Hydro deoxygenation catalyst, a fixed bed tandem catalytic reactor, a method for preparing hydrogen and a method for preparing biofuel from biomass

Abstract

The present invention relates to processes for the preparation of biofuel from biomass by fast hydropyrolysis or fast pyrolysis, using hydrogen generated by sorption enhanced steam reforming. The present invention also relates to fixed bed tandem catalytic-upgrading processes, and reactors and hydrodeoxygenation (HDO) catalysts useful in those processes.

Claims

1. A hydrodeoxygenation (HDO) catalyst which is Ru—MoFeP supported on Al.sub.2O.sub.3.

2. A fixed bed tandem catalytic-upgrading reactor suitable for upgrading bio-vapour or bio-oil from fast hydropyrolysis or fast pyrolysis into biofuel, wherein: the fixed bed comprises an upstream portion and a downstream portion, the upstream portion comprises an C—C coupling catalyst, and the downstream portion comprises the hydrodeoxygenation (HDO) catalyst according to claim 1.

3. The fixed bed tandem catalytic-upgrading reactor according to claim 2, wherein the C—C coupling catalyst is an aldol condensation and ketonization catalyst.

4. The fixed bed tandem catalytic-upgrading reactor according to claim 3, wherein the C—C coupling catalyst comprises TiO.sub.2 or TiO.sub.2 doped with Au, Ag, Cu, Pd or Ru.

5. The fixed bed tandem catalytic-upgrading reactor according to claim 4, wherein the C—C-coupling catalyst comprises TiO.sub.2 or TiO.sub.2 doped with Ru or Au.

6. A method for preparing biofuel from biomass, the method comprising: (a) fast hydropyrolysis or fast pyrolysis of biomass to provide bio-vapour and/or bio-oil; and (b) upgrading of the bio-vapour and/or bio-oil from step (a) in a fixed bed tandem catalytic-upgrading reactor to provide (i) the biofuel, and (ii) a stream comprising C.sub.1-C.sub.4 hydrocarbons, CO and CO.sub.2, wherein the fixed bed tandem catalytic-upgrading reactor comprises an upstream portion and a downstream portion, the upstream portion comprises an C—C coupling catalyst, and the downstream portion comprises the hydrodeoxygenation (HDO) catalyst according to claim 1.

7. A method for preparing biofuel from biomass, the method comprising: (a) fast hydropyrolysis or fast pyrolysis of biomass to provide bio-vapour and/or bio-oil; (b) upgrading of the bio-vapour and/or bio-oil from step (a) in a fixed bed tandem catalytic-upgrading reactor to provide (i) the biofuel, and (ii) a stream comprising C.sub.1-C.sub.4 hydrocarbons, CO and CO, and (c) steam reforming the stream comprising C.sub.1-C.sub.4 hydrocarbons, CO and CO.sub.2 from step (b) in the presence of a sorbent suitable for CO.sub.2 capture, to provide hydrogen, wherein the fixed bed tandem catalytic-upgrading reactor comprises an upstream portion and a downstream portion, the upstream portion comprises an C—C coupling catalyst, and the downstream portion comprises a hydrodeoxygenation (HDO) catalyst.

8. The method according to claim 7, wherein fast hydropyrolysis is used in step (a) and which method further comprises: (d) using the H.sub.2 produced in step (c) for fast hydropyrolysis in step (a).

9. The method according to claim 7, which further comprises: (e) introducing H.sub.2 from step (c) into the catalytic-upgrading reactor in step (b).

10. The method according to claim 7, wherein the steam reforming is pressure or temperature swing sorption enhanced steam reforming.

11. The method according to claim 10, wherein the steam reforming is pressure swing sorption enhanced steam reforming (PS SESR), whereby (i) steam reforming and CO.sub.2 sorption occurs in a first reactor and sorbent regeneration occurs in a second reactor, and then (ii) sorbent regeneration occurs in the first reactor and steam reforming and CO.sub.2 sorption occurs in the second reactor.

12. The method according to claim 7, wherein the steam reforming is based on carbonate looping by a circulating fluidized-bed (CFB) reactor, wherein steam reforming and CO.sub.2 sorption occurs in a first reactor and sorbent regeneration occurs in a second reactor, and the sorbent material is circulated between the first reactor and the second reactor.

