BACKGROUND OF THE INVENTION

Field of the Invention

The present invention relates to processes and apparatus for thermochemical conversion of feedstock, particularly but not exclusively gasification processes for thermochemical conversion of biomass. The present invention also provides processes and apparatus for removing carbon dioxide and/or sulphur compounds from mixtures of gases, for example those produced in thermochemical conversion processes.

Related Art

Thermochemical conversion processes, particularly for thermochemical conversion of biomass (e.g. plant material such as wood and wood bark) to produce fuels, have been studied significantly in recent years. Biomass is typically a mixture of hemicellulose, cellulose, lignin and small amounts of other organics. Thermochemical conversion of biomass (e.g. for fuel production) includes gasification and pyrolysis.

The term biomass pyrolysis is typically used to refer to the thermal decomposition of biomass substantially in the absence of oxygen, at temperatures in the range from about 300°C to about 600°C. A traditional example of biomass pyrolysis is the production of charcoal, where the main pyrolysis product is char. Alternative biomass pyrolysis techniques provide a product which, after cooling, includes a substantial proportion of liquid. This liquid is typically a dark brown liquid having a heating value that is around one half of the heating value of conventional fuel oil. This liquid is typically referred to as bio-oil. This bio-oil may be valuable, as it can be stored for later use, e.g. for heat and/or electricity generation. Bio-oil is typically a macroscopically homogeneous mixture of polar organics and water.

Gasification of biomass typically refers to the conversion of biomass into a mixture of gases (which is sometimes referred to as syngas) including hydrogen, carbon

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SUBSTITUTE SHEET RULE 26 monoxide, carbon dioxide and often methane. This can optionally be carried out by using air or steam as the gasifying agent. The latter, which may be referred to by pyrolytic gasification, produces higher calorific product gas quality. The yield of condensable liquid products is lower than for the biomass pyrolysis process outlined above. The gas product of gasification can be used for heat and power, or used to produce chemicals or synthetic fuels. Typically, gasification of biomass takes place at temperatures above about 700°C.

Substances other than biomass, including fossil fuel products such as coal, petroleum coke (petcoke) and heavy residues from distillation of crude oil, and waste materials such as plastic waste can also be broken down using thermochemical conversion processes such as pyrolysis and gasification. For example, coal may be gasified to produce syngas.

In recent years, there has been increasing attention on steam gasification (pyrolytic gasification) of biomass, due to its potential for the production of H ₂ rich gas. This H ₂ rich gas is useful, for example, in electricity generation using hydrogen fuel cells, and in gas turbines/engines. However, technological development of this process to industrial scale have been primarily hampered by (i) the high thermal energy demand required to complete the pyrolysis phase of the reactions, and (ii) contamination of the product gas with environmentally harmful gases and heavy condensable hydrocarbons (tars) that can cause problems for downstream uses of the product gas.

Typically, biomass gasification for fuel gas production is carried out in fluidised bed reactors. In the reactor, the biomass is typically pyrolysed when it reaches a temperature of above 300°C or more, which releases some of the volatile components of the biomass and produces char. In air gasification, this char is typically combusted within the reactor, due to the presence of oxygen. This combustion provides heat energy for the pyrolysis process, which is typically followed by gasification reactions when a temperature above about 700°C is reached. In steam gasification (pyrolytic gasification), either steam or steam-air mixture is introduced into the reactor, as this can enhance the production of H ₂ , via the water- gas shift reaction:

C +H ₂ 0 «→ H ₂ + CO

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SUBSTITUTE SHEET RULE 26 To satisfy the heat demand of the process, the proportion of feedstock material consumed by combustion may be as high as 40% by mass of the total solid feedstock. Additionally, the gas produced in air gasification has a low heating value, as a result of nitrogen diluting the product gas. Furthermore, the C0 ₂ level in the product gas is high, typically in the range of 25 to 50vol% on a dry basis.

In order to satisfy the heat demand of the process and increase the H ₂ yield, recent research has focused on steam gasification in Dual Fluidised Bed (DFB) reactors. This technique involves coupling two reactors: one for gasification and another for char combustion. In the first (gasification) reactor, which may be a bubbling or circulating fluidised bed, the fuel material (e.g. biomass) is fluidised by steam to produce a high heating value gas, while the resulting char and inert solid are sent to the second (combustion) reactor, where combustion takes place due to oxidation with fluidising air (or oxygen). The resulting hot solids, such as inert bed material or ash (and in some cases char which has not been combusted) are returned to the gasification reactor to provide the heat required for the pyrolytic gasification process, thus completing the cycle. Studies on this relatively new technology have confirmed good scale-up potential with possible throughput in excess of 100 tons/day. However, a number of technical, economical and environmental concerns hinder its adoption at an industrial scale.

In particular, the overall thermal balance of the system is complicated, and it can be difficult to accurately control the amount of heat provided to the gasification reactor. This is mainly because gasification and combustion in the two reactors strongly depends on the solid circulation rate of the heat carrier, which is difficult to control independently. For example, circulation of the heat carrier between the two reactors increases exponentially at temperatures greater than about 800 °C. Accordingly, typically it is not possible to run the process continuously, and frequently another fuel input is required to provide the energy required for the pyrolytic gasification, thus increasing the cost and environmental impact of the whole process. Furthermore, large quantities of C0 ₂ are emitted, both from the combustion reactor and the gasification reactor. An external source of steam is also often required.

