways for Complete Deoxygenation
Figure 5. Hydrogen consumption, water yield, and COx yield for jatropha oil hydrogenation (trend curves based on experimental data).
The hydrogen consumption is lower for incomplete oxygen removal at a temperature of 300 °C. It reaches a maximum and decreases with increasing temperature. Similar behavior is observed for water yield. The shift from a hydrogenation to a decarboxylation pathway can also be detected for the hydrogen consumption. 3.1.4. Hydrocarbon Chain Distribution. Furthermore, the molar ratio of the decarboxylation to hydrogenation (dec/hyd) can be calculated by calculating the molar hydrocarbon chain distribution, where hydrogenation is indicated by even numbered and decarboxylation by uneven numbered HC-chains. Only the main hydrocarbon chains in the product have been taken into account. x + xC17 dec/hyd = C15 xC16 + xC18 (7)
side reactions like formation of HC-gases except for propane are neglected. The CO formation reaction can be either decarboxylation or reverse water gas shift reaction. It was not possible to distinguish both with the help of experimental data. However, for both ways every CO molecule is coupled to the formation of one water molecule (for complete deoxygenating). In the scheme, the reversed water−gas-shift reaction is only shown. According to these idealized assumptions, theoretical yields can be calculated from the fatty acid composition of the feedstock when the dec/hyd ratio is known. The molar fraction of decarboxylated fatty acid chains (dec) and hydrogenated fatty acid chains (hyd) can be calculated.
The data in Table 4 were derived from GC-FID analysis of the organic liquid product. Table 4. Ratio dec/hyd of Pure Vegetable Oil Hydrogenation According Eq 7 jatropha oil temperature (°C) dec/hyd
300 1.21
320 1.60
340 1.97
rapeseed oil 360 2.10
340 1.73 2632
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
Energy & Fuels
dec =
Article
1 1 dec / hyd
x Propane =
+1
(8)
hyd = 1 − dec
(9)
YPropane =
All molar yields xi are calculated per mole fatty acid chain in the feed. The number of fatty acid chains per glycerine moiety FA/Gly is calculated considering the molar fraction of free fatty acids per total number of fatty acid chains in the feedstock. 1 ·3 (1 − x FFA )
FA/Gly =
(20)
MC3H8 FA/Gly·MFAChain
(21)
For complete deoxygenation of the liquid organic product the ratio of the volume fractions ν (equal to the molar fractions in case of ideal gas) of carbon oxides to propane can be calculated and compared to experimental data, even when a complete balancing is not possible. vCOx xCO2 + xCO dec = = vPropane x Propane FA/Gly (22)
(10)
The calculation of the molar mass M of the average fatty acid chain and the corresponding part of the glyceride moiety is based on the fatty acid configuration and the amount of free fatty acids FFA. MFAChain =
1 FA/Gly
Relative Hydrogen Consumption VΔH2,rel. xΔH2 = n +
∑ xCN,FA ·(MCN,FA − 1·MH)
3 1 1 7 dec · + dec + ·hyd + 2 CO2 /CO 2 FA/Gly 2 (23)
CN
+
YΔH2 =
3·MC + 5·MH FA/Gly
(n +
(11)
N-alkane Yields. The product yields Y for complete deoxygenation and hydrogenation can be calculated as follows. xn ‐ alkane = 1
mol mol(FAChain)
VΔH2,rel =
+ dec ·( −MCO2 + MH) − hyd( −2·MO + 3·MH) (13)
YCO2 =
(
(14)
)
MFAChain
xCO = dec ·
YCO =
1 CO2 / CO
dec ·MCO· CO
2
(16)
1 / CO
MFAChain
(17)
Water Yields. x water = (1 − dec) +
YCO =
dec CO2 /CO
(1 − dec) ·M water +
(18)
dec ·M water CO2 / CO
MFAChain
H2
YΔH2 ρH
(25)
2
vegetable oil
temperature (°C)
dec/hyd
CO2/CO
rapeseed oil jatropha oil
340 320 340 360
1.73 1.60 1.97 2.10
2.12 1.91 2.18 2.16
Figure 7 shows that the calculated yields for carbon oxides, water and organic liquid products are in good agreement with the measured yields. The measured relative hydrogen consumption is slightly lower than the calculated consumption. The difference in the calculated and measured hydrogen consumption is probably reasonable. The measurement of hydrogen consumption is more error susceptible than other yields since hydrogen was used in excess and the minor number of cracking reactions are not taken into account. Since these data are based solely on analyzing the gas and the liquid products and not referred to any balancing data, the theoretical calculation of the yields is practicable for experiments, where complete balancing is not possible, e.g., in very small reactors. 3.3. Co-Processing. The advantage of feeding a mixture of low percentage vegetable oil and conventional middle distillates is the opportunity to use the same equipment at the same time saving costs. The experiments were done with a vegetable oil content up to 20 wt %. The middle distillate is characterized in
(15)
1 CO2 /CO
)·M
Table 5. Analysis Data Used for Theoretical Calculations (Ratio dec/hyd According Eq 7)
Carbon Oxide Yields. With the molar ratio of CO2/CO by gas analysis or water gas shift equilibrium.
dec ·MCO2 · 1 −
dec CO2 / CO
For comparison with the experimental data, the following parameters for theoretical calculations were taken from analysis of liquid (dec/hyd) and gasous product (CO2/CO) and listed in Table 5.
