ing technologies is pyrolysis, which enables us to pro ... Keywords: biomass, bio oil, hydrodeoxygenation, sulfided catalysts, nickel containing catalysts, stability,.
ISSN 20700504, Catalysis in Industry, 2012, Vol. 4, No. 4, pp. 324–339. © Pleiades Publishing, Ltd., 2012. Original Russian Text © V.A. Yakovlev, M.V. Bykova, S.A. Khromova, 2012, published in Kataliz v Promyshlennosti.
BIOCATALYSIS
Stability of NickelContaining Catalysts for Hydrodeoxygenation of Biomass Pyrolysis Products V. A. Yakovleva, M. V. Bykovaa, b, and S. A. Khromovaa, b a
Boreskov Institute of Catalysis, Siberian Branch, Russian Academy of Sciences, Novosibirsk, 630090 Russia b Novosibirsk State University, Novosibirsk, 630090 Russia
Abstract—Heterogeneous catalysts for the hydrotreatment (hydrodeoxygenation) of biomass fast pyrolysis products (biooil) for the production of fuel hydrocarbons are considered. Hydrodeoxygenation catalysts are conventionally divided into three groups: catalysts based on noble metals, sulfided catalysts for desulfuriza tion, and nonsulfided catalysts based on supported transition metals. The main emphasis in this work is on nickelbased nonsulfided catalytic systems as the ones most promising for the hydrotreatment of feedstocks with low content of sulfur. In light of specific features of biooil (high acidity and viscosity, low thermal sta bility), requirements are formulated that must be met in developing hydrodeoxygenation catalysts and pro cesses based on them, especially specifications for the stability of catalysts of this type and their ability for multiple regeneration. Keywords: biomass, biooil, hydrodeoxygenation, sulfided catalysts, nickelcontaining catalysts, stability, deactivation DOI: 10.1134/S2070050412040204
INTRODUCTION There is now keen interest in using organic sub stances of plant origin as alternative fuels, and for the preparation of other useful products [1, 2]. However, many of the current processes of chemical wood pro cessing are inferior to the familiar technologies of oil refining and petroleum chemistry in their productivity and efficiency, requiring heavy equipment and high energy costs. One of the most promising wood process ing technologies is pyrolysis, which enables us to pro cess biomass into fuels, including combustible gas, charcoal, liquid products, and feedstocks for the chemical industry [3, 4]. In order to increase the frac tion of liquid products, the thermal destruction of biopolymers is performed under fast pyrolysis condi tions in which biomass particles are heated rapidly (>1000°C/s) upon contact with a heat carrier, while the pyrolysis products are rapidly cooled outside the pyrolysis zone [5]. The liquid product yield reaches 70 wt % of the initial dry biomass [6]. Liquid biomass pyrolysis products (also known as biooil) are a com plex mixture of oxygencontaining compounds whose elemental composition is close to that of the initial bio mass. Biooil has a higher volumetric energy capacity (21 MJ/L for biooil and 13.4 MJ/L for wood pellets [7, 8]) than wood, making it easier to transport and convert. Of special interest is its joint processing with oil fractions with the use of ordinary oilrefining equipment. However, preliminary biooil processing is needed to get rid of such undesirable properties as high viscosity and acidity (pH = 3) and low thermal stabil
ity, manifested as its tendency to repolymerization. These shortcomings are due to the high oxygen con tent in biooil. In order to improve its quality, technol ogies based on hydrodeoxygenation (HDO) processes must be developed [9]. Up to now, the development of biooil hydrodeox ygenation processes has been oriented primarily to using three types of catalysts: (1) such traditional sul fided catalysts of desulfurization as NiMo/Al2O3 and CoMo/Al2O3 [10, 11]; (2) catalysts based on noble metals (Rh, Ru, Pt, and Pd) deposited on carbon sup ports in addition to TiO2, SiO2, Al2O3, and zeolites β, ZSM5, and MCM41 [12–17]; and (3) heteroge neous nonsulfided catalysts based on transition met als as Ni3266 (50% Ni/SiO2Al2O3) [11], NiMoB [18], carbides and nitrides of Mo and W [19–22], phosphide catalysts Ni2P/SiO2, Co2P/SiO2, WP/SiO2 [23], as well as nickel catalysts modified by various metal compounds [24]. Even though sulfided catalysts are the ones most studied, there are certain difficulties related to their use. It is therefore necessary [25–28] to add a sulfiding agent to the reaction mixture in order to enhance the stability of sulfided catalysts, since the content of sulfur in biooil is not sufficient to keep the catalyst in an active state. In the absence of sulfur containing additives upon the hydrodeoxygenation of biooil, NiMo/Al2O3 and CoMo/Al2O3 catalysts lose sulfur quickly and undergo a reduction that in turn leads to catalyst coking. Additives of sulfiding agents (thiophene, H2S, CS2) do not solve the problem of the low efficiency of these types of catalyst, which reduce
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the total activity of sulfided catalysts due to the con current adsorption of sulfiding additives on catalytic active sites [29, 30]. Moreover, this approach leads to the pollution of model compounds and real biooil hydrodeoxygenation products with sulfurcontaining compounds. This type of catalyst is nevertheless the one most widely encountered in the current literature as a catalyst for the hydrotreatment of biomass fast pyrolysis products and thus deserves separate consid eration. The second type of hydrodeoxygenation catalysts based on noble metals is generally free of the disadvan tages characteristic of sulfided catalysts [31, 32]. Cat alysts based on noble metals exhibit high hydrodeoxy genation activity with respect to both aromatic and ali phatic oxygencontaining model compounds and real biooil. Another advantage of these catalysts is their relatively high stability, since they are less subjected to deactivation due to coking than sulfided catalysts of hydrodesulfurization. At the same time, catalysts based on noble metals also undergo coking. Since mostly catalysts supported on carbon are used in this case, the problem of their regeneration via calcination arises. Due to their high cost, it is unlikely that cata lysts based on noble metals would be used in multiton nage biooil hydrotreatment processes aimed at the production of fuel hydrocarbons. These catalysts are therefore beyond the scope of this work. The most promising type of catalyst for hydrodeox ygenation of biomass pyrolysis products is inexpensive nonsulfided catalyst. Since nickelbased catalysts are among the most widely used hydrogenation catalysts [33], it is understandable that nickelbased systems are the ones most often studied to develop other types of biooil hydrotreatment catalysts. These systems are also considered in this work. It should be noted that biomass fast pyrolysis products (biooil) have a num ber of properties and specific features that must be taken into account in developing conversion processes and catalysts for them. In this context, we must first consider the properties of the mixture of oxygencon taining organic compounds formed upon the fast pyrolysis of grinded wood.
H
OH
O
OH
H
H
H
The elemental composition of liquid products (biooil) obtained during fast pyrolysis differs slightly from that of the initial plant biomass. Biooil is a com plex multicomponent mixture of more than 400 dif ferent organic compounds [34]: aliphatic and aro matic alcohols, esters, ketones, carboxylic acids, sug ars, and so on. These are products of the depolymerization and degradation of plant biomass components (i.e. cellulose, hemicellulose, and lignin) during pyrolysis. Cellulose is a linear polysaccharide built of С6Н10О5 units with the structure shown in Fig. 1. CATALYSIS IN INDUSTRY
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HH O
O CH2OH
CH2OH O H
O
OH
H
H
H
OH n
Fig. 1. Structural formula of a cellulose monomer unit [35].
