Field of the invention

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.

Background to the invention

The increasing economic development facilitating globalization has caused a remarkably increases in the energy demand. Currently, fossil fuel accounts for about 80% of the global energy consumption and 95% of the transport energy demand ¹ . The transportation sector currently relies almost exclusively on liquid hydrocarbons as its energy source.

Production of liquid fuels from biomass can help solve the problem of CO2 emission from the transportation sector because CO2 released from vehicle exhaust is captured during biomass growth.

The aviation sector is deemed as one of the difficult-to-decarbonize areas and thus research into renewable jet-fuel range hydrocarbon is great, especially the liquid transport fuel and the aviation fuel. The typical composition of commercial jet- A range (aviation) fuel is made up paraffins, isoparaffins, olefins, naphthenes and aromatics of standard vol % limitation of 20, 33, 6, 13 and 26, respectively. The hydrocarbon carbon length usually contains ca.98 % of Cx-Cn linear, branched, cyclic and aromatic backbone ² \ The functionality of each components enhances either the combustion, thermal stability, fluidity, flying easiness and safety requirements and as such, the mixture of multi-functionality is mandatory for the aviation turbine engine ³ . Most especially the safety functionality produced by aromatic backbone such as aiding the elastomeric swelling O-rings in fuel tanks ³ and thereby reducing leakage of the most volatile fraction of the fuel is of high necessity ⁴ . However, the production of aromatics from lignocellulose biomass has attracted less research focus and attention.

The molecules present in biomass are complex and strongly functionalized, resulting in a low energy density (compared to hydrocarbons). However, the functionalized complex molecules provide diverse reaction pathways to different intermediate molecules. Biomass consists typically of 40-45% cellulose, 15-30% hemicellulose and 10-25% lignin. Full utilization of all the biomass components is essential for achieving high energy efficiency and becoming highly economically competitive to fossil fuels. There are three platforms available for biomass conversion to fuels and chemicals: the gasification, pyrolysis and hydrolysis (sugar) platforms ^{5 7} .

The sugar platform involves pre-treatment of biomass and hydrolysis to sugars and polyolic platform molecules followed by conversion of the platform molecules to chemicals and fuels. The drawback of this platform is that the lignin component cannot be utilized in the aqueous phase processing of biomass, and is typically combusted to generate heat and power ⁸ . Recently an alternative catalytic pathway has been reported, in which lignocellulosic biomass is directly converted to polyols over a heterogeneous catalyst in the presence of hydrogen. The process was named one-pot catalytic conversion of raw woody biomass ⁹ ’ ¹⁰ , which is a hydrothermal liquefaction process. Subsequent deoxygenation and C-C coupling reactions can convert these polyols to liquid transportation fuels. The remarkable fact about this catalytic pathway is that the expensive pre-treatment step is eliminated, and the catalytic system allows high selectivity towards certain polyols, such as ethylene glycol (EG) and propanediol (1,2-PG) ¹¹ . Although it is understood that catalysis is the central tool in improving efficiency in these conversion processes, from the technical realization is far away due to our inability to control the selectivity and by-product formation.

The (biomass) gasification platform produces synthesis (syn) gas as the intermediate carbon source for subsequent fuel synthesis, either Fischer-Tropsch (FT) synthesis or methane/dimethylether (DME) synthesis, to produce biofuels. The FT synthesis route is advantageous due to providing diesel range fuels.

Generally, there are three main steps in the biomass-to-liquid (BtL) FT synthesis. The FT synthesis was discovered in 1923 by Franz Fischer and Hans Tropsch. After that, several industrial plants have been built around the world using syngas from methane reforming or coal gasification. However, as of now, there are no commercial scale BtL plants like those installed for (coal-to-liquid) CtL or gas-to-liquid (GtL). Most of the BtL plants are either on demonstration- or experimental-scale. Although there are several large-scale biomass gasification systems employed for electricity generation and thermal applications, it is only recently that research has concentrated on converting bio-syngas to higher hydrocarbons via the FT process. The overall process involves complete bond cleavage of biomass molecules followed by bond re-formation for fuel production. This results in a relatively low energy efficiency ⁷ .

A main advantage of the BtL FT is that the basic technology is established, in contrast to the other routes. Nevertheless, the major disadvantage of this process is that this process has a low process thermal efficiency (PTE). Thus, a large amount of the energy content in biomass is irreversibly lost in the biomass conversion steps (typically around 16-50%).