CN
xCO2
7
(24)
Yn ‐ alkane = ( ∑ xCN,FA ·(MCN,FA − 1·MH)
⎞ ⎛ 1 = dec ·⎜1 − ⎟ CO2 /CO ⎠ ⎝
1
+ 2 dec + 2 ·hyd + MFAChain
(12)
+ 2·MH ·n)/MFAChain
3 1 · 2 FA / Gly
(19)
Propane Yields. 2633
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
Energy & Fuels
Article
Figure 7. Calculated (□ grey) and measured (■ black) yields as well as hydrogen consumption.
Figure 8. Sulfur removal of middle distillate with different percentages of vegetable oil (trend curves based on experimental data).
Table 2. The focus was set on the desulfurization efficiency at different temperatures, which is the main target for hydrotreating petroleum derivatives. The limit for diesel fuel is 10 ppm in the EU. As shown in Figure 8, for admixing vegetable oil, which is in principle sulfur free, dilution effects lead to a lower sulfur content in the products. The sulfur removal from the middle distillate was not inhibited by the admixed vegetable oil in the chosen temperature range. Only for the highest temperature, the sulfur content with 5 wt % vegetable oil was slightly higher than without vegetable oil. However, the 10 ppm limit could not be reached with the chosen catalyst up to 350 °C and the LHSV of 2 h−1. Therefore, the effect of the vegetable oil in ultralow sulfur diesel
production could not be investigated. Higher temperatures can cause the formation of higher boiling products from vegetable oil out of diesel range.23 Sulfur removal down to 10 ppm can probably be reached with lower LHSV or a more active catalyst. Despite this, the optimum temperature for production of sulfurfree diesel (10 ppm) and complete oxygen removal from vegetable oils differs by at least 20 K. This can be approximated by extrapolating the trend in sulfur content in Figure 8(left). Figure 9 illustrates that the yield of hydrocarbon gases from cracking reactions increases with increasing temperature. Thus, coprocessing leads to more cracking products from vegetable oil than from hydrotreating single feedstock. 2634
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
Energy & Fuels
Article
Figure 9. Hydrocarbon gases for different reaction temperatures for jatropha oil (trend curves based on experimental data).
Figure 11. Boiling behavior of hydrogenated vegetable oil and mixtures with conventional diesel.
3.4. Cold Flow Properties. Since the cold flow properties of hydrogenated vegetable oil are very poor, it is an important issue for determining the possible admixing rate to conventional middle distillates. The CFPP was measured for the products of the coprocessing experimental runs as well as mixtures of separately hydrogenated vegetable oils and straight run middle distillates. The amount of admixed hydrogenated vegetable oil corresponded to the diesel yield of the fed amount of vegetable oil in coprocessing. For example, a mixing rate of 10 wt % vegetable oil means an addition of 8.4 wt % of hydrogenated vegetable oil to the hydrogenated straight run diesel. Surprisingly, the CFPP for a mixture with less than 15 wt % is lower than the value for pure straight run diesel, represented by the value for 0 wt % jatropha oil (Figure 10). Also, the CFPP for
The boiling behavior of pure hydrogenated jatropha oil differs significantly from a typical diesel behavior and do not achieve the fuel specification for diesel fuel. The reason is the very narrow chain length distribution. However, mixtures up to 20 wt % does not change the boiling curve significantly compared to the investigated straight run diesel.
4. CONCLUSIONS The hydrogenation of jatropha oil and rapeseed oil was investigated in a continuous trickle bed reactor at 60 bar with a typical sulfided hydrotreating catalyst. For LHSV 2 h−1 a complete conversion is achieved at 340 °C. The reactor was balanced completely and the overall yields of the process could be determined. The two different reaction pathways are influenced by temperature, which could be shown by the chain length distribution, the hydrogen consumption, and the yields of the byproducts water and COx. Also, a set of equations is proposed for the theoretical calculation of the overall yields and the hydrogen consumption based on the fatty acid configuration of the feedstock, the chain length distribution of the liquid product, and the molar ratio of CO2/CO. This is applicable, for example, for calculating the yields for complete conversion in small scale laboratory reactors where a balance is not possible or very error susceptible. The yields for the experimental runs are in good agreement with the calculated values. In coprocessing experiments of jatropha oil with straight-run middle distillate, no negative influence on sulfur removal was observed down to 50 ppm sulfur (at 350 °C) in the product. Higher sulfur removal is not examined due to the thermal stability limit of vegetable oil to prevent formation of higher molecular products. However, the gap for the optimum reaction temperature for the vegetable oil hydrogenation and the production of ultralow sulfur diesel (ULSD) is at least 20 K for the catalyst used in the experiments. This means that the production of hydrogenated vegetable oil via coprocessing results in higher diesel component losses by hydrocracking and methane hydrogenolysis as compared to the pure vegetable oil as feedstock.