13. The method according to claim 7, wherein the sorbent comprises CaO.

14. The method according to claim 13, wherein the sorbent is derived from limestone or dolomite.

15. The method according to claim 7, wherein the steam reforming uses a catalyst which is Ni, Co, Ni—Co or noble metal (M: Pt, Pd, Ru, Rh) promoted above catalyst.

16. The method according to claim 15, wherein the steam reforming uses a catalyst which is Pd/Ni—Co.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

(1) FIG. 1A is a schematic view of the synergistic integration of biofuel (liquid hydrocarbon) synthesis by fast hydropyrolysis of biomass using pressure swing sorption enhanced steam reforming (PS SESR) to produce H.sub.2.

(2) FIG. 1B is a schematic view of the synergistic integration of biofuel (liquid hydrocarbon) synthesis by fast hydropyrolysis of biomass using carbonate looping by a circulating fluidized-bed (CFB) reactor to produce H.sub.2.

(3) FIG. 2A is a schematic view of the synergistic integration of biofuel (liquid hydrocarbon) synthesis with high yield and purity by fast hydropyrolysis biomass followed by upgrading using a fixed bed tandem catalytic-upgrading reactor, wherein C—C coupling and hydrodeoxygenation reactions occurs in two stages and in a single reactor.

(4) FIG. 2B is a schematic view of the synergistic integration of biofuel (liquid hydrocarbon) synthesis with high yield and purity by fast hydropyrolysis biomass followed by upgrading using a fixed bed tandem catalytic-upgrading reactor, wherein C—C coupling and hydrodeoxygenation reactions occurs in two stages and in a single reactor.

(5) FIG. 3 is a schematic view of a specific process of the invention described in Example 1.

(6) FIG. 4 represents total carbon-based yield obtained based on organic and gaseous product derived from HDO only as compared to tandem catalytic systems, when upgrading simulated bio oil to jet-fuel range aromatics in Example 3. Total pressure 20 bar, H.sub.2 partial pressure 17 bar, Temperature of reaction: 400° C., reaction time 6 h, WHSV: 0.49/h, weight of catalyst: HDO (Ru/MoFeP—Al.sub.2O.sub.3)=20 g, XTiO.sub.2=20 g

(7) FIG. 5 represents detailed product distribution in the organic and gaseous product stream derived from HDO only as compared to tandem catalytic systems, when upgrading simulated bio oil to jet-fuel range aromatics in Example 3. Total pressure 20 bar, H.sub.2 partial pressure 17 bar, Temperature of reaction: 400° C.

DETAILED DESCRIPTION OF THE INVENTION

(8) The present invention is concerned with the preparation of biofuel from biomass via fast hydropyrolysis or fast pyrolysis of the biomass to produce bio-vapour and/or bio-oil, followed by subsequent upgrading of the bio-vapour and/or bio-oil to produce biofuel.

(9) The overall process, which is described in further detail below, is illustrated by the specific process depicted in FIGS. 1A, 1B, 2A and 2B. In these figures, a conventional fast hydropyrolysis reactor 2 is an apparatus into which biomass 1 is fed, and a mechanism to supply heat 11 and H.sub.2 12 for the pyrolysis is provided. The gases exiting the fast hydropyrolysis reactor 2 are generally sent to a cyclone (not shown) where solid char 9 are separated. This char 9 may be burned to provide heat for pyrolysis and drying. Next, bio-vapour 3 is sent to a tandem catalytic upgrading reactor 4. The biofuel (liquid hydrocarbons) 5 are separated and light gaseous products 6 will go to pressure swing sorption enhanced steam reforming (PS SESR) reactors 7 and 8 in FIGS. 1A and 2A or to a carbonate looping by a circulating fluidized-bed (CFB) reactor 7 and 8 in FIGS. 1B and 2B to produce pure H.sub.2, 12. The CO.sub.2 captured in the PS SESR reactors will be released 13. The regeneration of CO.sub.2 is an endothermic process, requiring heat 11.

(10) The biofuel prepared using the present invention comprises liquid hydrocarbons. The liquid hydrocarbons typically contain five or more carbon atoms (C.sub.5+ hydrocarbons). Preferably the biofuel prepared using the present invention comprises a high proportion of C.sub.8-C.sub.13 hydrocarbons, such as C.sub.9 hydrocarbons. Preferably the biofuel prepared using the present invention also contains a high proportion of aromatics. These longer hydrocarbons and aromatics are particularly useful in the preparation of jet-fuel.