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SUBSTITUTE SHEET RULE 26 Some recent studies have investigated the possibility of providing heating for thermochemical conversion processes by direct irradiation of a reaction chamber with concentrated solar radiation, in order to reduce the environmental impact of the processes. For example, Reference 1 describes solar thermal cracking of methane, by direct irradiation with concentrated solar radiation. Similarly, Reference 2 describes thermochemical conversion of biomass by direct solar irradiation. The direct solar irradiation approach has been shown to provide sufficient process heating to drive highly endothermic reactions such as gasification of petcoke, coal and biomass. However, commercial investment and scaling-up is hindered because it can be technically difficult to carry out biomass thermal conversion simultaneously with solar irradiation through a transparent window in the same reactor. This is particularly challenging when operating at a high temperatures and/or pressure with a multi-component flow mixture of gas, solid and in many cases sticky tars.

References 2 and 3 describe providing indirect heating with concentrated solar radiation through the reactor walls, and Reference 4 discusses the possibility of heating a reactor for steam gasification of biomass, using a fluidised bed of sand heated using molten salt from a solar concentrator. However, significant challenges remain, including the need to provide control of heating of the reactor, and the need for a process suitable for scaling up to an industrial scale.

There remains a need for improved thermochemical conversion processes, and in particular thermochemical conversion processes for processing biomass to produce useful products.

SUMMARY OF THE INVENTION

The present invention has been made to reduce, address, ameliorate or avoid one of more of the problems set out above. In particular, the present inventors have found that heat energy for thermochemical conversion processes can conveniently be provided by solar radiation. As the skilled person will understand, this can significantly reduce the cost and environmental impact of thermochemical conversion processes such as those discussed above.

In particular, the present inventors have realised that solar radiation can be used to provide the required heat energy by heating heat carrier particles using concentrated

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SUBSTITUTE SHEET RULE 26 solar radiation, and supplying the heated heat carrier particles to the thermochemical conversion reaction chamber. In this way, heat for thermochemical conversion processes such as pyrolysis and gasification can be provided, and the need to combust part of the feedstock or an external fuel to provide energy is reduced or avoided. Conveniently, this also allows for improved control of the heat supplied for thermochemical conversion, for example by varying the length of time the heat carrier particles are exposed to concentrated solar radiation, and/or by varying the flow rate of the heat carrier particles into the reaction chamber. Varying the material of the heat carrier particles can also provide control of the heating process.

Accordingly, in a first preferred aspect, the present invention provides a process for the thermochemical conversion of feedstock, comprising heating the feedstock to a thermochemical conversion temperature in a reaction chamber to produce one or more products of thermochemical conversion of the feedstock, wherein heating the feedstock to said thermochemical conversion temperature comprises:

flowing heat carrier particles through a focal zone of a solar concentrator to heat the heat carrier particles; and

supplying the heated heat carrier particles to the reaction chamber to heat the feedstock to said thermochemical conversion temperature.

Preferably, the feedstock is biomass. However, the present inventors consider that the process of the first aspect may also be suitable for thermochemical conversion of feedstocks other than biomass. Suitable feedstocks are not particularly limited, but may include for example fossil fuels such as coal, petroleum coke (petcoke) or heavy residues from distillation of crude oil, and waste material such as municipal and plastics waste material. Alternatively, mixtures of different feedstocks may be used, for example a blend of biomass with a fossil fuel such as coal.

In the process of the present invention, preferably the heat carrier particles are circulated. It will be understood that the heat carrier particles may be circulated by removing them from the reaction chamber for heating in the solar concentrator, then returning the heated heat carrier particles to the reaction chamber, as described above. It will be understood that this process may be repeated to provide ongoing circulation of the heat carrier particles. For example, the heat carrier particles may be circulated substantially continuously. It will be understood that the solar concentrator is provided at a position external to the reaction chamber.

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SUBSTITUTE SHEET RULE 26 Preferably, the heat carrier particles are fluidised to provide a fluidised bed comprising heat carrier particles in the reaction chamber. Accordingly, it will be understood that the thermochemical conversion process may be carried out in a circulating fluidised bed reactor.

It will be understood that the thermochemical conversion may be, for example, pyrolysis and/or gasification (e.g. steam gasification). Typically, the one or more products of thermochemical conversion include a product gas. The product gas may comprise, among other components, carbon dioxide and/or sulphur compounds such as S0 ₂ and H ₂ S. For downstream uses of the product gas, it may be preferable to remove at least a portion of these compounds from the product gas.

Accordingly, the process of the first aspect may further comprise treating the product gas to remove at least a portion of the carbon dioxide and/or sulphur compounds from the product gas. Typically, this treatment comprises contacting the product gas with sorption particles, to remove at least a portion of the carbon dioxide and/or sulphur compounds from the product gas by adsorption, e.g. chemisorption. It will be understood that the sorption particles comprise a material to which carbon dioxide and/or sulphur compounds may be adsorbed, to remove the carbon dioxide and/or sulphur compounds from the product gas. For example, the sorption particles may comprise CaO, which can act to remove at least a portion of the carbon dioxide and/or sulphur compounds from the product gas by chemisorption (see Refs 8 and 9). It will be understood that chemisorption of the carbon dioxide and/or sulphur compounds typically involves a chemical reaction of the carbon dioxide or sulphur compounds with the sorption particles. For example, CaO typically reacts with carbon dioxide to produce CaC0 ₃ . Similarly, CaO typically reacts with sulphur containing compounds to produce CaS, CaS0 ₄ and CaS0 ₃ .

The reacted sorption particles may be regenerated (i.e. the adsorbed substances may be released from the particles, e.g. by conversion of reacted CaO back to CaO) by heating the reacted sorption particles to a regeneration temperature. This heating may be carried out by flowing the sorption particles through a focal zone of a solar concentrator, to heat the particles to a regeneration temperature. Reference 10

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SUBSTITUTE SHEET RULE 26 describes thermochemical regeneration CaO from CaC0 ₃ , followed by formation of Ca(OH) ₂ in a slaker.