Hemicellulose is composed of branched polysaccha rides built mainly from С5Н8О4 units with chains shorter than those of cellulose. The aromatic structure of lignin consists of a combination of phenylpropane structural units bound to one another by ether and carboncarbon bonds (Fig. 2). Of the different functional groups, methoxy and phenol groups predominate in lignin. Biooil differs substantially from traditional fuels in composition and does not satisfy the established requirements for liquid motor fuels. It has a lower cal orific value than diesel fuel, high viscosity and acidity; it is also unstable when stored and is prone to polymer ization, especially upon heat treatment [36]. The high acidity and thus corrosiveness of biooil make it unsuitable for internal combustion engines. All of these properties are due to the high content of oxygen (Table 1) found in almost all organic compounds (including biooil components) and water. However, liquids obtained by means of plant biomass pyrolysis have also some advantages over oil. Among these are low contents of nitrogen and sulfur in biooil ( Ni–Mo > Ni–W > Co–W Ni–Mo = Ni–W > Co–Mo > Co–W Ni–Mo > Co–Mo > Ni–W > Co–W Ni–W > Ni–Mo > Co–Mo > Co–W
Nickel and Co–Mo/Al2O3 catalysts were therefore used for hydrodeoxygenation in [43]. When Ni cata lysts were used at 250°C, a hydrogen pressure 14 MPa, and a biooil hourly space velocity (LHSV) of 0.32– 0.45 h–1, the content of oxygen in biooil fell from 45 to 20–25 wt %. The use of Co–Mo catalyst produced a similar result at 270°C. However, both catalysts were rapidly deactivated due to coking. To solve the prob lem of catalysts coking, a twostep biooil processing CATALYSIS IN INDUSTRY
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scheme was proposed in [44]. In the first step, the pro cess is performed in the presence of Ni or Co–Mo cata lyst at reduced temperatures (250–300°C), yielding stabilized products with an oxygen content of 26– 27%. At this stage, aliphatic and aromatic oxygen containing compounds partially lose their oxygen groups, that reduces their ability to polymerize at higher temperatures. (This is the reason for the stabili zation effect.) At the second stage, the mixture is treated with hydrogen at 14 MPa and a temperature of 350°C on a sulfided CoMo/Al2O3 catalyst until the content of oxygen is reduced to 2–3 wt %. In the two step process, the sulfided catalyst is more stable, since biooil acidity falls after the first step. This is a positive factor, as γAl2O3 is soluble in acidic media. An inter esting approach was employed by Zhang et al. [10], who separated liquid pyrolysis products into two phases: aqueous and organic. The organic fraction was subjected to hydrodeoxygenation in the presence of Co–Mo–P/γ–Al2O3 catalyst at a hydrogen pressure of 2 MPa and a temperature of 360°C. (It was shown that this temperature is optimal for the given process and reactants.) Under these conditions, the oxygen content was successfully reduced from 42 to 3 wt %. Considering the multicomponent nature of biooil, the examples presented above show that more detailed investigations with model compounds [30, 45, 46] and
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their mixtures [47–50] are required for its efficient deoxygenation. This would allow us to determine both the individual reactivities for each type of oxygen containing organic biooil components and the effect of other parameters of the process on the activity and stability of sulfided catalysts. Without going into details of the hydrodeoxygen ation mechanism on sulfided hydrotreating catalysts, we must emphasize some of its specific features. The regularities of hydrodeoxygenation reactions have been investigated for certain types of biooil compo nents: phenol derivatives [51, 52], benzofuran [53, 54], and aliphatic compounds [55, 56]. The obtained data on the selectivity of product formation revealed two parallel paths for the reaction, one of which is direct hydrodeoxygenation without aromatic ring hydrogenation and the other is benzene ring hydrogenation without oxygen removal. Two types of active site that participate in the above mentioned processes were described in [57]. Sites for aromatic ring hydrogenation have electronacceptor properties, with reactant being adsorbed on them in its plane, i.e., the π system of the ring. Sites of the second type are of an electrondonor character, with reactant being adsorbed on them vertically with respect to the catalyst’s surface, i.e., through a heteroatom. The same regularities were also obtained in [58], devoted to studying the hydrodeoxygenation kinetics of mono and disubstituted phenols on a commercial sulfided CoMo/γAl2O3 catalyst at 300°C and 5 MPa H2. In addition, the selectivity of toluene formation dimin ished considerably upon adding H2S, which evidences the different effects hydrogen sulfided has on reaction rates of the first and second routes. Kinetic investiga tions confirmed the authors’ assumption that different active sites of the catalyst are responsible for the reac tion proceeding along both routes. On sites with direct hydrodeoxygenation, reactants are adsorbed via oxy gen σbonds and scissions of C–O bonds are observed; on the second type of sites, aromatic rings are adsorbed along with their π bonds and then hydroge nated. The authors of [59] believed that, as in the case of hydrodesulfidation, oxygen can be removed from organic compounds both before and after benzene ring hydrogenation, which does not contradict the concept of two types of active sites. The deactivation of sulfided catalysts during the hydrodeoxygenation of oxygencontaining organic compounds is usually attributed to their coking and the formation of highmolecular weight products, and to oxygen replacing the sulfur in sulfided structures. In addition, the water that forms during the reaction can also serve as a weak inhibitor of the hydrodeoxygen ation reaction [60]. In view of the presence of two types of active sites described above, it is of interest to consider the effects of various factors (H2O, coking, etc.) on their stability.