(11) As used herein, the term “biomass” means carbon-containing organic matter, generally derived from plants. Some examples include switchgrass, poplar tree, sugarcane, corn, tree barks, aquatic material including algae, plankton. However, the precise nature of the biomass is not believed to be an important aspect of the invention. Typically, the biomass is lignocellulosic biomass. Preferably, the biomass is woody or agricultural biomass. Preferably the biomass is woody biomass. Typically the biomass is processed, for example by grinding or shredding, such that it is suitable for introduction into the reactor where fast hydropyrolysis or fast pyrolysis occurs.

(12) Any conventional fast hydropyrolysis or fast pyrolysis reactor can be used for the fast hydropyrolysis or fast pyrolysis steps. Fast pyrolysis involves heating the biomass in the reactor in the absence of air (particularly O.sub.2). Fast hydropyrolysis involves heating the biomass in the reactor in the presence of hydrogen and the absence of other air (particularly O.sub.2). Otherwise, the conditions for fast hydropyrolysis and fast pyrolysis are generally the same. Thus, the processes typically involve heating the biomass to 300° C. to 600° C., preferably 400 to 500° C., for example about 400° C. or about 500° C. The heating is rapid, hence it being “fast” in contrast to slow pyrolysis which mainly produces char. Typically, processes occurs at a pressure of 5 to 50 bar, preferably 5 to 20 bar. Typically heat is supplied at a rate of 1500 to 2500 J/g of biomass, for example about 2000 Jg/biomass. Typically, the flux is above 50 W/cm.sup.2.

(13) Fast hydropyrolysis and fast pyrolysis convert the biomass into bio-vapour and/or bio-oil. Bio-vapour and bio-oil contain the same components, but referred to as “bio-oil” when in liquid phase at lower temperatures and “bio-vapour” when gaseous phase at higher temperatures. Bio-vapour and bio-oil contain a complex mixture of compounds (e.g. acetic acid, acetol, furfural, phenol, guaiacol, eugenol etc), which will depend to some extent on the exact starting biomass material. In addition, solid char is generally produced during fast hydropyrolysis and fast pyrolysis, and this solid char is typically separated from the bio-vapour and/or bio-oil using routine separation techniques, such as a cyclone. The solid char may be burned to provide heat for other steps in the process of the invention, including the fast hydropyrolysis or fast pyrolysis steps.

(14) The fast hydropyrolysis or fast pyrolysis steps may be carried out in the presence of an HDO catalyst. In that case HDO catalyst is present in the fast hydropyrolysis or fast pyrolysis reactor during pyrolysis. For example, if heat for fast hydropyrolysis or fast pyrolysis is provided through a fluidized bed, then catalyst particles can either be mixed with the material of the fluidized bed or supported on the particles being fluidized. An example would be sand used as a circulating fluidized material to supply heat for fast hydropyrolysis or fast pyrolysis. In such a case, the HDO catalyst may either be mixed with the sand or supported on the sand particles. Typically, however, the fast hydropyrolysis or fast pyrolysis step is carried out in the absence of an HDO catalyst, since upgrading is carried out ex-situ in a subsequent step.

(15) The bio-vapour and/or bio-oil formed from the fast hydropyrolysis or fast pyrolysis step contains a complex mixture of materials, including a significant proportion of light oxygenates. The bio-vapour and/or bio-oil is sent next to a fixed bed tandem catalytic-upgrading reactor. Typically, the bio-vapour and/or bio-oil is cooled (for example in a cooler) prior to entering the fixed bed tandem catalytic-upgrading reactor. That is because the operating temperature of the fixed bed tandem catalytic-upgrading reactor (300 to 400° C.) is may be lower than that of the fast hydropyrolysis reactor (300 to 600° C.). Alternatively, if the bio-vapour and/or bio-oil has been cooled to below the operating temperature of the fixed bed tandem catalytic-upgrading reactor, then the bio-vapour and/or bio-oil can be heated prior to entering the fixed bed tandem catalytic-upgrading reactor.