Accordingly, in some embodiments the present invention provides a process for the thermochemical conversion of feedstock comprising heating the feedstock to a thermochemical conversion temperature in a reaction chamber to produce one or more products of thermochemical conversion of the feedstock,

wherein heating the feedstock to said thermochemical conversion temperature comprises:

flowing heat carrier particles through a focal zone of a solar concentrator heat the heat carrier particles; and

supplying the heated heat carrier particles to the reaction chamber to heat the feedstock,

wherein the one or more products of the thermochemical conversion of feedstock includes a product gas comprising carbon dioxide and/or sulphur compounds, and wherein the method further comprises:

contacting the product gas with sorption particles in a gas cleaning zone, to remove at least a portion of the carbon dioxide and/or sulphur compounds from the product gas by chemisorption to the sorption particles, thereby producing reacted sorption particles; and

regenerating the sorption particles by flowing the reacted sorption particles through a focal zone of a solar concentrator to heat the sorption particles to a regeneration temperature.

Conveniently, the heat carrier particles and the sorption particles may be heated using the same solar concentrator, e.g. they may be flowed through the same focal zone of the solar concentrator. It will be understood that a mixture of the heat carrier particles and the sorption particles may be flowed through the focal zone. This method of removing carbon dioxide and/or sulphur compounds from a gas, in which sorption particles are regenerated using concentrated solar radiation, provides convenient and efficient sorbent regeneration without the need to supply an external fuel source, and without the need to rely on process heat from elsewhere in the reactor, which can limit the efficiency of regeneration. Accordingly, the present inventors consider that this method is independently novel and inventive. In a second preferred aspect, therefore, the present invention provides a method for

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SUBSTITUTE SHEET RULE 26 removing carbon dioxide and/or sulphur compounds from a mixture of gases including carbon dioxide and/or sulphur compounds, the method comprising:

contacting the mixture of gases with sorption particles in a gas cleaning zone, to remove at least a portion of the carbon dioxide and/or sulphur compounds from the mixture by chemisorption to the sorption particles, thereby producing reacted sorption particles; and

regenerating the sorption particles by flowing the reacted sorption particles through a focal zone of a solar concentrator to heat the sorption particles to a regeneration temperature.

Preferably, the sorption particles comprise (e.g. consist essentially of) CaO. The carbon dioxide and/or sulphur compounds may be removed from the gas by chemisorption to the CaO of the sorption particles. Other suitable sorbent materials which may be suitable include pure MgO and dolomite (CaC0 ₃ +MgC0 ₃ ) particles.

Preferably, the regenerated sorption particles are returned to the gas cleaning zone, and the method repeated. In this way, the sorption particles may be circulated, e.g. substantially continuously. Over time, it may be desirable to replace "spent" sorption particles with fresh sorption particles.

In a further preferred aspect, the present invention provides thermochemical conversion apparatus for thermochemical conversion of feedstock, comprising:

a reaction zone (e.g. chamber) for heating the feedstock to a thermochemical conversion temperature to produce one or more products of thermochemical conversion of the feedstock; and

a solar concentrator arranged to concentrate solar radiation into one or more focal zones, the solar concentrator being arranged to flow heat carrier particles through a focal zone of the solar concentrator to heat the heat carrier particles, wherein the apparatus is configured to circulate the heat carrier particles between the reaction zone and the solar concentrator to provide a circulating fluidised bed of heat carrier particles.

In a further preferred aspect, the present invention provides gas cleaning apparatus for removing carbon dioxide and/or sulphur compounds from a mixture of gases, comprising:

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SUBSTITUTE SHEET RULE 26 a gas cleaning zone (e.g. chamber) for contacting the mixture of gases with sorbent particles to remove carbon dioxide and/or sulphur compound from the mixture by chemisorption; and

a solar concentrator arranged to concentrate solar radiation into one ore more focal zones, the solar concentrator being arranged to flow sorbent particles through a focal zone of the solar concentrator to heat the sorbent particles to a regeneration temperature,

wherein the apparatus is configured to circulate the sorption particles between the gas cleaning zone and the solar concentrator.

The features of any aspect of the invention may be combined, singly or in combination, with any other aspect, unless the context demands otherwise. Any preferred or optional features may be combined, either singly or in combination, and may be combined with any aspect of the invention, unless the context demands otherwise.

BRIEF DESCRIPTION OF THE DRAWINGS

Fig. 1 illustrates a biomass pyrolytic gasification process according to an embodiment of the present invention.

Fig. 2 illustrates a method for removing carbon dioxide and/or sulphur compound from a mixture of gases according to an embodiment of the present invention.

Fig. 3 illustrates apparatus according to a preferred embodiment of the present invention.

Fig. 4 is a flow diagram illustrating a process according to an embodiment of the present invention.

Fig. 5 shows the results of a mass and energy balance calculation for a solar concentrator as set out in the Example.

Fig.6 shows a scaling up analysis of the calculation of Fig. 5.

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SUBSTITUTE SHEET RULE 26 Fig. 7 shows the results of a mass and energy balance calculation for a gasifier as set out in the Example.

Fig. 8 provides a breakdown of the total energy demand of a process according to an embodiment of the present invention.

Fig. 9 illustrates the thermal energy available in a conventional dual fluidized bed system gasifier (from combustion of char) compared with the energy demand for steam gasification as explained in the Example.

DETAILED DESCRIPTION

Further preferred and optional features of the invention will now be set out.