Water is a product of the hydrodeoxygenation reac tion (oxygen being eliminated from oxygencontain ing compounds in the form of H2O), and its effect on the process can be noticeable if it is formed in appre ciable amounts (at high conversion). Viljava et al. [29] thought that the unselective water blocking of active sites should lower the rates of product formation along both reaction routes in the absence of a sulfiding agent. According to Vogelzang et al. [25], however, who inves tigated the stabilities and activities of the sulfided and oxide forms of Mo/γAl2O3 and NiMo/γAl2O3 cata lysts in the hydrodeoxygenation of 1naphthol at 225°C and 12 MPa H2, the deactivation attributed to the presence of water formed during the reaction was accompanied by changes in the selectivity of product formation, the rate of formation along the HDO route (C−O bond hydrogenolysis) falling faster than for the ring hydrogenation route. This effect was more sharply pronounced for the sulfided form of the catalyst. After calcination in hydrogen at a temperature of 400°C, the activity of catalysts was virtually restored, confirming the authors’ assumption that the catalyst’s active sites are blocked upon the adsorption of water. On the other hand, the data obtained by the authors support the assumption that there are two types of active sites, one for HDO and one for hydrogenation, for both sulfided and oxide forms of the catalyst. The electron donor sites responsible for the direct hydrogenolysis of C–O bond are affected by water to a greater extent than the electron acceptor sites of the second type. In [55], the effect of water on the activity of com mercial NiMo/γAl2O3 and CoMo/γAl2O3 sulfided catalysts was also investigated in the hydrodeoxygen ation reaction of methyl and ethyl esters of heptanoic acid. A reactor was supplied with hydrogen saturated with water vapors. The ester/water molar ratio was 1.6 at the maximum content of water. In all experi ments, the catalyst was deactivated over time, the activity falling by ~10% in 4 h. When water was added to the reaction mixture, conversion fell in all cases; additional hydrogen sulfided increased the activity of the catalysts; and the simultaneous addition of water and H2S lowered conversion to the value attained in the absence of water and hydrogen sulfide. Even though the presence of water reduces the conversion of reactants, the authors considered the losses of sulfur and the coking of the catalyst (the content of carbon was 6.3 wt % for NiMo and 7.5 wt % for CoMo after 3 h of reaction) to be the primary reasons for deactiva tion. Bredenberg et al. also noted [26] it is not only the presence of water that affects the stability of sulfided catalysts. Upon hydrodeoxygenation of some aro matic compounds on the commercial sulfided catalyst NiMo/Al2O3–SiO2 in the temperature range of 350– 400°C, the catalyst was deactivated quite rapidly (in several hours) even when there was no solvent, i.e., when the reactor was supplied with pure oxygencon taining reactants. The authors attributed this to the CATALYSIS IN INDUSTRY
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effect of water formed during reactions on the catalyst, to losses of sulfur by active sites, and to coking of the catalyst. In contrast, Odebunmi et al. [49] observed no effect of water on deactivation. They performed the hydrodeoxygenation of o, m, and pcresols on a commercial sulfided CoMo/γAl2O3 catalyst at a hydrogen pressure of 68 atm and T = 400°C with nhexadecane as the solvent, with the reactor also being supplied with water. The main conversion prod ucts of all cresols were toluene and methylcyclohex ane. The results showed that water had no appreciable inhibiting effect on the reaction, since the mcresol conversion did not depend on the presence of water. Since sulfided catalysts must be kept in their sul fided form in order to avoid loss of their activity [26– 28, 61], it is of interest to investigate the effect of sul fiding agents (usually H2S or CS2) on the stability of catalysts and the selectivity of reaction products for mation along the two reaction routes described above. Most of the works devoted to this topic reported a neg ative effect of sulfiding agents on the activity of cata lysts during hydrodeoxygenation, their presence hav ing little or no effect on the conversion of the initial substance but lowering the selectivity of targeted prod uct formation. Bredenberg et al. showed [26] that dur ing the hydrodeoxygenation of a model compound (anisole) on a sulfided CoMo/C catalyst in the pres ence of hydrogen sulfide, the degree of deoxygenation fell considerably as the reaction of Car−O bond hydro genolysis with the formation of aromatic compounds was suppressed, while aromatic ring hydrogenation was weakly affected by H2S. This also confirms the assumption that there are two types of active catalyst sites that participate in the hydrodeoxygenation pro cess. Like most researchers, Senol et al. [62] believed that the process of C−O bond hydrogenolysis proceeds on coordinatively unsaturated active sites. It is this fea ture that is associated with the inhibiting effect of H2S on the reactions, leading to the formation of hydrocar bons from phenol. They investigated the effect of H2S as a sulfiding agent on the activity of commercial NiMo/γAl2O3 and CoMo/γAl2O3 catalysts in the hydrodeoxygenation of aromatic and aliphatic biooil model compounds. It was found that hydrogen sul fided suppresses hydrogenolysis and hydrogenation reactions on NiMo catalyst, while hydrogenolysis was mainly suppressed on CoMo if phenol was used as the reactant. The inhibiting effect of H2S was related by the authors to the concurrent adsorption of phenol and hydrogen sulfided on coordinatively unsaturated active sites. Ferrari et al. [28] investigated the effects of hydrogen partial pressure and sulfiding temperature on the activity and selectivity of CoMo/C catalysts during hydrodeoxygenation. The carbon support had a specific surface of 1100 m2/g. This support was selected to avoid catalyst deactivation due to coking, since it occurs on alumina due to the presence of acidic centers on its surface. A mixture of model com CATALYSIS IN INDUSTRY
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pounds (guaiacol, ethyl decanoate, and 4methyl acetophenone) was used for the reaction. HDO was conducted at a temperature of 270°C for guaiacol and ethyl decanoate, and at 200°C for 4methylacetophe none, with a hydrogen pressure of 7 MPa. The prod ucts of guaiacol conversion were phenol, catechol, benzene, cyclohexane, and cyclohexene. It was shown that the pressure of hydrogen sulfided does not affect guaiacol conversion on CoMo/C catalyst. However, hydrogen sulfided does suppress direct Cаr−O bond hydrogenolysis, thus affecting the selectivity of the process. Viljava et al. [27] investigated the effect of a sulfiding agent (H2S) on the activity and stability of CoMo catalyst in the hydrodeoxygenation of phenol and anisole. It was shown that the degree of deoxygen ation drops considerably in the presence of hydrogen sulfide, since hydrogenolysis with the formation of aromatic compounds is suppressed while aromatic ring hydrogenation is weakly affected by H2S. On the other hand, there are data that also show H2S has no effect on selectivity. Huuska et al. [61] investigated the influence of H2S on the activity of catalysts in the hydrodeoxygenation of anisole (the reaction mixture was supplied with 0.5 vol % CS2, which was trans formed into H2S under the reaction conditions). The concentration of CS2 in these experiments was 10 times higher than that normally required to keep the catalyst in its sulfided state. The test results showed that an increase in the concentration of sulfiding agent affects neither anisole conversion nor the selectivity of its transformation products. In general, the data of the abovementioned works complement each other. We may therefore conclude that the primary role in the hydrodeoxygenation of oxygencontaining aromatic compounds belongs to the coordinatively unsaturated centers on which the elementary acts of both aromatic ring hydrogenation and C−O bond hydrogenolysis occur. Gandarias et al. [63] studied the activity and stabil ity of phosphorusdoped NiMo catalyst in its sulfided and reduced forms during phenol hydrodeoxygen ation. They associated the drop in the activity of reduced NiMo catalyst after 1100 min of reaction with catalyst coking. With sulfided NiMo catalyst, the cat alyst activity increased with the reaction time at low temperatures and decreased at high temperatures. The authors explained this behavior by saying NiMo cata lyst is more active in reduced form than in sulfided form: it loses sulfur at low temperatures without sulfur donors, and its activity increases. At high temperatures (300–350°C), the positive effect due to catalyst sulfur losses is compensated by a drop in activity due to cok ing, the latter proceeding more intensively at high temperatures than at low ones. Bui et al. [64] investi gated the stability of CoMo and NiMo catalysts on var ious supports in their sulfided and reduced forms during guaiacol hydrodeoxygenation. The authors noted that using γAl2O3 as a support raises the yield of phenol