Gasification of the biomass has a PTE of 75%, which represents the maximum PTE possible from syn-gas derived fuels. Taking into account the energy required to produce and transport the biomass decreases the thermal efficiency even further. The losses of energy in a bio based FT-plant are linked with the gasification (23% loss), steam generation (9% loss) and energy recovery as power (24% loss). A maximum attainable energy efficiency is reported to be 46.2%, including about 4% as electricity ⁷ . The process is multi-step, each step with specific challenges ¹² .

In the fast pyrolysis platform, biomass is heated rapidly to 300°C to 600°C in the absence of air, thereby producing black bio-oils containing tars, aromatics, acids, alcohols, aldehydes and other mixed oxygenates. It is a promising process to convert large volume biomass in-situ to bio-oils with a high yield (typically 75%) ¹³ . It makes it easy to transport bio-oils instead of biomass. However, properties such as low heating value due to high moisture content (15-30%), high corrosivity (pH about 2.5), high viscosity, incompatibility with conventional fuels, ash content, incomplete volatility and chemical instability (high contents of aldehydes, ketones and carboxyl acids) negatively affect the quality of bio-oil as a transportation fuel ⁷ . An upgrading process is necessary to convert bio-oil into a ready alternative to the petroleum fuel in the transportation industry. Although

hydrodeoxygenation (HDO) has been extensively studied to upgrade bio-oils, upgrading bio oils to fuels and chemicals remains a formidable challenge, partially due to the complexity in composition with more than 300 compounds ^{14, 15} .

In recent years, fast hydropyrolysis (FHP) has gained significant attention ^{16, 17} . In fast hydropyrolysis, biomass is heated rapidly to 450°C to 600°C in a ¾ environment. In this process, the reducing ¾ gas generates radicals, which reacts with volatiles released by biomass usually with catalyst named as catalytic FHP, that removes oxygen in the form of H2O, CO, and CO2, which lowers the possibility of coking of catalyst ¹⁸ . However, this process based on‘in-situ’ catalysis due to which the pyrolysis and catalysis processes such as catalytic HDO are constrained by the same reaction conditions (e.g. same temperature) even though the optimum conditions for each of the process steps might be different. To avoid this problem, Gas Technology Institute, USA, (GTI) proposed catalytic FHP with ex-situ HDO named as“integrated hydropyrolysis and hydroconversion” (IH2®) technology ^{19, 20} .

Nevertheless, from the catalysis prospective, the effect of having catalysis in both the primary FHP unit and secondary HDO unit is unclear ¹⁶ . As an alternative, there is the H2BioOil process proposed by Purdue University ^{21, 22} , which has non-catalytic FHP with ex-situ HDO. In this process, vapour phase catalytic upgrading is done to avoid secondary reactions during condensation and re-vaporization of the pyrolysis vapours. The main advantage of this process is that the conversion process is envisioned to operate at high ¾ partial pressures, so the HDO reaction rates will be higher due to increased availability of hydrogen at high partial pressures ^{21, 22} . However, introduction of ¾ and high pressure is expensive. Therefore, to make the process energy efficient and economical, a novel process is needed, which can operate at low pressure and ¾.

The ¾ needed for hydropyrolysis has generally been derived from steam reforming of methane or gasification of coal or biomass, both of which will lead to the discharge of CO ₂ in the atmosphere. Recently, it has been proposed that ¾ can be produced from the CO and C ₁ -C ₄ hydrocarbons from HDO reactors, by reforming ²² . However, that process needs costly hydrogen separation process to remove CO ₂ before the hydrogen can be fed into the FHP reactor.

Bio-oil, generated from lignocellulosic biomass, usually consists of C1-C4 light functionalized oxygenates (20%) including aliphatic alcohols, ketones, aldehyde. These oxygenates are thermally and chemically unstable and can be further oxidized to form corrosive acids or low-heating value liquid products. Bio-oil can be upgraded to transport fuel and chemicals via complex reactions (e.g., such as aldol condensation ^{28, 29} ,

oligomerization ³⁰ , ketonization ³¹ , hydroalkylation/alkylation ³² , diels-alder ³³ , guerbet ³⁴ and acylation ³⁵ over catalysts such as T1O2 ^{36, 37} and HZSM-5 ³⁸ ). However, there are usually three to four reactors operated in cascade mode ³⁹ , and the catalysts usually suffer from severe deactivation and carbon loss as light alkenes and alkanes. The most products have less aromatics hydrocarbon content but greater light gases, gasoline and diesel range products ^20,