As discussed above, the nature of the feedstock of the thermochemical conversion process is not particularly limited. However, preferably it is biomass. Preferably, the feedstock is supplied to the reaction chamber as feedstock particles. A suitable number average particle diameter for the feedstock particles is up to about 3cm equivalent spherical diameter. Typically, the feedstock particles have a number average particle equivalent spherical diameter of at least about 500μπι, at least about 1 mm or at least about 10mm.

Preferably, gas is supplied to the reaction chamber to fluidise the feedstock particles. It will be understood that the feedstock particles are typically fluidised together with the heat carrier particles. Accordingly, the reaction chamber may be a fluidised bed reactor.

The nature of the thermochemical conversion reaction also is not particularly limited in the present invention. However, particularly suitable thermochemical conversion processes include gasification (e.g. steam (pyrolytic) gasification) and pyrolysis. Where the process is gasification, especially gasification of biomass or a fossil fuel such as coal, preferably the feedstock particles are fluidised in a gas comprising steam, since this can advantageously enhance the production of H ₂ via the water-gas shift reaction. C0 ₂ (e.g. C0 ₂ produced in the process) may also be supplied to the reaction chamber, as this can improve the quality of the product gas. Where the process is pyrolysis only, especially pyrolysis of biomass, the feedstock particles are

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SUBSTITUTE SHEET RULE 26 preferably fluidised in a gas having a very low concentration of oxygen, or in the (complete) absence of oxygen, in order to produce a high quality bio-oil

It will be understood the heat carrier particles are typically heated to a temperature greater than the thermochemical conversion temperature, to enable them to heat the feedstock to the thermochemical conversion temperature in the reaction chamber. Preferably, the heat carrier particles are heated to a temperature at least 50°C or at least 100°C higher than the thermochemical conversion temperature. Where the process is a gasification process, typically the thermochemical conversion temperature is at least about 700°C, at least about 800°C or at least about 900°C. Where the process is a pyrolysis process, typically the thermochemical conversion temperature is at least about 350°C, at least about 450°C, or at least about 550°C.

Preferably, the heat carrier particles are metal heat carrier particles (e.g. metal alloy heat carrier particles). Metal heat carrier particles are preferred, for example, due to their high thermal conductivity, meaning that they can be sufficiently heated following exposure to the focal zone of a solar concentrator, e.g. for a relatively short period of time.

It may be preferable that the heat carrier particles are magnetic. Particularly preferable are magnetic metals having a Curie point of at least about 900°C, or at least about 1000°C, since such metals may retain their magnetic properties at the temperatures employed in the process of the invention. This allows for convenient separation of the heat carrier particles using a magnetic separator. For example, the heat carrier particles may be separated from other components from the product stream leaving the reaction chamber using a magnetic separator. In embodiments where a mixture of heat carrier particles and sorption particles are flowed through the same focal zone of a solar concentrator, the heated heat carrier particles may be separated from the sorption particles after leaving the solar concentrator (e.g. using a magnetic separator or one or more cyclone separators, e.g. arranged in series). Typically, at this stage the particles may be at a temperature of about 1000°C, e.g. in gasification processes. It will be understood that the apparatus may be provided with one or more magnetic separators, arranged to separate the heat carrier particles as described above.

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SUBSTITUTE SHEET RULE 26 Preferably, the heat carrier particles are catalytically active, e.g. in catalysing the thermochemical decomposition (cracking) of tars produced in the thermochemical conversion process, since this can improve the quality of the product. Where the metals are catalytically active, over time the activity may be degraded by catalyst poisoning, e.g. due to adsorption of sulphur compounds. Regeneration of catalytic activity is possible by heating the particles, e.g. to a temperature of about 1000°C. It will be understood that numerous metals and metal alloys are suitable for the heat carrier particles, and accordingly the metal of the heat carrier particles is not particularly limited. For example, the heat carrier particles may comprise iron or iron alloys such as steel. Nickel or nickel alloys may also be suitable. However, cobalt and cobalt alloys are particularly preferred.

Cobalt is particularly preferred because it has particularly high thermal conductivity (-70 w/mK), it has high resistance to corrosion and wear (even under the challenging conditions of high temperature thermochemical conversion processes), it has activity in catalysing cracking of tar (see Reference 7), and remains magnetic at high temperatures (it has a Curie point above 1100°C). Similar comments apply to cobalt alloys. Preferably, the heat carrier particles have an average equivalent spherical diameter of at least about 100pm. Preferably, the heat carrier particles have a number average diameter of 1 mm or less, or 0.5mm or less. Typically, the heat carrier particles have a number average diameter in the range from about 200μιπ to about 500μηι. Heat carrier particles within this size range are suitable for fluidising in the reaction chamber, and exhibit suitable heat transfer properties in the solar concentrator and in the reaction chamber.

Typically, the products of thermochemical conversion include a gas with entrained solids, such as char and ash. Typically, the heat carrier particles are also carried with the product gas stream. The solids may be removed e.g. by cyclone separators, and/or by a magnetic separator as described above. Similarly, a cyclone separator and/or a magnetic separator may be used to separate the heat carrier particles from other solids, e.g. for reheating the heat carrier particles in the solar concentrator, and/or for downstream uses of the other solids.

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SUBSTITUTE SHEET RULE 26 Where the products include condensable vapours, the products may be passed through a condenser to condense these vapours. For example, tar may be removed from the product stream in this way. For example, tar may be condensed by spraying with water to produce water and tar emulsion. An anti-fouling or dispersant composition may be added at this point. The water of this emulsion may be heated to form steam, e.g. by heat exchange with the hot product stream exiting the reaction chamber. This steam may be flowed into the reaction chamber, e.g. to enhance H ₂ production in the gasification of biomass.