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methylation products due to the presence of Lewis acidic centers on the support’s surface. With γAl2O3, a high yield is observed for the highmolecular weight products that are precursors of carbon deposits. The deactivation of sulfided catalysts during the hydrodeoxygenation of oxygencontaining organic compounds is thus related to their coking and the for mation of highmolecular weight products, along with oxygen replacing sulfur in sulfided structures. The support for sulfided catalysts is usually γAl2O3. Due to the presence of acidic centers on their surfaces, such catalysts are quite rapidly deactivated as a result of coking, even if the reaction is conducted with a mix ture of model compounds. Carbon supports present no such disadvantage, but the regeneration of catalysts is difficult when they are used. The sulfiding agents (thiophene, H2S, CS2) used as additives do not solve the problem of the low efficiency of this type of cata lyst, which lowers the overall activity of sulfided cata lysts due to concurrent adsorption of sulfiding addi tives on the active sites of a catalyst. This approach also pollutes the hydrodeoxygenation products of both model compounds and real biooil with sulfurcon taining compounds. Considering the negative effect of water on sulfided catalysts, the prospects for using them in the hydrodeoxygenation of biooil with high water content is also doubtful. In this context, the rel evance of developing nonsulfided catalysts for biooil hydrodeoxygenation seems obvious. However, their susceptibility to coking, their stability in acidic media in the presence of water, and the possibilities of the catalyst regeneration must also be taken into account. NONSULFIDED NICKELCONTAINING HYDRODEOXYGENATION CATALYSTS The abovementioned specific features of sulfided catalysts in the hydrodeoxygenation of model oxygen containing organic compounds and real biooil stimu lated researchers to search for new types of biooil hydrotreatment catalysts. One such approach was test ing oxide forms of conventional oil hydrotreating cat alysts in the hydrodeoxygenation of biomass pyrolysis products. The data from experiments by Vogelzang et al. [25] on the hydrodeoxygenation of 1naphthol indicate that the activity of Mo/γAl2O3 and Ni Mo/γAl2O3 catalysts in their nonsulfided form is higher than that of the sulfided form. At the same time, contradictory results were obtained in [27], in which a commercial hydrotreating catalyst (Ketjen fine 7421.3Q, Akzo Chemie), containing 4.4 wt % CoO and 15 wt % MoO3 on γAl2O3, was used in its sulfided and oxide forms. The results from testing the catalyst in the hydrodeoxygenation of anisole at 1.5 MPa H2 and 250°C showed a higher initial conver sion on sulfided catalyst (88.2 %) than on the oxide from of the catalyst (63.3 %). Moreover, the degree of deoxygenation (i.e., the yield of oxygenfree products)
was 10.4% on sulfided catalysts, while no such prod ucts were found on its oxide analog. It was shown that the conversion of anisole on the oxide catalyst was due to the presence of acidic centers on the alumina sur face. Hydrodeoxygenation apparently did not occur in this case, since no reduction of NiO to Ni0 was observed at 250°C, the latter being responsible for cat alyst activity in the hydrogenation and hydrodeoxy genation reactions. It should be noted that the reactiv ity of oxygencontaining aromatic compounds is much lower than that of oxygencontaining aliphatic compounds. Bunch et al. [30] compared the activity of NiMo/γAl2O3 in its reduced and sulfided forms in the reaction of benzofuran hydrodeoxygenation. The con version of benzofuran on the reduced sample was higher than that on the sulfided sample over the inves tigated range of temperatures (200–320°C). At a tem perature of 200°C, conversion was 80 and 35%, respectively. Nevertheless, no hydrodeoxygenation was observed, while products of the partial hydrogena tion of benzofuran with fivemembered ring opening products being formed without the elimination of oxy gen. When the temperature was raised to 320°C, the yield of fivemembered ring opening products increased, while the degree of deoxygenation was still close to zero. Consequently, regardless of the state of bimetallic Ni–Mo or Co–Mo phase (sulfided, oxide, or reduced), this type of catalysts is not promising for biooil hydrodeoxygenation, and other types of hydro genolysis catalysts are needed. Traditional nickelbased hydrogenation catalysts have also been considered as biooil hydrotreatment catalysts. Ni1404 nickel catalyst (68% nickel) exhib its approximately the same activity as that of sulfided CoMo/γAl2O3 catalyst during biooil hydrodeoxygen ation at 340°C and a hydrogen pressure of 16.8 MPa [65]. The nickel catalyst was, however, deactivated after several hours, exhibiting lower activity than the sulfided catalyst. The main product of biooil hydrotreatment was in this case a light gasoline frac tion with Tboil = 50–225°C. The reason for the drop in the activity of nickel catalyst was apparently nickel phase coking. In addition, the yield of undesirable products (i.e., light C1–C4 hydrocarbons) was as high as 15% in the presence of Ni1404, reducing the yield of targeted products and increasing H2 consumption. At the same time, the yield of C1–C4 hydrocarbons on sulfided catalysts was only 1–2%. A high yield of gas eous products was also obtained in [11] during biooil hydrodeoxygenation at 400°C in the presence of Raney nickel. As was mentioned above [44], one advanced approach is a twostep biooil processing in which the first stage is performed at a lower temperature (250– 280°C), making it possible to avoid early catalyst cok ing and the repolymerization of oxygencontaining organic components. As a result of lowtemperature treatment, the content of oxygen in biooil falls from CATALYSIS IN INDUSTRY
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35–45 wt % to 22–28%, while the H/C ratio is ∼1.5. The second stage is performed at higher temperatures (350–360°C), allowing us to obtain products with low oxygen contents (2–3%) and H/C ratios of ∼1.7. It should be noted that nonsulfided catalyst like Ni 3266 (50%Ni/SiO2–Al2O3) is used only in the first stage of hydrotreatment, its primary aim being the par tial deoxygenation and enrichment of processing products with hydrogen. As was shown in [66], the use of Ni3266 at elevated temperatures (310–340°C) results in blocking the flow reactor with coking prod ucts. It should be noted that typical nickel catalysts of hydrogenation are more susceptible to coking upon biooil hydrotreatment than sulfided hydrotreating catalysts. Since hydrodeoxygenation on the catalyst’s surface proceeds via an oxred mechanism, the corre sponding reaction conditions are needed for the reversible reduction of surfaceoxidized nickel. At lower temperatures (250–280°C is the recommended temperature for the first stage), the reversible reduc tion of oxidized nickel in a monometallic catalyst would seem to be unlikely [67], which in turn leads to leaching of the active components of nickel catalysts. Nickel catalysts for biooil hydrodeoxygenation therefore appear to be in an equivocal situation: on one hand, they should not be used at low temperatures to avoid leaching; on the other hand, they undergo rapid coking at elevated temperatures. To raise the activity and stability of nonsulfided catalysts, includ ing nickelcontaining systems, special approaches to developing efficient biooil hydrotreatment catalysts must thus be adopted, including unconventional methods for preparing and modifying catalysts with different compounds. Examples are nickel and cobalt catalysts prepared by reduction of salts with sodium borohydride NaBH4 [18, 68–70]. Amorphous catalysts prepared in this way have been a focus of attention for the last 20 years, due to their highly disperse morphologies and high con centrations of coordinatively unsaturated centers. These catalysts have also been tested in hydrogenation [70, 71] and hydrodechlorination [72, 73] reactions, and their high activity in hydrogenolysis has been demonstrated. It is obvious that these catalysts should also exhibit a certain level of activity in hydrodeoxy genation reactions, which was proved in [18, 68–70]. When Ni and Co salts are reduced by sodium borohy dride, an amorphous highly disperse phase of Ni2B or Co2B composition is formed that is responsible for the high activity of these catalysts during hydrogenolysis under mild conditions. Wang et al. [18] used XPS to show that the activity of a bimetallic NiMoB system depends on the amount of Mo+4 phase formed as a result of the reduction of ammonium heptamolybdate by NaBH4: Mo+6 → Mo+4. A synergic effect was noted for Ni and Mo in the hydrodeoxygenation reaction of phenol. We must therefore conclude that Ni active sites are responsible for hydrogen activation, while CATALYSIS IN INDUSTRY