(16) The fixed bed tandem catalytic-upgrading reactor has a fixed bed which comprises an upstream portion and a downstream portion. The upstream portion comprises a C—C coupling catalyst, and the downstream portion comprises a hydrodeoxygenation (HDO) catalyst. This arrangement is depicted in 4 of FIG. 2. For the avoidance of doubt, the upstream portion comprising a C—C coupling catalyst and downstream portion comprising a hydrodeoxygenation (HDO) catalyst are present in a single fixed bed reactor, and are not present in separate reactors.

(17) The preferred operating temperature of the fixed bed tandem catalytic-upgrading reactor will depend upon the specific C—C coupling and HDO catalysts used, but is typically 300 to 400° C. The preferred operating pressure of the fixed bed tandem catalytic-upgrading reactor will depend upon the specific C—C coupling and HDO catalysts used, but is typically 5 to 20 bar. Given that the upstream and downstream portions are present in a single reactor, the reaction conditions, and in particular the temperature and pressure, are generally the same in the upstream portion and the downstream portion. This is advantageous, as there is no need to change the temperature and/or pressure between the C—C coupling catalyst (upstream) portion and the HDO catalyst (downstream) portion, which increases the overall efficiency of the processes. Inefficient changes in temperature and/or pressure are generally necessary when separate reactors are used for the C—C coupling catalyst and the HDO catalyst.

(18) Any conventional C—C coupling catalyst can be used in the present invention. A C—C coupling catalyst is a catalyst which catalyses reactions which form C—C bonds between hydrocarbon compounds, and thereby increase the number of carbon atoms in the resulting hydrocarbon products.

(19) Typically the C—C coupling catalyst is an aldol condensation and ketonization catalyst (sometimes known as an “aldol catalyst”), which catalysts a reaction in which C—C bonds are formed by (e.g.) aldol condensations. Preferably the aldol condensation and ketonization catalyst comprises TiO.sub.2 or TiO.sub.2 doped with Au, Ag, Cu, Pd or Ru. Preferably the aldol condensation and ketonization comprises TiO.sub.2 doped with Au, Ag, Cu, Pd or Ru, more preferably TiO.sub.2 doped with Au or Ru. Typically the Au, Ag, Cu, Pd or Ru is present in an amount of 0.1 to 0.3 wt %, for example about 0.2 wt %.

(20) Typically C—C coupling catalyst as described above is in the form of pellets.

(21) Any conventional HDO catalyst may be used in the present invention (both in the fixed bed tandem catalytic-upgrading reactor and, when present, in the fast hydropyrolysis or hydropyrolysis or fast pyrolysis reactor). An HDO catalyst catalyses reactions in which oxygen is removed from oxygen-containing compounds by reaction of the oxygen-containing compounds with H.sub.2 to form water. For example, in the present invention oxygen-containing hydrocarbon products of the C—C coupling reactions react with H.sub.2, thereby removing oxygen from those hydrocarbons and generating water.

(22) Typically the HDO catalyst comprises (a) Fe—S, Ni—Co or Co—Mo supported on a support, or (b) M.sup.1—MoM.sup.2P supported on support, wherein M.sup.1 and M.sup.2 represent transition metals. Typically the HDO catalyst is in the form of pellets. The M.sup.1—MoM.sup.2P supported on support is preferred as an HDO catalyst.

(23) M.sup.1 and M.sup.2 represent different transition metals.

(24) M.sup.1 generally acts as a promotor in the HDO catalyst. A promotor is component which has little or no catalytic effect itself, but improves the performance of the catalyst in which it is present. Thus, in the present case, M.sup.1 typically improves the performance of the HDO catalyst, without generally catalysing the HDO reactions itself. Typically, M.sup.1 represents Rh, Ru, Pt, Pd, Ni, Co or Cu, and preferably represents Ru.

(25) Typically, M.sup.2 represents Ni, Co, Fe or Cu, and preferably represents Fe.

(26) The molar ratio of Mo:M.sup.2:P is 0.8-1.2:0.8-1.2:0.8-1.2, typically about 1:1:1. M.sup.1 is typically present in an amount of about 0.05 to about 0.1 wt %, based on the total weight of M.sup.1—MoM.sup.2P. The Mo, M.sup.2 and P are typically atomically dispersed on the surface of the acid support. The M.sup.1 typically forms a nano-layer (or single atom layer) on the surface of the Mo, M.sup.2, P and acid support.