Where the thermochemical process is a biomass pyrolysis process, for example, the product vapours may be condensed to form bio-oil, as the skilled person will readily understand.

The heat carrier particles are then preferably returned to the focal zone of the solar concentrator for reheating, as described above. The product gas may be treated, e.g. to remove carbon dioxide and/or sulphur compounds, as described herein.

The char may be stored, e.g. to provide process heat by combustion when the solar concentrator is not operating (e.g. during the hours of darkness) or used as a fertiliser. The ash may be used, for example, as a fertiliser or cement filler.

As discussed above, a mixture of gases (e.g. a product gas from the thermochemical conversion process) may be contacted with sorbent particles to remove carbon dioxide and/or sulphur compounds from the gas.

The sorbent particles may be provided to in the reaction chamber itself. However, this can be undesirable, particularly where the thermochemical conversion process is gasification (where the temperature in the reaction chamber is typically >700°C). In order to ensure sufficient sorption of the carbon dioxide and/or sulphur compounds, it may be necessary to reduce the temperature in the reactor, which can lower the thermochemical conversion efficiency, and in some cases can considerably increase the amount of tars present in the product gas.

Accordingly, preferably the gas is contacted with the sorption particles at a location external to the reaction chamber. For example, the gas may be contacted with the sorption particles in a gas cleaning zone, e.g. chamber, which is separate from the

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SUBSTITUTE SHEET RULE 26 reaction chamber where thermochemical conversion takes place. The product gas from the thermochemical conversion may be transferred to the gas cleaning zone for removal of carbon dioxide and/or sulphur compounds.

Typically, the sorbent particles have a number average equivalent spherical diameter of at least about 200pm, and a number average equivalent spherical diameter less than about 2mm. Typically, the mixture of gases is contacted with the sorption particles at a temperature of below 400°C, or below 250°C. Typically, the temperature is about 200°C or less. Typically, the regeneration temperature is about 800°C or more. For example, it may be 850°C or more or 900°C or more.

Using the process of the present invention, the chance of the sorption particles sintering during high temperature regeneration may be reduced. The rapid circulation of the particles causes particle-particle and particle-wall collisions and shearing which can reduce agglomeration. Where a mixture of heat carrier particles and sorption particles is flowed through the focal zone of the solar concentrator, the risk is further reduced due to the presence of a large fraction of heat carrier particles which may reduce the interaction between the sorption particles. Furthermore, where a small amount of moisture is present in the product gas, this may help to maintain the pore volume of the sorbent, which can increase the time its activity is maintained.

If excess CaC0 ₃ is produced in the sorption process, this may be employed as slurry feed for scrubbing any S0 ₂ traces from the flue gas mixture leaving the solar concentrator.

The skilled person will readily understand the terms solar concentrator and concentrated solar radiation. Solar concentrators typically employ mirrors and/or lenses to concentrate solar radiation onto a small area. Commercially, solar concentrators have been used for example to generate steam to drive steam turbines for electricity generation. A review of solar concentration technologies is provided in References 5 and 6, which are hereby incorporated by reference in their entirety and in particular for describing and defining solar concentrators and solar concentration systems.

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SUBSTITUTE SHEET RULE 26 At a general level, a solar concentrator typically comprises a concentration system arranged to focus sunlight at a focal zone, and is typically arranged to allow a substance to be heated to pass through the focal zone. It will be understood that the concentration system typically comprises one or more mirrors and/or lenses.

Existing solar concentration systems may broadly be divided into two classes - linear focusing systems (wherein the focal zone is a line) and point focussing systems (wherein the focal zone is a point). Suitable linear focusing systems include parabolic troughs and linear Fresnel systems. A parabolic trough has a linear parabolic reflector (e.g. mirror) which concentrates solar radiation along the reflector's focal line. Fresnel reflectors instead comprise a plurality of flat reflector (e.g. mirror) strips arranged to concentrate solar radiation into a focal line. The use of flat reflectors means that Fresnel systems may provide a larger surface area of reflectors than a linear parabolic system. Typically, the substance to be heated by the solar radiation is provided at the focal line of these linear reflectors.

A parabolic dish reflector is an example of a point focussing solar concentration system, and is typically provided with a receiver positioned at the reflector's focal point. Solar towers typically comprise an array of movable mirrors (e.g. which track the sun), which are typically flat and are arranged to focus solar radiation at a focal point.

Typically, solar radiation received by the concentrator is about 1kW m ^"2 . A typical solar concentration system concentrates this solar energy by a factor of at least 500, at least 1000, at least 2500. A solar concentrator may concentrate the solar energy by a factor as high as 5000.

Preferably, the solar concentrator is arranged to allow flow of the heat carrier particles and/or the sorption particles through its focal zone to heat them. For example, a conduit or chamber (solar furnace) may be provided to allow flow of the particles through the focal zone. It will be understood that the conduit or chamber may be provided with a window or other opening to allow direct irradiation of the particles by the concentrated solar radiation. However, it may instead be preferable that the particles are not exposed to direct irradiation, but instead pass through a closed conduit or chamber positioned at the focal zone, and so are heated by the action of the concentrated solar radiation on the walls of the conduit or chamber. In

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SUBSTITUTE SHEET RULE 26 this way, the heat carrier particles and/or the sorbent particles are exposed to concentrated solar radiation directly or through the walls of the conduit.