21

In this perspective is desirable to rationally design catalysts of multifunctional nature and more effective catalytic process that can selectively produce biofuels with tuneable selectivity and low yield of light hydrocarbons as well as low oxygen content. However, selective oxygen removal reaction from methoxy and hydroxy phenolics is still a challenge due to a kinetic competition between the selective hydrodeoxygenation and the selective hydrogenation of aromatic rings. The selective hydrodeoxygenation entails the direct C-0 bond cleavage. The phenolic C-0 bond energy is large (468 kJ/mol) and therefore makes direct hydrodeoxygenation a challenge. Hydrodeoxygenation (HDO) of phenolic compounds was typically carried out at very high pressures (100 bar ¾) on catalysts such as Ni/Zr0 ₂ ⁴⁰ . The high hydrogen transfer activity of phosphides incorporated transition metal catalyst (MoP, Fe ₂ P, N P, CoP and WP) have been widely applied in fossil fuel upgrading towards hydrodenitrogenation and hydrodesulphurization reactions ^41-43 . FeMo phosphide catalyst could selectively cleave C-0 instead of hydrogenating the aromatic ring. Recently, Jason C Hicks et al. have reported highly selective catalysts, FeMoP, NiMoP, RuMoP, for C-0 bond cleavage of aryl ethers or phenolics to aromatics at low pressure, 2.1 MPa ¾ and

temperature, 400°C at near conversion of model reactants ^43-47 . Besides, the understanding of hydrodeoxygenation reactions has been mostly studied using single modelled phenolic components ^{48, 49} , such as creosol, guaiacol, phenol, anisole, to yield monocyclic aromatics ⁵⁰ . However, it is highly desired to improve the catalysts and explore them to more realistic bio-oil mixture at industrial relevant conditions.

In view of the above, there remains a need for improved processes for preparing biofuel, particularly jet fuel, from biomass.

Summary of the invention

The present invention arises from the surprising finding that it is possible to prepare biofuel (i.e. liquid hydrocarbons) by fast hydropyrolysis or fast pyrolysis of biomass, and then using ¾ generated by sorption enhanced steam reforming (SESR) of light (i.e. C ₁ -C ₄ ) hydrocarbons for the fast hydropyrolysis and/or for hydrodeoxygenation (HDO) reactions in a catalytic upgrading reactor. The C ₁ -C ₄ hydrocarbons are themselves generated during the fast hydropyrolysis or fast pyrolysis step, and a significant proportion survive the catalytic upgrading reactions. The SESR produces ¾ at a high yield and purity, and without the requirement for subsequent purification (e.g. to remove CO ₂ ) prior to introduction into the fast hydropyrolysis reactor and/or into the catalytic upgrading reactor.

A further surprising finding of the present invention is that it is possible to upgrade the bio-vapour and/or bio-oil that is initially generated following fast hydropyrolysis of biomass using a fixed bed tandem catalytic-upgrading reactor that contains both an C-C coupling catalyst and a hydrodeoxygenation (HDO) catalyst, and in which the C-C coupling and HDO catalysts form a dual bed system. By integrating C-C coupling and HDO catalyst in this manner, light (i.e. C ₁ -C ₄ ) oxygenates (which are present at significant levels in bio vapour) are converted into C ₅₊ hydrocarbons (i.e. hydrocarbons containing 5 or more carbon atoms, particularly C ₉₊ hydrocarbons) with high yield and purity, such that these

hydrocarbons are useful as biofuel. Further, high levels of aromatic components have also been found following upgrading. The high levels of C ₉₊ hydrocarbons and aromatics are particularly desirable, as these are useful in jet fuels. Moreover, it was possible to use a low temperature of 300-400°C and low pressure of 5 to 20 bar with the fixed bed tandem catalytic-upgrading reactor.

The present inventors have also demonstrated that the Ru modified MoFeP supported on AI2O3 acts as an effective HDO catalyst. Together with representative C-C coupling catalysts, this HDO catalyst was used in a fixed bed tandem catalytic-upgrading reactor to upgrade a simulated bio-vapour feed (comprising acetic acid, acetol, furfural, phenol, guaiacol and eugenol with 30 wt% water, which is representative of the bio-vapour generated by fast hydropyrolysis and fast pyrolysis). The resultant product contained high levels of C5 ₊ aromatic backbone hydrocarbons, with negligible oxygen content.

Overall, the present invention provides a biomass conversion process that is efficient, can be built on a relatively small to large scales and provides biofuel with a high yield and purity, which can be directly used by the transportation sector.

Accordingly, the present invention provides a hydrodeoxygenation (HDO) catalyst which is M '-MOM ² P supported on a support, wherein M ¹ and M ² represent transition metals.

The invention further provides a fixed bed tandem catalytic-upgrading reactor suitable for upgrading bio-vapour or bio-oil from fast hydropyrolysis or fast hydropyrolysis 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 a hydrodeoxygenation (HDO) catalyst.