Figure 10. CFPP for different amounts of hydrogenated vegetable oil and middle distillate (trend curves based on experimental data).
the samples generated after separate hydrotreating are slightly better than the coprocessing products. An explanation for these phenomena could not be found. It is expected that it depends strongly on the properties of the petroleum derived middle distillate. However, no negative effect of hydrogenated vegetable oil admixing is observed up to 15 wt %. For a more precise evaluation, the effect of flow improvers should also be taken into account. 3.5. Boiling Behavior. The boiling behavior is an important issue for diesel fuels achieving the fuel specification. Boiling curves for pure hydrogenated vegetable oil and hydrogenated straight run diesel as well as two mixtures of both are received by Engler distillation (DIN 51751) (Figure 11).
■
AUTHOR INFORMATION
Corresponding Author
*E-mail:
[email protected] (M.E.); thomas.
[email protected] (T.K.);
[email protected]. de (J.R.). Notes
The authors declare no competing financial interest. 2635
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
Energy & Fuels
■
Article
ABBREVIATIONS LHSV, Liquid hourly space velocity; WABT, Weighted average bed temperature; RFA, X-ray fluorescence adsorption; IR, Infrared; FA, Fatty acid; CFPP, Cold filtration plugging point; CN, Cetane number
■
REFERENCES
(1) Richtlinie 2009/28/EG des Europäischen Parlaments and des Rates zur Förderung der Nutzung von Energie aus erneuerbaren Quellen. (2) Idem, R. O.; Katikaneni, S. P. R.; Bakshi, N. N. Fuel Process. Technol. 1997, 51, 101−125. (3) Churasiri, W.; Yongchareon, W.; Viditsant., T. Kor. J. Chem. Eng. 2006, 23 (3), 349−355. (4) Churasiri, W.; Viditsant., T. Energy Fuels 2005, 19, 1783−1789. (5) Gusmao, J.; Brodzki, D.; Djega-Mariadassou, G.; Fretty., R. Catal. Today 1989, 5, 533−544. (6) da Rocha Filho, G. N.; Brodzki, D.; Djega-Mariadassou., G. Fuel 1993, 72 (4), 543−549. (7) Stumborg, M.; Wong, A.; Hogan, E. Bioresour. Technol. 1996, 56, 13−18. (8) Craig, D. W. . U.S. Patent No. 4,992,605, 1991. (9) Jakkula, J.; et al. European Patent No. 1,396,531, A2, 2003. (10) Endisch, M.; Olschar, M.; Balfanz, U.; Kuchling, T. Proceedings of the DGMK conference “Future feedstocks for fuels and chemicals”, Berlin, 2008. (11) Endisch, M.; Olschar, M.; Balfanz, U.; Kuchling, T. Proceedings of the DGMK conference: “Energetische Nutzung nachwachsender Rohstoffe”, Velen, 2008. (12) Simacek, S.; Kubicka, D.; Sebor, G.; Pospisil., M. Fuel 2009, 88, 456−460. (13) Kubicka, D.; Kaluza., L. Appl. Catal. A: Gen. 2010, 72, 199−208. (14) Toba, M.; Abe, Y.; Kuramochi, H.; Masahiro, O.; Mochizuki, T.; Yoshimura, Y. Catal. Today 2010, 164, 533−537. (15) Kubicka, D.; Horacek., J. Appl. Catal. A: Gen. 2010, 394 (1−2), 9− 17. (16) Kikhtyanin, O. V.; Rubanov, A. E.; Ayupov, A. B.; Echevsk, G. V. Fuel 2010, 89, 3085−3092. (17) Yakolov, V. A.; Khromova, S. A.; Sherstyuk, O. V.; Dundich, V. O.; Yu Ermakov, D.; Novopashina, V. M.; Yu. Lebedev, M.; Bulavchenko, O.; Pamon, V. N. Catal. Today 2009, 144, 362−366. (18) Kubickova, I.; Kubicka, D. Waste Biomass. Valor. 2010, vol. 1, 293−308. (19) Baldauf, W.; Balfanz, U. VDI Ber. 1994, 1126, 153−168. (20) Huber, G. W.; O’Connor, P.; Corma, A. Appl. Catal. A: Gen. 2007, 329, 120−129. (21) Lappas, A. A.; Bezergianni, S.; Vasalos, I. A. Catal. Today 2009, 145, 55−62. (22) Simacek, P.; Kubicka, D.; Sebor, G.; Pospisil, M. Fuel 2010, 89, 611−615. (23) Simacek, P.; Kubicka, D.; Kubickova, i.; Homola, F.; Pospisil, M. Fuel 2011, 90, 2473−2479. (24) Olschar, M.; Endisch, M.; Dimmig, T.; Kuchling, T. Eur. Oil Gas J. 2007, 33 (4), 187−193. (25) Weitkamp, J. Erdoel. Kohle. Erdgus.−Petrochem. Brennst.-Chemie 1978, 31 (1), 13−22. (26) J. Reihnardt. PhD-Thesis, TU Bergakademie Freiberg, 2003
2636
dx.doi.org/10.1021/ef400007e | Energy Fuels 2013, 27, 2628−2636