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Mo+4 sites are responsible for the additional activation of oxygencontaining organic compounds. It was also shown that the optimum Ni/Mo ratio is 1/2, the same as for sulfided catalysts. Raising the temperature to 250°C nevertheless causes a sharp drop in the activity of the boride catalyst due to its sintering. The thermal stability of boride catalysts can be increased either by depositing the active component on a support [74] or by their stabilization with other transition metal com pounds. The effect of adding Co and La (15% with respect to Ni) to Ni–Mo–B catalyst on phenol hydro deoxygenation at 4.0 MPa H2 was investigated in [68]. After introducing Co and La into the catalyst, the mean size of amorphous phase particles fell and the specific surface area grew from 38.4 m2/g for Ni–Mo– B to 52.7 m2/g for Co–Ni–Mo–B and 116.6 m2/g for La–Ni–Mo–B. Based on the degrees of crystallinity of amorphous samples at the temperature of 275°C determined by XRD, an increase in thermal stability was also observed, this effect being higher upon the introduction of La than in the case of Co additives. The highest yield of deoxygenated products from phe nol hydrodeoxygenation at 250 and 275°C (91.2 and 73.9%, respectively) was produced by Co–Ni–Mo–B catalyst. When the reaction temperature is raised to 275°C, however, the total phenol conversion fell, probably indicating the sintering of the catalyst’s amorphous phase after 10 h of reaction or the intensi fication of catalyst coking. In another work by the same authors [69], the effect of the composition of CoMoB catalyst on its activity in phenol, benzaldehyde, and acetophenone hydrode oxygenation was investigated by varying the ratio Co/Mo. The highest activity was displayed by catalyst with a Co/Mo ratio of 1/3, which is close to the opti mum metal ratio for conventional sulfided Co– Mo/Al2O3 desulfurization catalyst [75]. It was also shown by EM and XRD that this catalyst displays noticeable thermal stability upon calcination (300°C), maintaining its amorphous state. At the same time, it should be noted that biooil hydrotreatment must be performed at higher temperatures (300–400°C), and the multiple calcination of coked catalysts for their regeneration at temperatures of 500–600°C is implied. Will boride systems be stable at such high temperatures? The question remains open, but this line of biooil hydrodeoxygenation catalyst develop ment is undoubtedly a promising one that deserves to be continued. When considering hydrodeoxygenation catalysts for biomass pyrolysis products, attention is inevitably drawn to the number of works devoted to investigating Mo and W carbides and nitrides as catalysts for hydrodesulfidation, hydrogenation, and hydrodeoxy genation [19–22], primarily in oil hydrotreating. It was shown for βMo2C that there are two types of active sites on the catalyst’s surface when the three indicated processes occur simultaneously. Molybde
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Table 2. Comparison of activities of phosphide catalysts based on transition metals and commercial catalysts in HDO of guaiacol (τ = 0.339 min, LHSV = 59 h–1, 300°C [23]) Sample
Ni2P/SiO2
Co2P/SiO2
WP/SiO2
5% Pd/Al2O3
CoMo/Al2O3
Conversion, %
19
35
12
70
1
28
1
12
0
0
4
0
0
0
0
38
0
0
0
100
0
99
88
100
0
30
0
0
0
0
Selectivity, % Phenol Benzene Methoxybenzene Catechol Cresol
num carbide is responsible for the activity during hydrogenation and hydrodeoxygenation. The second type of active site is molybdenum carbosulfide, to which the activity in all three processes is attributed. It should be noted that as a substrate for deoxygenation, only 2% of benzofuran in the reaction mixture lowers the rate of cumene hydrogenation at 250°C and 5.1 MPa H2 by half. It was nevertheless shown that oxygencontaining organics are reversible inhibitors of hydrogenation, while sulfurcontaining compounds (e.g., dibenzothiophene) are catalytic poisons for Mo2C as a cumene hydrogenation catalyst. Since bio oil has a low content of sulfur, we may suggest that these types of catalysts, along with nickel systems, have high potential for biooil hydrotreatment. The use of phosphide catalysts deposited on SiO2 in hydrodeoxygenation reactions of oxygencontaining organic compounds [23, 76] deserves separate conside ration. This type of catalyst was prepared by impreg nating SiO2 with a solution of the appropriate transi tion metal salts and ammonium phosphate, followed by calcination and reduction in situ in H2. This process is described in detail in [77–79]. The results from investigating guaiacol (2methoxyphenol) hydrodeox ygenation at 300°C and 0.1 MPa H2 in a flow reactor at various contact times were presented in [23]. Ni2P/SiO2, Fe2P/SiO2, MoP/SiO2, Co2P/SiO2 and WP/SiO2 were used as catalysts and were compared to commercial catalysts 5%Pd/Al2O3 (BASF) and CoMoS (Haldor Topsoe). After a rather long contact time (τ = 20.2 min), the relative activity series for phosphide catalysts was found to be Ni2P/SiO2 > Co2P/SiO2 > Fe2P/SiO2 > WP/SiO2 > MoP/SiO2. The primary products in these cases were benzene and phenol. Testing the catalysts at different temperatures (200–300°C) for 100 h showed that Ni2P/SiO2 and Co2P/SiO2 had much higher activities and stabilities than those of all other catalysts. The decline in catalyst activity related to their coking was sharpest for Fe2P/SiO2. Three phosphide catalysts (Ni2P/SiO2, Co2P/SiO2, and WP/SiO2) were selected for compari
son with commercial catalysts 5%Pd/Al2O3 and CoMoS/Al2O3. The test results are presented in Table 2. As was expected, shortening the contact time from 20.2 to 0.339 min considerably reduces the conversion of guaiacol for all catalysts. The fully deoxygenated product (benzene) was observed only in the presence of Ni2P/SiO2. Partial deoxygenation of guaiacol to phenol and methoxybenzene occurred in the presence of Ni2P/SiO2 and CoMo/Al2O3. For all of the other catalysts, the primary product was catechol C6H4(OH)2, which was not a target deoxygenation product. At short contact times, the series of relative activities for hydrodeoxygenation catalysts is some what different from that at τ = 20.2 min: Ni2P/SiO2 > WP/SiO2 > CoMo/Al2O3 > Co2P/SiO2 > 5% Pd/Al2O3. The prospects for this type of catalysts should be noted due to the high thermal stabilities of both phosphide and phosphate phase of the abovementioned transi tion metals. This is important, since it prevents sinter ing of the active component at elevated temperatures of the target process, and at the stage of aglomeration. The authors of [24, 80–82] suggested using bime tallic nickel–copper hydrodeoxygenation catalysts that display high activity in both the hydrogenation and hydrodeoxygenation of aromatic and aliphatic oxygencontaining compounds. Such typical model compounds as anisole (methoxybenzene) and guaia col were selected in the first stage of developing cata lysts for hydrodeoxygenation of biomass pyrolysis products, while δAl2O3 calcined at 1000°C was cho sen as the support after catalysts screening. A series of Ni–Cu catalysts was prepared with different Ni/Cu ratios and were tested in the HDO reaction of anisole at 300°C and 1.0 MPa H2 in a flow reactor (Table 3). Silica sand served as an inert diluent for the catalyst. Data on the selectivity of products formation in HDO of anisole on Ni–Cu/δAl2O3 catalysts are pre sented in Table 4. The 24.5Cu sample demonstrated the highest anisole conversion (Table 5) but a very low degree of deoxygenation (1%), with phenols as the main products. The higher ratio of aliphatic/aromatic CATALYSIS IN INDUSTRY
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Table 3. Activities of NiCu/δAl2O3 catalysts in anisole HDO reaction at 300°C, 1.0 MPa H2, and LHSV = 6 h–1 [24] Sample1
Anisole conversion Degree of oxygenation2, Yield of deoxygenated Specific activity, (XA), % mol % products (DP), mol % 107 ⋅ mol DP/s ⋅ gNi–Cu
Quartz δAl2O3 20.8Ni/δAl2O3 16Ni2Cu/δAl2O3 13.8Ni6.83Cu/δAl2O3 13.3Ni11.8Cu/δAl2O3 5.92Ni18.2Cu/δAl2O3 24.5Cu/δAl2O3
2.8 11.8 66.1 78.6 73.8 70.3 76.9 95.3
0 0 97.8 95.9 90.6 82.8 72.8 1.0
0 0 64.6 75.4 66.9 58.2 56.0 1.0
– – 5.0 7.3 6.4 6.1 6.9 0.2
1 The numbers in the catalyst designations correspond to the weight percentage of transition metals in the catalysts. 2
The degree of deoxygenation corresponds to the selectivity of deoxygenated product (cyclohexane, benzene, methylcyclohexane, tol uene) formation.