The invention further provides a method for preparing hydrogen, the method comprising steam reforming a stream comprising C1-C4 hydrocarbons, CO and CO2 from a catalytic upgrading reactor, preferably a fixed bed tandem catalytic-upgrading reactor of the invention, in the presence of a sorbent suitable for CO2 capture, thereby to produce ¾.

The invention further provides 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 of the invention to provide (i) the biofuel, and (ii) a stream comprising C1-C4 hydrocarbons, CO and CO2.

The invention further provides a method for converting bio-vapour and/or bio-oil into biofuel, which method comprises upgrading of the bio-vapour and/or bio-oil in a fixed bed tandem catalytic-upgrading reactor of the invention to provide the biofuel. Brief description of the drawings

Figure 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 _2.

Figure IB 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 _2.

Figure 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.

Figure 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.

Figure 3 is a schematic view of a specific process of the invention described in Example 1.

Figure 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, Fh partial pressure 17 bar, Temperature of reaction: 400°C, reaction time 6 h, WHSV : 0.49/h, weight of catalyst: HDO (Ru/MoFeP-AhOs) =20 g, XTi0 ₂ = 20 g

Figure 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 ₂ partial pressure 17 bar, Temperature of reaction: 400°C.

Detailed description of the invention

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.

The overall process, which is described in further detail below, is illustrated by the specific process depicted in Figures 1 A, IB, 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 ¾ 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 Figures 1 A and 2A or to a carbonate looping by a circulating fluidized-bed (CFB) reactor 7 and 8 in Figures IB and 2B to produce pure Fh, 12. The CO2 captured in the PS SESR reactors will be released 13. The regeneration of CO2 is an endothermic process, requiring heat 11.

The biofuel prepared using the present invention comprises liquid hydrocarbons. The liquid hydrocarbons typically contain five or more carbon atoms (C5 ₊ hydrocarbons).

Preferably the biofuel prepared using the present invention comprises a high proportion of C8-C13 hydrocarbons, such as C9 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.

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.

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 O2). Fast hydropyrolysis involves heating the biomass in the reactor in the presence of hydrogen and the absence of other air (particularly O2). 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 ² .

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.

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.

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.

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 Figure 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.

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.

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.

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 T1O2 or T1O2 doped with Au, Ag, Cu, Pd or Ru. Preferably the aldol condensation and ketonization comprises T1O2 doped with Au, Ag, Cu, Pd or Ru, more preferably T1O2 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%.

Typically C-C coupling catalyst as described above is in the form of pellets.

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 ¾ to form water. For example, in the present invention oxygen-containing hydrocarbon products of the C-C coupling reactions react with ¾, thereby removing oxygen from those hydrocarbons and generating water.

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

M ¹ and M ² represent different transition metals.

M 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 ¹ typically improves the performance of the HDO catalyst, without generally catalysing the HDO reactions itself. Typically, M ¹ represents Rh, Ru, Pt, Pd, Ni, Co or Cu, and preferably represents Ru.

Typically, M ² represents Ni, Co, Fe or Cu, and preferably represents Fe.

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

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 AI2O3, T1O2, ZrCh or CeCb, a carbon material such as activated carbon, mesoporous carbon or carbon nanomaterials, or S1O2. Preferably the support is AI2O3, most preferably g-AEO,.

Thus, a preferred HDO catalyst is Ru-MoFeP supported on AI2O3. Typically the AI2O3 is y-AhCh _. 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.

M^MoM ² ? 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 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 AI2O3). Following impregnation, the coated support is dried and calcined. Next, a solution containing the desired amount of M 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 ¾ and N ₂ ) to convert the M ¹ , Mo, M ² and P components into elemental form.

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

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.

Upgrading using the fixed bed tandem catalytic-upgrading reactor provides (i) the biofuel, and (ii) a stream comprising C ₁ -C ₄ hydrocarbons, CO and CO _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 ₁ -C ₄ hydrocarbons, CO and CO ₂ from the upgrading step may further comprise oxygenates which were not converted into (i) biofuel.

The stream comprising C ₁ -C ₄ hydrocarbons, CO and CO 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 ₁ -C ₄ hydrocarbons, CO and CO ₂ in the presence of a sorbent suitable for CO ₂ capture, thereby to produce ¾. The SESR reactor contains the catalyst required for the steam reforming process together with a sorbent suitable for CO ₂ capture for the in-situ removal of carbon dioxide from the gaseous phase. The steam reforming [including water gas shift (WGS)] and CO ₂ capture reactions are thus conducted simultaneously in one single reactor.