(27) Any conventional support suitable for catalysts can be used, and will generally be one with a high surface area. Typically the support is an acidic oxide such as Al.sub.2O.sub.3, TiO.sub.2, ZrO.sub.2 or CeO.sub.2, a carbon material such as activated carbon, mesoporous carbon or carbon nanomaterials, or SiO.sub.2. Preferably the support is Al.sub.2O.sub.3, most preferably γ-Al.sub.2O.sub.3.

(28) Thus, a preferred HDO catalyst is Ru—MoFeP supported on Al.sub.2O.sub.3. Typically the Al.sub.2O.sub.3 is γ-Al.sub.2O.sub.3. Preferably the ratio or Mo:Fe:P is preferably 1:1:1. Ru is preferably present in an amount of about 0.05 to about 0.5 wt %, more preferably about 0.05 to about 0.1 wt %, for example about 0.1 wt %, based on the total weight of Ru, Mo, Fe and P.

(29) M.sup.1—MoM.sup.2P catalysts can be prepared by a modified Pichini method, for example that described in Example 2. In summary, the method typically involves a sequential wetness impregnation technique. For example, the Mo, M.sup.2 and P components are first added to a solution (typically a citric acid solution) in the desired amounts, this solution is then used to impregnate an acid support (such as Al.sub.2O.sub.3). Following impregnation, the coated support is dried and calcined. Next, a solution containing the desired amount of M.sup.1 is used to impregnate the coated acidic support, followed by a second drying and calcination step. Finally, the active catalyst is obtained by heating (typically in the presence of H.sub.2 and N.sub.2) to convert the M.sup.1, Mo, M.sup.2 and P components into elemental form.

(30) H.sub.2 is typically introduced into the fixed bed tandem catalytic-upgrading reactor, in order to promote the HDO reactions. Generally, any such H.sub.2 added to the fixed bed tandem catalytic-upgrading reactor will have been generated in the sorption enhanced steam reforming step described further below.

(31) The bio-vapour and/or bio-oil entering the fixed bed tandem catalytic-upgrading reactor first comes into contact with the C-Caldol catalyst in the upstream portion of the fixed bed. The C—C coupling catalyst promotes reactions (e.g. aldol condensations) which convert light oxygenates into heavier compounds containing higher numbers of carbon atoms. Next, the materials generated following the bio-vapour and/or bio-oil contacting the C—C coupling catalyst will come into contact with the HDO catalyst in the downstream portion of the fixed bed. The HDO catalysts promotes hydrodeoxygenation reactions, thereby reducing the oxygen content of the resulting effluent that leaves the fixed bed tandem catalytic-upgrading reactor.

(32) Upgrading using the fixed bed tandem catalytic-upgrading reactor provides (i) the biofuel, and (ii) a stream comprising C.sub.1-C.sub.4 hydrocarbons, CO and CO.sub.2. Typically, (i) and (ii) exit the fixed bed tandem catalytic-upgrading reactor together as a combined mixture or effluent containing both (i) and (ii). The effluent is typically condensed and (i) is separated from (ii). Typically, (i) is a relatively low oxygen, high energy density bio-fuel product, which may be directly used in many applications without further upgrading or processing. Stream (ii) comprising C.sub.1-C.sub.4 hydrocarbons, CO and CO.sub.2 from the upgrading step may further comprise oxygenates which were not converted into (i) biofuel.

(33) The stream comprising C.sub.1-C.sub.4 hydrocarbons, CO and CO.sub.2 from the upgrading step is next subjected to sorption enhanced steam reforming (SESR). SESR is an integrated process involving steam reforming of a stream comprising C.sub.1-C.sub.4 hydrocarbons, CO and CO.sub.2 in the presence of a sorbent suitable for CO.sub.2 capture, thereby to produce H.sub.2. The SESR reactor contains the catalyst required for the steam reforming process together with a sorbent suitable for CO.sub.2 capture for the in-situ removal of carbon dioxide from the gaseous phase. The steam reforming [including water gas shift (WGS)] and CO.sub.2 capture reactions are thus conducted simultaneously in one single reactor.