A biomass gasification process according to an embodiment of the present invention will now be described with reference to Fig. 1. Biomass particles are supplied to a circulating fluidised bed gasifier 1 for pyrolytic gasification. In the gasifier 1 , the biomass particles are fluidised by steam to enhance H ₂ yield via water-gas shift reaction. The highly endothermic pyrolysis step of the gasification process is driven by heat from circulating heat carrier particles, which have been heated in a solar concentrator 4. This pyrolysis reaction is followed by a range of gasification reactions, including the water-gas shift reaction as shown in Table 1 below (reactions 1-5, side reactions omitted). These reactions collectively contribute to increasing H ₂ formation, and also to the production of C0 ₂ . Typically, sulphur will also be present in the biomass feedstock (typically in the range of 0.5-2wt% on dry basis). Optionally, a small amount of air (well below the stoichiometric ratio for complete combustion) may be introduced to the gasifier to promote formation of sulphur dioxide (S0 ₂ ) and at a later stage, hydrogen sulphide (H ₂ S) according to reactions 6 and 7 in Table 1. The hot product gas leaving the gasifier passes through at least one heat exchanger 2 to cool the gas to a temperature of about 300°C. This is followed by rapid cooling/scrubbing in concentric water cooling and spray unit 3 to reduce the gas temperature to less than 40°C, causing condensation of water vapour and scrubbing of heavy organics (tars). Unit 3 is arranged to provide simultaneous internal water spray cooling/scrubbing and external water cooling. The condensed liquid, comprising mainly water and tars, forms a reasonably uniform emulsion suitable for pumping and vaporization. Alternatively, oil or an oil-water mix may be used for spray cooling and scrubbing to provide a uniform emulsion. An antifouling chemical and/or a dispersant may optionally be supplied to unit 3 if necessary.

The vaporized emulsion, and additional steam recovered from the cooling step, if needed, are sent back to the circulating fluidised bed gasifier 1 , while the substantially tar free and relatively dry product gas is sent for further cleaning to eliminate the pollutant gases such as C0 ₂ , S0 ₂ and H ₂ S using CaO particles in a second circulating fluidised bed reactor, as described in below. The tar content of the product gas may be as low as 5 mg/Nm ³ . The solids (heat carrier particles, char

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SUBSTITUTE SHEET RULE 26 and ash) and the product gas leaving the gasifier pass through a series of cyclone separators, first to separate the cobalt and send it to the solar concentrator 4 for heating; second to separate the char and send it to storage, and finally to separate the ash for disposal. The char may be combusted to provide heat energy for the process when solar radiation is not available, or alternatively sold as a bio-char fertiliser.

Table 1. Summary of chemical reactions

A method for removing carbon dioxide and/or sulphur compound from a mixture of gases according to an embodiment of the present invention will now be described with reference to Fig. 2. In this process, a mixture of gases (product gas from the circulating fluidised bed gasifier 1 of Fig. 1) is subjected to dry cleaning at a temperature below about 300°C using CaO sorbent, in a circulating fluidised bed sorber 5 (gas cleaning zone). At this temperature, and substantially in the absence of oxygen, chemisorption takes place at the surface of the sorbent material to separate C0 ₂ according to reaction 8 in Table 1. Similarly, the sulphur compounds (S0 ₂ and H ₂ S) may undergo a range of reactions (reactions 9-12 in Table 1). Through these reactions, the harmful C0 ₂ , S0 ₂ and H ₂ S gases may be eliminated from the product gas and retained in a particulate phase of CaS, CaS0 ₄ , CaS0 ₃ and CaC0 ₃ .

The CaO sorbent can also serve in catalytic cracking of any remaining tar in the product gas. The carbonated/sulphated solid mixture leaving the sorber 5 is then sent to the focal zone of a solar concentrator 4, where the internal temperature reaches about 1000 °C. At this temperature, CaC0 ₃ thermally decomposes to produce regenerated CaO and C0 ₂ gas, while the remaining solids may be involved in different reaction as follows:

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SUBSTITUTE SHEET RULE 26 (i) the CaS and CaS0 ₄ undergo solid-solid reaction to produce S0 ₂ and CaO according to reaction 13 in Table 1 ; and

(ii) the CaS0 ₃ thermally decomposes to produce CaO and S0 ₂ according to reactions 14 in Table 1.

The regenerated sorbent is sent back to the sorber 5 to complete the cycle. The regenerated sorbent may be passed through a loop seal, which serves to control the solid mass flow as well as cooling the sorbent down to the desired temperature for sorption. To replace any spent and/or elutriated CaO which is carried out of circulation (e.g. in the product gas), fresh CaO particles may be provided. Finally, the high calorific value dry product gas leaving the sorber 5 is subjected to gas cooling and filtration before being compressed and stored or directly sent to a gas turbine/engine for electricity generation. The flue gas exiting the solar concentrator 4, comprising mainly C0 ₂ and S0 ₂ , is sent for cooling and limestone scrubbing, to eliminate possible S0 ₂ traces, thus producing a high purity C0 ₂ with less than 1 ppm S0 ₂ concentration. Further oxidation of the by-product (CaS0 ₃ ) produces CaS0 (gypsum), which is a valuable commercial product.

A biomass gasification process according to a particularly preferred embodiment of the present invention will now be described with reference to Figs. 3 and 4. This preferred process integrates the thermochemical conversion and gas cleaning processes of the present invention.

The process is summarised in the flow diagram shown in Fig. 4, and the apparatus is illustrated in Fig. 3.

The apparatus comprises a fluidised bed gasifier 101. Biomass particles are fed into the gasifier 101 for gasification. They are fluidised by steam, and optionally C0 ₂ and/or air, which are preferably pre-heated in a heat exchanger 112. Heat carrier particles are also supplied to the gasifier 101. Prior to entering the gasifier 101 , the heat carrier particles are heated to about 1000°C in a solar concentrator 104.