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 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 ^{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.

Steam reforming involves the reaction of C ₁ -C ₄ hydrocarbons and CO with water to provide hydrogen and CO _2. The reactions involved can be illustrated for methane as follows:

CH ₄ + H ₂ O ^ CO + 3 H ₂ [steam reforming reaction] A H _r ° = +184 kJ mol ¹

CO + H ₂ O ^ CO ₂ + H ₂ [water gas shift reaction] A H _r ° = -41 kJ mol ¹

Both of the above reactions are reversible, and so the reactions can be driven towards H ₂ production by removal of CO _2. Removal of CO ₂ is achieved by conducting the steam reforming steps in the presence of a sorbent suitable for CO ₂ capture. The sorbent reacts with the CO ₂ , generally as soon as it is formed, thereby driving the equilibrium towards ¾ production.

Any sorbent that is suitable for CO ₂ capture can be used, but generally CaO-based sorbents are preferred. Natural limestone (primarily CaCCh) and dolomite (primarily

CaCCh.MgCCh) based sorbents being particularly preferred due to their low cost and ready availability (despite suffering from a decay in their CO ₂ capture capacity after several cycles of carbonation/regeneration ²³ ). These natural materials can be converted into their oxides by heating, thereby to provide the sorbent.

For example, a sorbent material can be prepared from limestone (CaCCh) by heating it to provide CaO (and CO ₂ ). The CaO sorbent can then react with CO ₂ to reform the CaCCh,

CO ₂ + CaO ^ CaC0 ₃ [steam reforming reaction] A H _r ° = -178 kJ mol ¹

thereby removing the CO ₂ from the atmosphere. This reaction occurs at low CO 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 CaC0 ₃ by heating, with the relatively pure stream of CO ₂ produced as a by-product being suitable for other uses or sequestration.

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 Figures 1 A and 2A). Initially, steam reforming and CO 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 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 ₂ sorption occurs in the second reactor. The process of switching between reactors can be continued. In this way, a continual stream of Eh can be produced without needing to interrupt the process (i.e. PS SESR is a continuous process).

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

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 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 Fh is at relatively high pressure and so can generally be introduced directly into the fast hydropyrolysis process without the need for re- pressurisation.

The Fh produced by SESR is then typically introduced into the fast hydropyrolysis reactor, though some Fh 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

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

Integration of ¾ production from CO and C1-C4 from an HDO reactor with FHP is presented is represented in Figure 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.

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 (C4+) 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 AI2O3

The MoFeP active phase on supported alumina spheres was prepared by a modified Pichini method ^{51, 52} , using sequential wetness impregnation method. 0.4 M 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 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.

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 ⁰ 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/AhCb was subsequently dried 100°C (for 4 h) and calcined at 500°C using l°C/min heating rate, dwell time at final temperature is was 5 h.

The active phase, Ru-MoFeP, was obtained using temperature reduction method at heating rate of l°C/min at 250 and 700°C to react all phosphorus and reduced oxides to metallic form ^44-46 in the presence of 75 % Fh in Nitrogen.

Example 3

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.

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).

A fixed bed tandem catalytic-upgrading reactor according to the invention was prepared. The aldol catalysts were T1O2 pellets or 0.2 wt%X-Ti0 ₂ , where X represents Au,

Pd or Ru. The X-T1O2 catalysts were prepared by a standard wetness impregnation technique. Ru-FeMoP/AhCb was used as the HDO catalyst. The Ru-FeMoP/AhCb HDO 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.

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

In Figure 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/T1O2+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 Ti02 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.

Figure 5 illustrates the detailed hydrocarbon carbon-number yield as a function of the tandem catalysts. Clearly, the selectivity to CO+CO2 was higher for the Pd/TiC + HDO while C3 (propene and propane) was also greater for only TiCk + HDO. Acetone yield observed was higher for Pd/Ti0 ₂ +HDO. The increasing in CO2 and acetone possibly suggest that Pd/Ti0 ₂ was the most active catalyst for ketonization of acetic acid to acetone, CO2 and water. However, in considering the CO2/CO molar ratio, which gives indication of the preferred reaction pathway, thus either combined decarboxylation (CO2 release) and ketonization (CO2 release) vs decarbonylation (CO release). The CO2/CO ratio of 6.8, 5.6,

5.5, 3.4 and 2.7 was observed for Au/Ti0 ₂ +HDO, Pd/Ti0 ₂ + HDO, Ru/Ti0 ₂ +HDO, TiCk 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/TiCk 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 C10 and C11 ₊ 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/TiCk.

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