(34) The steam reforming aspect of this step uses a steam reforming catalyst, such as Ni, Co or Ni—Co, or noble metal (i.e. Pt, Pd, Ru, Rh) promoted versions of Ni, Co or Ni/Co. Pd promoted Ni—Co (i.e. Pd/Ni—Co) is particularly preferred. Ni catalysts are commonly used in steam reforming processes because they have high activity and selectivity towards hydrogen products. However, Ni catalysts do not offer particularly high resistance to the deactivation caused by coke deposition on nickel particles.sup.24. In previous work from the inventors, a Pd/Ni—Co catalyst derived from a hydrotalcite-like material (HT) has been demonstrated to be an effective catalyst.sup.25,26,27,22 with high activity and selectivity towards hydrogen products and high resistance to deactivation caused by coke deposition. In particular, this Pd/Ni—Co catalyst is a highly active catalyst per weight as well as volume, has proper redox properties and spatial confinement against sintering, and is superior to the commercial reforming catalysts. The Pd/Ni—Co catalyst can therefore be advantageously used in the SESR techniques of the present invention.

(35) Steam reforming involves the reaction of C.sub.1-C.sub.4 hydrocarbons and CO with water to provide hydrogen and CO.sub.2. The reactions involved can be illustrated for methane as follows:
CH.sub.4+H.sub.2O CO+3H.sub.2[steam reforming reaction]ΔH.sub.r.sup.0=+184 kJ mol.sup.−1
CO+H.sub.2OCO.sub.2+H.sub.2[water gas shift reaction]ΔH.sub.r.sup.0=−41 kJ mol.sup.−1

(36) Both of the above reactions are reversible, and so the reactions can be driven towards H.sub.2 production by removal of CO.sub.2. Removal of CO.sub.2 is achieved by conducting the steam reforming steps in the presence of a sorbent suitable for CO.sub.2 capture. The sorbent reacts with the CO.sub.2, generally as soon as it is formed, thereby driving the equilibrium towards H.sub.2 production.

(37) Any sorbent that is suitable for CO.sub.2 capture can be used, but generally CaO-based sorbents are preferred. Natural limestone (primarily CaCO.sub.3) and dolomite (primarily CaCO.sub.3.MgCO.sub.3) based sorbents being particularly preferred due to their low cost and ready availability (despite suffering from a decay in their CO.sub.2 capture capacity after several cycles of carbonation/regeneration.sup.23). These natural materials can be converted into their oxides by heating, thereby to provide the sorbent.

(38) For example, a sorbent material can be prepared from limestone (CaCO.sub.3) by heating it to provide CaO (and CO.sub.2). The CaO sorbent can then react with CO.sub.2 to reform the CaCO.sub.3,
CO.sub.2+CaOCaCO.sub.3[steam reforming reaction]ΔH.sub.r.sup.0=−178 kJ mol.sup.−1

(39) thereby removing the CO.sub.2 from the atmosphere. This reaction occurs at low CO.sub.2 partial pressures and at moderate temperatures and has fast kinetics and good adsorption capacities. When desired, the CaO sorbent can be regenerated from the thus-formed CaCO.sub.3 by heating, with the relatively pure stream of CO.sub.2 produced as a by-product being suitable for other uses or sequestration.

(40) Given the reversibility of the sorbent reaction, it is particularly preferred to use temperature or pressure swing sorption enhanced steam reforming (PS SESR), with PS SESR preferred. A typical arrangement for PS SESR involves a first reactor and a second reactor (as depicted in FIGS. 1A and 2A). Initially, steam reforming and CO.sub.2 sorption occurs in the first reactor, whilst sorbent regeneration occurs in a second reactor. At an appropriate point (e.g. when the sorbent in the first reactor has all absorbed CO.sub.2 and/or sorbent regeneration is complete in the second reactor), the reactors are switched so that sorbent regeneration instead starts to occur in the first reactor and steam reforming and CO.sub.2 sorption occurs in the second reactor. The process of switching between reactors can be continued. In this way, a continual stream of H.sub.2 can be produced without needing to interrupt the process (i.e. PS SESR is a continuous process).