Products of the gasification reaction are removed from the gasifier 101 , along with the heat carrier particles. Cyclone separators 106 separate solids from the product gas. Char and ash are removed. The char may be stored in a char storage chamber 108 (e.g. for providing heating when solar radiation is unavailable, or used as a fuel

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SUBSTITUTE SHEET RULE 26 in other processes or sold as a bio-char fertiliser. Ash is removed, e.g. for use as a fertilizer or cement filler. The heat carrier particles are returned to the solar concentrator 104 for reheating, after which they are returned to the gasifier 101 to provide further heating for the ongoing gasification process.

The product gas is sent to a tar/water eliminator 107, where water and tar vapours present in the product gas are removed by cooling and scrubbing. Antifouling chemical and/or a dispersant may be supplied to the quencher. Further water may be added if required. The condensed water may be re-vaporised for supplying to the gasifier 101 as steam. The tar may also be returned to the gasifier, for further decomposition.

After water and tar removal, the product gas is sent to a sorber 105, where it is contacted with CaO sorption particles to remove C0 ₂ and sulphur containing compounds such as S0 ₂ and H ₂ S, by chemisorption to the CaO, as described above with reference to Fig. 2. The reacted CaO sorption particles leaving the sorber 105 (gas cleaning zone) are separated from the clean product gas in a cyclone separator 109. The clean product gas, which is rich in H ₂ , may be used for a variety of uses as described herein. For example, metal heat carrier particles with catalytic activity (e.g. cobalt particles) may be treated with some of the H ₂ rich product gas, to restore their catalytic properties, where these have been poisoned e.g. by carbon and/or sulphur in the gasifier 101.

The reacted CaO sorption particles are sent to the solar concentrator 104 for heating to a temperature of about 1000°C, to regenerate the CaO. The flue gas produced, which contains mainly C0 ₂ and S0 ₂ is cooled in a heat exchanger 112, then scrubbed using a limestone slurry in a scrubbing unit 110, to produce relatively pure C0 ₂ . The C0 ₂ may, for example, be compressed and shipped to market, sent to deep sea/geological storage, or sent back to the gasifier 101 to further enhance H ₂ production.

The regenerated CaO particles are returned to the sorber 105. Fresh CaO particles may be added as needed.

The metal heat carrier particles (e.g. cobalt or cobalt alloy heat carrier particles) and the CaO particles are heated together in the solar concentrator. After heating, the

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SUBSTITUTE SHEET RULE 26 heat carrier particles and CaO sorption particles are separated in a magnetic separator 111 , and returned to the gasifier 101 and the sorber 105 respectively.

Example - mass and energy balance calculation

Set out below are mass and energy balance calculations for the solar concentrator (SRC) and the circulating fluidised bed gasifier (CFB-1 ), in a process according to an embodiment of the present invention.

The calculation was carried out using a spread sheet, using available experimental data of a typical steam gasification process (solid fuel feed, product gas quality, etc) in addition to a number of assumptions and simplifications. A summary of the parameters used in this analysis are given in Table 2.

Table 2. Assumed steam asification arameters

1.0 Solar Concentrator It is assumed that the solar concentrator/reactor (SRC) is operating at a maximum pressure of 10 bars. The solids entering the SRC are mainly introduced for two purposes:

1. Supply of the heat required to complete the highly endothermic steam gasification in CFB-1 , in the form of hot metal heat carrier particles

2. Supply of the heat required for the regeneration of the carbonated/sulphated sorbent to produce CaO, C0 ₂ and traces of S0 ₂ inside the SRC.

The operating temperature of the SRC is assumed to be around 1000 °C. All the solids and gas are expected to leave at the SRC at the same temperature. The flue gas leaving the SRC (mainly C0 ₂ and traces of S0 ₂ ) will be sent to cooling and lime

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SUBSTITUTE SHEET RULE 26 scrubbing, to eliminate S0 ₂ traces, thus producing a high purity C0 ₂ with expected range of less than 1ppm S0 ₂ (this is ignored in the mass/energy balance calculation). The C0 ₂ can optionally be sent to the gasifier, to enhance H ₂ yield, or compressed and shipped to market or sent to deep sea and geological storage (sequestration). The regenerated sorbent (CaO) and hot metals (metal heat carrier particles) leaving the SRC will pass through a magnetic particle separator, where the metal heat carrier particles will be selectively separated and diverted to the CFB gasifier to provide the heat necessary for completing the gasification reaction, while the CaO particles will be sent back to CFB gas cleaning unit for C0 ₂ and sulphur compound sorption, thus completing the closed loop gasification cycle as shown in Figs. 3 and 4.

1.1. Mass and energy balance

The particulate material involved in the SRC includes three main phases. These are:

(i) calcium carbonate (CaC0 ₃ )

(ii) metal heat carrier particles (e.g. cobalt or cobalt alloy)

(iii) calcium oxide (CaO)

For the sake of simplicity and due to their limited contribution to the overall mass and energy balance, sulfur compounds are neglected in this analysis. The mass flow of CaO sorbent is estimated based on recommended CaO to C0 ₂ mass flow ratio of

According to the flow diagram shown in Fig. 5, the following energy balance applies (heat losses are assumed to be 10% of the total heat supplied to the system: where Q _solar represents the required solar heat, the first term on the right side represents the heat of reaction and the second and third terms represent the heat in the feed and products. The mass flow of metal heat carrier is based on a typical range of operating condition of a steam biomass gasifier, which is given in terms of solid inert to dry biomass teed ratio of m _iner m _biomass = 10.

The sorbent regeneration reaction is highly endothermic and is described by the following reversible reaction:

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SUBSTITUTE SHEET RULE 26 CaCOs <→ CaO + C0 ₂ AH=176.0 kJ/kg

Other reactions, which may take place inside the SRC, are of limited contribution to the overall energy balance, thus will be neglected. Reported experimental studies have shown complete regeneration of the sorbent material at temperature of >900 °C.