Table 4. Selectivities (%) of products formation in anisole HDO reaction on Ni–Cu/δAl2O3 catalysts at 300°C, 1.0 MPa H2, and LHSV = 6 h–1 Active component composition, wt % Product 24.5Cu
5.92Ni 18.2Cu
13.3Ni 11.8Cu
13.8Ni 6.83Cu
16Ni 2Cu
20.8Ni
Cyclohexane
0
29.2
27.7
32.8
24.3
14.9
Benzene
0
28.6
42.2
43.5
59.9
72.5
Methylcyclohexane
1.0
7.8
5.3
5.5
8.8
2.4
Toluene
0
7.1
7.5
8.8
2.9
8.1
Cyclohexanol
0
0
2.3
5.9
2.3
0
Cyclohexanone
0
2.7
0
0
1.8
0
99
22.6
13.3
1.2
0
2.2
0
1.9
1.6
2.3
0
0
Phenols Other products
hydrocarbons obtained using Ni–Cu catalysts relative to Ni catalyst indicates that Ni–Cu systems are more active in the hydrogenation of the anisole aromatic ring. The relatively low yield of aliphatic oxygencontaining anisole conversion products (cyclohexanol and cyclo hexanone) on the series of Ni–Cu/δAl2O3 (0–4.4%) samples indicates that Car–O bond scission either dominates over benzene ring hydrogenation or pre cedes it. The optimum relationship between conversion and the degree of deoxygenation is reached when using sample with small amount of copper–16%Ni2%Cu/ δAl2O3. Investigation of this series of samples by tem peratureprogrammed reduction (TPR) and XRD showed that a solid Ni1 – xCux solution is formed after adding copper to nickelcontaining catalysts, which is one of the reasons for the reduction of NiO at lower temperatures (410°C for Ni/δAl2O3 and 270–330°C for NiCu/δAl2O3). Adding copper to nickel catalysts CATALYSIS IN INDUSTRY
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thus allows the nickel oxide to be reduced, transform ing it into an active state under the target process con ditions. A series of Ni–Cu/δAl2O3 samples was also tested in the hydrodeoxygenation of biooil (VTT, Finland) at the university of Groningen in the Nether lands [24]. The process was conducted in an autoclave at an initial pressure of 11 MPa H2. The temperature was kept at 150°C for the first hour and was then raised to 350°C. The process continued at this temperature for another 3 h. The experimental results are presented in Fig. 4 in a Van Krevelen plot characterizing the ele mental composition of HDO products by O/C and H/C ratios. It follows from the data that the H/C ratio in the reaction products grows along with the nickel content in the samples (except for 20.8Ni). This is a positive factor in biooil hydrotreatment, since it low ers the viscosity and molecular weight of the obtained products. The increase in the H/C ratio testifies to the enhanced activities of the investigated samples during
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YAKOVLEV et al. 0.6 Biooil (VTT) 0.5
O/C, at
0.4 0.3 Ni/Сu 13.3/11.8
0.2 0.1
Without catalyst
Ni/Сu 13.8/6.83 Ni/Сu 5.92/18.2
Сu 24.5 0 0.8
0.9
1.0
Ni/Сu 16.2/2
Ni 20.8
1.1
1.2 H/C, at
1.3
1.4
1.5
1.6
Fig. 4. Van Krevelen diagram for biooil hydrodeoxygenation products on NiCu/Al2O3 catalysts in an autoclave at 11 MPa H2, the temperature regime: 150°C (1 h) + 350°C (3 h).
hydrogenation, with the same effect also observed upon anisole hydrodeoxygenation. Investigations of this series of catalysts after biooil HDO showed that the phase composition and crystal lite size of active components changed in all of the samples. The precursors of the active components in the initial NiCu/δAl2O3 catalysts were NiO and CuO oxides, their crystallite sizes growing along with the content of Ni and Cu in the samples. After reacting with biooil, the precursors of the active components are reduced to their metallic states. Considerable growth in the size of nickelcontaining particles is observed relative to the original ones, due to the sinter ing of particles during the reaction. In contrast to nickelcontaining particles, the size of copper crystal lites falls after the reaction due to the dispersion of copper after reduction as a result of the formation of nickel–copper phase. The content of carbon in sam ples of NiCu/δAl2O3 catalysts after biooil hydrode oxygenation was determined by TGA: Active component composition Content of carbon deposits, wt % Active component composition 24.5Cu 5.92Ni18.2Cu 13.3Ni11.8Cu 13.8Ni6.83Cu 16Ni2Cu 20.8Ni
Based on the test results, it was concluded that the employed Ni–Cu catalysts should be improved to increase their stability with respect to active compo nent agglomeration and coking. Investigations of dif ferent supports under conditions close to those for biooil HDO (high acidity and the presence of water) showed that such supports as CeO2–ZrO2, ZrO2, and TiO2 were promising [24]. A series of Ni–Cu catalysts was therefore prepared on various supports and tested for biooil HDO under the same conditions as those for the series of NiCu/Al2O3 catalysts. The experimen tal results are presented in Fig. 5. The weight fraction of nickel in the case of NiCu/C(1) and NiCu/C(2) catalysts was ~18%, while that of copper was ~8%. In all other cases, the nickel and copper contents were 8 and 3 wt %, respectively.
Content of carbon deposits, wt %
The data on commercial Ru/C (5% Ru, Sigma Aldrich) catalyst are presented for comparison. It can be seen that at close degrees of deoxygenation corre sponding to approximately the same O/C atomic ratio in the reaction products, the highest H/C ratio in the series of Ni–Cu catalysts was achieved on NiCu/ZrO2 and NiCu/TiO2: 1.37 and 1.43, respectively.