(41) The H.sub.2 production can alternatively be done based on carbonate looping by a circulating fluidized-bed (CFB) reactor (as depicted in FIGS. 1B and 2B), where one fluidized-bed acts as a reformer where steam reforming, water gas shift and CO.sub.2 removal by the solid sorbent occurs simultaneously and the other release CO.sub.2 from the solid sorbent (thereby regenerating the sorbent). The solid sorbent circulates between the two reactors.

(42) The use of SESR in the manner described above has a number of significant advantages compared to, for example, standard steam reforming. First, the amount of hydrogen produced increases, due to the reversible reactions being driven towards hydrogen production. Second, the hydrogen that is produced contains very little residual CO.sub.2, and so generally can be fed directly back into the fast hydropyrolysis process without further purification. Finally, the use of carbon from the starting biomass to generate hydrogen increases the overall energy efficiency of the process. The use of PS SESR has the additional advantage that the process does not need to be interrupted for sorbent regeneration, and is therefore highly efficient. Further, the resulting H.sub.2 is at relatively high pressure and so can generally be introduced directly into the fast hydropyrolysis process without the need for re-pressurisation.

(43) The H.sub.2 produced by SESR is then typically introduced into the fast hydropyrolysis reactor, though some H.sub.2 may also be introduced into the fixed bed tandem catalytic-upgrading reactor. The overall process is thus circular and continuous, and consequently is generally highly efficient.

EXAMPLES

(44) The following are Examples that illustrate the present invention. However, these Examples are in no way intended to limit the scope of the invention.

Example 1

(45) Integration of H.sub.2 production from CO and C.sub.1-C.sub.4 from an HDO reactor with FHP is presented is represented in FIG. 3. The pure hydrogen is produced by pressure swing sorption enhanced steam methane reforming (SESR) and is fed to the FHP reactor. This SESR process can produce relatively pure hydrogen and helps increase hydrogen production versus regular steam reforming, by using a solid sorbent material that removes carbon dioxide from the reactor, shifting the water gas equilibrium to favour hydrogen production. Additionally, the charcoal by-product of the fast-pyrolysis process can be combusted and used to provide energy to the plant to reduce energy costs.

(46) Aspen Plus simulation software was used for process modelling. Two parameters were considered in evaluating the bio-fuel production process, carbon efficiency and energy efficiency. The total carbon efficiency of the plant was determined by calculating the amount of the carbon in the feed stream and the amount of carbon in the bio-fuel product. The wood feed has 10750 kg of carbon and the bio-oil (C.sub.4+) has 5801 kg of carbon. This gives a carbon efficiency of 54%; if charcoal production is included, the carbon efficiency increases to 73.7%. The biofuel has an ethanol gallon equivalent (ege) of 163.4 ege/ton biomass. The energy efficiency of the process was calculated by finding the total amount of energy produced by the process. Biofuel is assumed to have an energy content of 42.1 MJ/kg whereas biomass only has an energy content of 17.1 MJ/kg. Using the flow rate of biofuel produced, the total energy produced was estimated to be 2.40 million GJ. The total energy used in the plant was calculated as the total energy added to the plant plus the energy content of the biomass. This gave an energy efficiency of 74.9% for the process.

Example 2—Preparation of Ru—MoFeP Supported on Al.SUB.2.O.SUB.3

(47) The MoFeP active phase on supported alumina spheres was prepared by a modified Pichini method 51,52 using sequential wetness impregnation method. 0.4M of aqueous citric acid, organic chelating agent, was first prepared to create acidic environment to free the metal salts from precipitating. A 1:1:1 molar concentration of Mo:Fe:P was added stepwise onto the 1M citric acid solution. In typical experiment in which 100 g of alumina sphere.sup.22 were used, the Mo, Fe and P precursor weight used was 37.2, 84.8 and 23.9 grams respectively, representing 20 wt % loading of the active metal phase.

(48) The light yellowish homogenous solution formed was impregnated sequentially on the spherical alumina support within 24 period. The formed catalyst stayed at room temperature overnight and was dried at 100° C. for 12 h. The catalyst was further calcined in air at a heating rate of 1° C./min, with a dwell time at 350° C. of 6 h. 0.1 wt % of Ru was then impregnated onto the calcined catalyst using incipient wetness impregnation method. The Ru promoted catalyst on MoFeP/Al.sub.2O.sub.3 was subsequently dried 100° C. (for 4 h) and calcined at 500° C. using 1° C./min heating rate, dwell time at final temperature is was 5 h.