The mass and energy balance around the solar reactor/collector (SRC) is shown Fig. 5. This is based on a proposed 0.57MW thermal input. The scaling up analysis shown in Fig. 6 indicates linear increase in the solar energy demand with increasing the solid fuel feed rate (directly related to the inert solid and C0 ₂ flow in the closed loop system).

2.0 Circulating fluidised bed gasifier

2.1 Brief process description

In this analysis it assumed that the CFB gasifier is fluidised by steam to enhance H ₂ yield, mainly via water-gas shift reaction. The highly endothermic pyrolysis phase of the gasification process will be indirectly driven by the intense solar heat captured in the solar concentrator (SRC) by the circulating metal heat carrier particles. The gasifier is assumed to operate at a maximum pressure of 10 bars. The metal particles, steam and solid fuel are assumed to enter the gasifier at 1000°C, 158°C and 25°C respectively, while all the exiting materials (metal particles, product gas, char and ash) are all assumed to leave at 700°C. The solid feed fuel, which is assumed to be 100% biomass wood (see Table 3), will be thermally decomposed to produce char, pyrolysis gas and tar as shown in Table 4. Note that the ash formation is ignored the heat and mass balance due to its insignificant effect. The various pyrolysis gases components are then believed to undergo a range of homogenous and heterogeneous reactions to produce the product gas given in Table 5.

Table 3. Biomass fuel quality

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SUBSTITUTE SHEET RULE 26

Table 4. Pyrolysis products of fuel feed

Table 5. Assumed product gas quality

2.2 Mass and energy balance

The mass and energy balance around the gasifier is shown in Fig. 7. The calculation is based on a gasifier of 1.2 MWe. In the calculation, the heat losses are assumed to be 10% of the total heat supplied to the system. The energy balance is given by:

where Q represents the energy added or removed from the reactor. The first and second terms in the right side represent the heat of formation and enthalpy in the feed and product respectively. The heat of formation of the solid fuel feed is given by:

AH _f ° _{m feed} = HHF-(327.63C + 1417.94H+92.575' + 158.67iF)

Where HHV is the fuel higher heating value, , H, S and Ware the percentage weight of carbon, hydrogen, sulphur and water content in the fuel respectively (see Table 3).

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SUBSTITUTE SHEET RULE 26 The total thermal energy demand for the gasification of the solid fuel feed compared to the performance of a conventional DCFB (without external heating) is shown in Fig. 9. The shaded area demonstrates the energy deficiency in the DCFB system, which amount to around 55% of the total thermal energy required to maintain the gasifier at the desired temperature of >900°C. This is clearly confirms the potential of the proposed integrated solar gasification system in energy saving.

In terms of the energy requirement for C0 ₂ capturing, the standard wet scrubbing method, using various alkali solutions, requires 30 to 390 kJ/mol of C0 ₂ captured. On the other hand, the use of CaO for C0 ₂ capture requires a regeneration energy estimated at 256.0 kJ/mol C0 ₂ . (taking into consideration the sensible heat required to take the sorbent from 25 to 1000 °C). In the process of the present invention, this energy demand for sorbent regeneration is met by solar energy as described herein.

Calculations by the present inventor show that the energy demand for the process illustrated in Fig. 3 is 230 kJ/mol of H ₂ produced, which can be entirely supplied by solar heating. Conventional H ₂ production via natural gas steam reforming with C0 ₂ capture requires a higher energy demand estimated at 440 kJ/mol of H ₂ produced. Besides, this process is undesirable due to depleting fossil fuel reserves.

All references cited herein are hereby incorporated by reference in their entirety and for all purposes.

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SUBSTITUTE SHEET RULE 26 G. Maag et al, Solar thermal eracking of methane in a particle-flow reactor for the co-production of hydrogen and carbon; Int. J. Hydrogen Energy 34 (209) 7676-7685

Lede, J., Solar Thermochemical Conversion of Biomass; Solar Energy (1999( 65 No.1 3-13

Steinfeld, A., Solar thermochemical production of hydrogen - a review; Solar Energy 78 (2005) 603-615

Hertwich, E. G. and Zang, X., Concentrating-Solar Biomass Gasification Process for a 3 ^rd Generation Biofuel; Environ. Sci. Techno!. 2009 43

4207-4212

Muller-Steinhagen, H. and Trieb, F., Concentrating Solar Power - a review of the technology, Quarterly of the Royal Academy of Engineering Ingenia 18, February/March 2004, p43-50

Muller-Steinhagen, H. and Trieb, F., Concentrating solar power for sustainable electricity generation - Part 2: Perspectives, Ingenia p 35-42

Tasaka et al, Steam Gasification of Cellulose with Cobalt Catalysts in a Fluidised Bed Reactor, Energy & Fuels 2007, 21 , 590-595

Weimer et al, Lime enhanced gasification of solid fuels: Examination of a process for simultaneous hydrogen production and C0 ₂ capture, Fuel 87 (2008) 1678-1686

Nikulshina and Steinfeld, C0 ₂ capture from air via CaO-carbonation using a solar driven fluidised bed reactor - Effect of temperature and water vapour concentration, Chem. Eng. J. 155 (2009) 867-873

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SUBSTITUTE SHEET RULE 26 Nikulshina et al, C0 ₂ capture from air and co-production of H ₂ via the Ca(OH) ₂ -CaC0 ₃ cycle using concentrated solar power - Thermodynamic analysis, Energy 31 (2006) 1379-1389

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SUBSTITUTE SHEET RULE 26