24.4 15.6 15.6 19.3 8.9 15.0
It is evident that the next stage in the development of biooil hydrodeoxygenation catalysts should be the multicycle testing of catalysts with intermediate regeneration by means of carbon deposit calcination. The revealed regularities will be used for further opti mization of the catalysts’ composition in order to increase their thermal stability. CATALYSIS IN INDUSTRY
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0.6 Atomic ratio O/C
EFFECT OF BIOOIL PROPERTIES ON THE EFFECTIVENESS OF ITS HYDROTREATMENT Specific features of biooil include its high viscos ity, chemical complexity and high content of high molecular weight compounds, and its low thermal sta bility. As a result, major roles in hydrotreatment are played by the reactor geometry, pressure, and temper ature of the process, along with the type of catalyst used [83–86]. Venderbosch et al. [85] proposed a mechanism for the catalytic hydrotreatment of biomass pyrolysis products that allowed for these features of biooil. According to this mechanism, pyrolysis liquid (bio oil) can be transformed in two ways: (1) With no catalyst and/or hydrogen, repolymer ization reactions proceed vigorously at temperatures above 175°C, with the formation of highmolecular weight compounds and subsequent coking. 2) With a catalyst, the pyrolysis liquid is stabilized at temperatures 175–250°C and high hydrogen pres sure (20 MPa), due mainly to the hydrogenation reac tions and partial hydrodeoxygenation. A product is then formed that can undergo further upgrading by means of hydrodeoxygenation and hydrocracking at temperatures above 250°C with no considerable con tribution from the undesirable repolymerization pro cess. Recent experiments performed by Wildschut et al. [87] showed that similar processes are observed in the hydrotreatment of glucose, a model compound of the carbon hydrate fraction of biooil. Glucose was hydrotreated in an aqueous solution at 250°C and 10 MPa Н2 with Ru(Pd)/C catalyst in an autoclave under vigorous stirring. Upon thermal treatment with no catalyst, waterinsoluble humins (coke) formed. With catalyst, the glucose was hydrogenated with the formation of polyols (sorbitol). Venderbosch et al. [88] proposed a simplified temperature dependence for the transformation rate of biooil components (mol/min) with allowance for the following processes: the mass transfer of hydrogen from the gas to the liquid phase, hydrogenation (Rhyd), and the polymerization of bio oil components (Rpol) (Fig. 6). The following was kept in mind: The rate of hydrogen mass transfer from the gas to the liquid phase was determined as k g −1a1ΔcH2 , where kg – l is the diffusion coefficient, al is the phase contact area, and Δ cH 2 is the difference between the hydrogen concentrations in the gas and liquid phases. The mass transfer of hydrogen thus depends on the reactor’s geometry and the reaction parameters (T and PH 2 ). Using static reactors (autoclaves) with stirring, the stirring rate is important, as it affects kg –l al; for flow reactors with a fixed catalyst layer, the wetting of cata lyst particles is important, as it affects al. In addition to the size of the catalyst particles and temperature, the
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0.5
Biooil
0.4 0.3 0.2 Without NiCu/C(1) NiCu/ZrO2 NiCu/TiO2 catalyst NiCu/SiO 2 0.1 Ru/C NiCu/C(2) NiCu/CeO2ZrO2 0 0.8 0.9 1.0 1.1 1.2 1.3 1.4 1.5 1.6 Atomic ratio H/C
Fig. 5. Van Krevelen diagram for biooil hydrodeoxygen ation products on NiCu catalyst with different supports at 11 MPa H2 and temperatures of 150°C (1 h) and 350°C (3 h).
partial pressure of hydrogen is important for both sys tems, as it directly affects the parameter Δc H2 . Tem perature has a considerably smaller effect on hydrogen mass transfer than pressure, as was revealed by the change in the diffusion coefficient kg – l at activation energies in the range of 10–40 kJ/mol. The rate of hydrogenation reactions Rhyd 3 (mol/(mreact s)) is a function of the kinetic rate kr 2 (mol/(ms s)), which is in turn determined by the con centration of hydrogen and the concentrations of reactants in biooil. In addition, Rhyd is a function of the accessible surface area of the catalyst as, divided by 3 the reactor volume (ms2 mreact ). In the case presented in Fig. 6, it is assumed that only a subsurface layer of a catalyst particle participates in the reaction, the acces sible surface area of the catalyst being proportional to the surface area of the catalyst particle as ∼ 6/dpar, where dpar is the diameter of the catalyst particle. The rate of polymerization Rpol of reactants in bio oil depends mainly on their concentration coil and temperature. The polymerization reaction order for these reactants is > 1 (up to 2 or 3). It was shown that hydrogenation reactions proceed at temperatures as low as 150°C, their products being less susceptible to polymerization. The bold line in Fig. 6 represents the temperature dependence of the effective conversion rate for biooil components when both batch reactors and flow sys tems are used. Three temperature intervals should be emphasized: (1) from room temperature to TM; (2) between TM and Tpol; and (3) above Tpol. Within the first temperature range, the initial conversion rate of reactants is limited by hydrogenation reactions, their rate also being rather slow due to low temperatures. The conversion rate for hydrogenation reactions rises between TM and Tpol, but the process is limited by the mass transfer rate within this temperature range. At
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Conversion rate, mol/min
2. Hydrogenation limited by mass transfer plus polymerization 3. Hydrogenation limited by mass transfer plus intensive polymerization
1. Hydrogenation
Hydrogenation (Rhyd) Rhyd = krap ∝ kr • 6/dpar ∝ f(T, cH2)
Polymerization(Rpol) Rpol = kpol ∝ f(T, coil)
Gas–liquid mass transfer of Н2 (reactor, T, cH2)
Tpol TM Temperature, °С Fig. 6. Temperature dependences of the conversion rates of biooil components (mol/min) by hydrogenation (Rhyd) and poly merization (Rpol) reactions, accounting for the effect of hydrogen masstransfer from gas into liquid. The bold line marks the effi cient conversion rate. TM and Tpol are critical temperatures at the boundaries of the corresponding regions: the hydrogenation reaction (T < TM); hydrogenation limited by mass transfer (TM < T < Tpol); the predominance of polymerization reactions (T > Tpol) [88].