(49) The active phase, Ru—MoFeP, was obtained using temperature reduction method at heating rate of 1° C./min at 250 and 700° C. to react all phosphorus and reduced oxides to metallic form.sup.44-46 in the presence of 75% H.sub.2 in Nitrogen.

Example 3

(50) Biofuel was produced from simulated bio-vapour using tandem catalytic upgrading, whereby high yield and purity of the biofuel was achieved by integrating C═C coupling via aldol condensation/dehydration/ring closure reactions and hydrodeoxygenation in one single reactor over dual bed catalyst system.

(51) A simulated bio vapour feed (water, acetic acid, acetol, furfural, phenol, guaiacol and eugenol) was prepared based on the product distribution observed from the pyrolysis of several biomass using pyrolysis gas chromatography mass spectrometry (PyGCMS).

(52) A fixed bed tandem catalytic-upgrading reactor according to the invention was prepared. The aldol catalysts were TiO.sub.2 pellets or 0.2 wt % X—TiO.sub.2, where X represents Au, Pd or Ru. The X—TiO.sub.2 catalysts were prepared by a standard wetness impregnation technique. Ru—FeMoP/Al.sub.2O.sub.3 was used as the HDO catalyst. The Ru—FeMoP/Al.sub.2O.sub.3HDO catalyst was applied using spherical alumina support. The active phase was impregnated sequentially on the support using citrate acid as passivating agent as described in Example 2.

(53) The simulated bio vapour feed was then passed through the fixed bed tandem catalytic-upgrading reactor in a pilot plant investigation.

(54) In FIG. 4, the overall carbon recovery for all tandem catalyst was greater than 95%. The gas phase yield declined from 17% (for HDO alone) to 9% (Au/TiO.sub.2+HDO), with corresponding increases in the liquid yield. The other tandem catalysts also provided a reduction in gas yield and increase in liquid yield, particularly with the doped TiO.sub.2 aldol catalysts. The total carbon recovered in the liquid phase was greater than 80% for the tandem catalyst, thus a higher carbon efficiency was attained.

(55) FIG. 5 illustrates the detailed hydrocarbon carbon-number yield as a function of the tandem catalysts. Clearly, the selectivity to CO+CO.sub.2 was higher for the Pd/TiO.sub.2+HDO while C.sub.3 (propene and propane) was also greater for only TiO.sub.2+HDO. Acetone yield observed was higher for Pd/TiO.sub.2+HDO. The increasing in CO.sub.2 and acetone possibly suggest that Pd/TiO.sub.2 was the most active catalyst for ketonization of acetic acid to acetone, CO.sub.2 and water. However, in considering the CO.sub.2/C.sub.0 molar ratio, which gives indication of the preferred reaction pathway, thus either combined decarboxylation (CO.sub.2 release) and ketonization (CO.sub.2 release) vs decarbonylation (CO release). The CO.sub.2/C.sub.0 ratio of 6.8, 5.6, 5.5, 3.4 and 2.7 was observed for Au/TiO.sub.2+HDO, Pd/TiO.sub.2+HDO, Ru/TiO.sub.2+HDO, TiO.sub.2 and HDO, respectively. Clearly, decarboxylation and ketonization activity was higher than decarbonylation for these catalysts. In addition, the lower amount of acetone observed for Au as compared to Pd suggest Au/TiO.sub.2 promoted much higher carbon coupling activity than Pd based catalyst. Therefore, it is expected that Au based catalyst should have higher chain growth than Pd based catalyst. This was further confirmed in the liquid phase analysis where the C.sub.10 and C.sub.11+ carbon yield was higher under Au based upstream catalyst. The polycyclic aromatic formation activity that significantly depends on the HDO catalyst was similar in all dual bed tested catalyst. The observed products once again have higher value in the usage as gasoline additive or jet-fuel blend fuel. The tandem strategy led to reduction in gas phase products while increasing the organic liquid yield with much chain growth observed for Au impregnated catalyst. The high chain growth activity of Au may be due to its mild hydrogenation activity, which is required in condensation reactions. However, Ru based upstream catalyst gave the highest degree of deoxygenation due to the phenolics hydrodeoxygenation ability for Ru/TiO.sub.2.

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