temperatures slightly below Tpol, the rate of polymer ization increases, being limited by the mass transfer process as well. If the process is conducted in an auto clave, the mass transfer rate can be enhanced by increasing the stirring rate. At temperatures above Tpol, polymerization reactions predominate, which is cer tainly undesirable. Considering the above, we may assume that the most efficient hydrodeoxygenation method is a grad ual rise in temperature during biooil hydrotreatment. For static reactors, this approach can be implemented just by raising the reactor temperature. If a flow regime is attained, it is advisable to use a cascade connection of several reactors working at different temperatures. Since hydrogenation reaction products are less sus ceptible to polymerization, it should be remembered that with a very rapid rise in temperature, the third regime of the process (intense repolymerization) is achieved earlier than sufficient conversion by the hydrogenation reactions (regions 1 and 2), and this leads to considerable polymerization and finally to coking. It was stressed by the authors of [88] that the inner surface of catalyst particles is blocked during polymerization reactions, with only the outer surfaces of particles remaining active. In sum, the most important factor restraining the intensification of biooil hydrotreatment is its suscep tibility to repolymerization at elevated temperatures. It is therefore necessary to minimize polymerization by simultaneously increasing the rates of hydrogena tion, hydrodeoxygenation, and mass transfer. In order to attain the stated objectives, we must: 1. raise the gas–liquid mass transfer rate, by, e.g., increasing the contact area of gas and liquid phases
(al), diffusion coefficient kg – l, and/or the value of Δ cH 2 by increasing the partial pressure of H2. As was mentioned above, these requirements can be met by using a certain reactor geometry, small catalyst parti cles, and solvents (preferably nonpolar) in which hydrogen is more soluble than in biooil itself; 2. accelerate hydrogenation and hydrodeoxygen ation reactions by increasing the effective surface of a catalyst, reducing the diameter of catalyst particles, and raising the concentration of hydrogen dissolved in the liquid phase washing catalyst particles. The effi ciency of targeted chemical transformations can unquestionably be enhanced by increasing the contact (reaction) time until we attain degrees of deoxygen ation and hydrogenation at which biooil conversion products are not susceptible to polymerization upon a further rise in temperature; 3. slow the rate of polymerization by using reduced temperatures and solvents stable with respect to poly merization processes. CONCLUSIONS Today’s main trends in processing biomass into fuel products are well established, with a key place in this field being occupied by fast pyrolysis with the forma tion of biooil. The main approaches to subsequent biooil processing are fractionation; the separation of valuable chemical products; gasification aimed at obtaining synthesis gas; and hydrotreatment, includ ing hydrodeoxygenation and hydrogenation. The main objective of biooil hydrotreatment is to reduce the content of oxygen in it, since a high oxygen con tent provides such biooil properties as high viscosity, CATALYSIS IN INDUSTRY
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nonvolatility, aggressiveness, immiscibility with fossil fuels, instability, and susceptibility to polymerization. The most promising approach to biooil upgrading seems to be twostep hydrotreatment at elevated hydrogen pressures (5–10 MPa) with the partial deoxy genation and hydrogenation of biooil components occurring in the first stage at 150–250°C, which in turn increases the thermal stability of partially deoxy genated products upon further hydrotreatment at higher temperatures of 300–380°C. It is then advis able to employ the approaches described at the end of the above section on the effect of biooil properties on the effectiveness of its hydrotreatment. Due to the specific properties of biooil, particu larly its high water content and acidity, the additional requirement of stability under the conditions of hydrothermal treatment in acidic media (pH = 3–4) is needed for biooil hydrotreatment catalysts. It has been proved definitely that the conventional catalysts for oil hydrotreating (sulfided Ni–Mo/ γAl2O3 and Co–Mo/γAl2O3) are not suitable for biooil hydrotreatment, due to the low sulfur content in the initial biooil. Adding sulfur compounds to it does not solve the problem, since they pollute the hydrodeoxygenation products with sulfurcontaining components. Catalysts based on noble metals (especially Rh, Pt, and Ru catalysts deposited on carbon supports) are highly efficient in biooil hydrotreatment. These types of catalysts are free of the shortcomings of sulfided cat alysts for desulfurization. Problems arise with their regeneration, however, and they are too expensive to use in technologies for producing biofuels from biooil. The search for inexpensive hydrodeoxygenation catalysts free of the shortcomings of sulfided systems is now oriented toward nickel systems. Nickel catalysts are prone to rapid coking at elevated temperatures, but conducting the process at reduced temperatures (150–250°C) leads to nickel phase dissolution in bio oil processing products, as the oxidized nickel on the catalyst surface is not capable of reduction at such low temperatures. On one hand, researchers face the prob lem of modifying nickel catalysts with various compo nents in order to lower the reduction temperature for nickelcontaining phase to the targeted reaction tem perature. On the other hand, we must stabilize the active nickelcontaining component on a support by using specific approaches to catalyst preparation and/or stabilizing additives to increase catalyst ther mal stability in general. In light of the specific features of biooil, we should be oriented toward hydrotreat ment in which catalysts of any type are rapidly deacti vated due to coking. Such hydrodeoxygenation cata lysts should be thermally stable in order to be multiply regenerated without loss of activity. An additional requirement for biooil hydrodeoxygenation catalysts is stability of their supports under hydrothermal treat ment in acidic media. The main approach to meeting CATALYSIS IN INDUSTRY
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the indicated requirements is to modify nickel systems with elements such as Cu, Mo, La, Co, Fe, B, P. Researchers are oriented toward using stable catalyst supports (C, ZrO2, TiO2, SiO2) or completely elimi nating them. The existing trends in the development of biooil hydrotreatment catalysts allow us to conclude that the number of new catalysts for hydrodeoxygen ation of biomass processing products will grow with an emphasis on nonsulfided catalytic systems. If the development of nickelcontaining biooil hydrodeox ygenation catalysts is successful, more rapid develop ment should be expected in such fields of bioenerget ics as producing liquid motor fuels from wood, edible and nonedible oil crops, and microalgae lipids. ACKNOWLEDGMENTS This work was supported by the federal target pro gram Research and Development of Russia’s Scien tific and Technological Complex, 2007–2012 (state contract no. 16.516.11.6049) REFERENCES 1. Varfolomeev, S.D., Efremenko, E.N., and Krylova, L.P., Usp. Khim., 2010, vol. 79, no. 6, p. 544. 2. Moiseev, I.I., Teor. Prikl. Khim., 2010, vol. 46, no. 6, p. 360. 3. Bridgwater, A.V., Meier, D., and Radlein, D., Org. Geochem., 1999, vol. 30, no. 12, p. 1479. 4. Bridgwater, A.V. and Peacocke, G.V.C., Renew. Sust. Energy Rev., 2000, vol. 4, no. 1, p. 1. 5. Zhang, Q., Chang, J., Wang, T., and Xu, Y., Energy Convers. Manag., 2007, vol. 48, p. 87. 6. Geletukha, G.G. and Zheleznaya, T.A., Ekotekhnol. Resursosber., 2000, no. 2, p. 3. 7. Azarov, V.I., Burov, A.V., and Obolenskaya, A.V., Khimiya drevesiny i sinteticheskikh polimerov (Chemis try of Wood and Synthetic Polymers), St. Petersburg: SPbLTA, 1999. 8. Igathinathane, C., Tumuluru, J.S., Sokhansanj, S., Bi, X., Lim, C.J., Melin, S., and Mohammad, E., Bioresour. Technol., 2010, vol. 101, p. 6528. 9. Mercader, F. de Miguel, Groenevel., M.J., Kersten, S.R.A., Venderbosch, R.H., and Hogendoorn, J.A., Fuel, 2010, vol. 89, p. 2829. 10. Zhang, S., Yan, Y., Li, T., and Ren, Z., Bioresour. Tech nol., 2005, vol. 96, no. 5, p. 545. 11. Elliott, D.C., Energy Fuels, 2007, vol. 21, no. 3, p. 1792. 12. Bejblova, M., Zamostny, P., Cerveny, L., and Cejka, J., Appl. Catal., A, 2005, vol. 296, no. 2, p. 169. 13. Dundich, V.O. and Yakovlev, V.A., Khim. Interes. Usto ich. Razvit., 2009, vol. 17, no. 5, p. 527. 14. Prochazkova, D., Zamostny, P., Bejblova, M., Cerveny, L., and Cejka, J., Appl. Catal., A, 2007, vol. 332, no. 1, p. 56. 15. Lunzer, S. and Kramer, R., Stud. Surf. Sci. Catal., 2000, vol. 130, p. 2303. 16. Wildschut, J., MelianCabrera, I., and Heeres, H.J., Appl. Catal., B, 2010, vol. 99, nos. 1–2, p. 298.
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