Sorption enhanced methanation of biomass

Abstract

Disclosed embodiments provide a system and method for producing hydrocarbons from biomass. Certain embodiments of the method are particularly useful for producing substitute natural gas from forestry residues. Certain disclosed embodiments of the method convert a biomass feedstock into a product hydrocarbon by hydropyrolysis. Catalytic conversion of the resulting pyrolysis gas to the product hydrocarbon and carbon dioxide occurs in the presence of hydrogen and steam over a CO.sub.2 sorbent with simultaneous generation of the required hydrogen by reaction with steam. A gas separator purifies product methane, while forcing recycle of internally generated hydrogen to obtain high conversion of the biomass feedstock to the desired hydrocarbon product. While methane is a preferred hydrocarbon product, liquid hydrocarbon products also can be delivered.

Claims

1. A method for converting a biomass feedstock into methane, comprising: introducing the biomass feedstock, a transition and/or noble metal catalyst, a sorbent, hydrogen and steam to a hydrogasifier; heating the biomass feedstock, the catalyst, the sorbent, the hydrogen and the steam in the hydrogasifier at a temperature and pressure suitable to form a gas stream comprising the methane, while simultaneously removing carbon dioxide by carbonation of the sorbent; regenerating the catalyst and the sorbent; and separating the methane from the gas stream.

2. The method of claim 1, wherein the hydrogasifier comprises a pyrolyzer and a hydroconverter, and the sorbent and the catalyst are located in the hydroconverter.

3. The method of claim 2, wherein heating the biomass feedstock comprises hydropyrolysis to form a pyrolysis gas and hydroconversion of the pyrolysis gas, the hydropyrolysis and hydroconversion being performed in staged reactors.

4. The method of claim 2, wherein the pyrolyzer is heated at temperature of from 300 C. to 500 C.

5. The method of claim 2, comprising subjecting the biomass feedstock to pyrolysis in the pyrolyzer to form a pyrolysis gas, and converting at least a first portion of the pyrolysis gas by hydroconversion in the hydroconverter that is separate from the pyrolyzer.

6. The method of claim 1, wherein regenerating the sorbent comprises heating the sorbent by combustion of char and any coke deposited on the sorbent or the catalyst.

7. The method of claim 6, wherein regenerating the catalyst comprises regenerating or decoking the catalyst while regenerating the sorbent.

8. The method of claim 1, wherein regenerating the sorbent comprises heating the sorbent with superheated steam.

9. The method of claim 8, wherein regenerating the catalyst comprises regenerating or decoking the catalyst while regenerating the sorbent.

10. The method of claim 1, wherein the sorbent is selected from calcined dolomite, calcium hydroxide, lithium zirconate, lithium orthosilicate, CaO, K-promoted hydrotalcite, K-promoted MgO, K-promoted dolomite, or a combination thereof.

11. The method of claim 1, wherein the catalyst is selected from Ni, Mo, W, Co, Pt, Pd, Ru, Rh, ceria, or a combination thereof.

12. The method of claim 1, where the pressure is from 5 bara to 50 bara.

13. The method of claim 12, wherein the pressure is from 10 bara to 20 bara.

14. The method of claim 1, where regenerating the sorbent occurs at a temperature of from 700 C. to 850 C.

15. The method of claim 1, wherein the gas stream comprises excess hydrogen and the method further comprises separating the excess hydrogen by pressure swing adsorption or membrane permeation.

16. The method of claim 1, wherein regenerating the catalyst and introducing the catalyst comprise regenerating the catalyst and introducing the catalyst using one of the following reactor configurations so that the catalyst will cycle between distinct reaction zones for hydrogasification and regeneration: a. a moving bed with a granular catalyst; b. at least one fixed bed with a granular packing or a monolithic catalyst, the at least one fixed bed having rotary or directional valve logic for cyclically switching the at least one fixed bed between reaction and regeneration; or c. a bubbling fluidized bed or a circulating fluidized bed.

17. The method of claim 1, further comprising generating electrical power with an internal-reforming solid oxide fuel cell fuelled by the methane and hydrogen converted from the biomass feedstock.

18. The method of claim 1, further comprising generating power with a gas turbine to recover heat from regenerating the sorbent.

19. The method of claim 1, further comprising removing the catalyst and the sorbent from the hydrogasifier prior to regenerating the catalyst and the sorbent.

20. The method of claim 1, wherein the hydroconverter is heated at temperature of from 200 C. to 650 C.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

(1) FIG. 1 is a schematic diagram of one embodiment of an apparatus according to the present invention.

(2) FIG. 2 shows an embodiment with integration to a gas turbine power plant.

(3) FIGS. 3 and 4 show an embodiment with a rotary reactor including cyclically switched hydrogasification and regeneration zones.

(4) FIG. 5 shows a fluidized twin bed embodiment of the invention.

(5) FIG. 6 shows a two stage double twin bed fluidized embodiment.

(6) FIG. 7 shows a two stage embodiment, with the first stage a fluidized bed hydropyrolysis reactor and the second stage a methanation or hydroconversion reactor with a cyclic rotary switching mechanism.

(7) FIG. 8 shows an embodiment for coproduction of methane and liquid hydrocarbons.

(8) FIG. 9 shows an embodiment with a hydrogasification system coupled to a solid oxide fuel cell (SOFC) for generation of electricity.

(9) FIG. 10 is a graph of methane conversion, methane concentration, hydrogen concentration, and the ratio of hydrogen output to hydrogen input from the hydrogasification reactor of the invention, versus the ratio of carbon carbonated on the sorbent to carbon content of the biomass feed to the process.

DETAILED DESCRIPTION

(10) FIG. 1 shows a simplified schematic of a system or apparatus 1 according to the invention. Apparatus 1 includes a feed preparation section 2, a feed pressurization section 3, a hydrogasification reactor 4, a solids separation section 5, a regeneration reactor 7, a heat recovery section 8, a tar scrubber 9, and a gas separation section 10. The gas separation section 10 separates hydrogen from methane, and may be based on selective membranes or on pressure swing adsorption (PSA). Gas separation section 10 delivers a substantially purified methane product stream from conduit 11, and a hydrogen-enriched recycle stream back to hydrogasification reactor 4 via conduit 12. It is a key feature in some preferred embodiments of the present invention that the gas separation system 10 can be operated so as to minimize hydrogen concentration in the product methane stream, thus preventing most of the hydrogen from exiting apparatus 1 until consumed in hydrogasification to produce the desired methane product.

(11) A membrane permeation system may be advantageously used as gas separation system 10 for purification of product methane and separation of hydrogen-rich gas for recycle to hydrogasification. The polymeric membrane will selectively permeate hydrogen, carbon monoxide, carbon dioxide and water vapour relative to methane. In order to obtain product methane containing no more than 1% hydrogen, three membrane stages may be used in series to progressively concentrate methane with high recovery and purity in the retentate stream. The feed gas is introduced to the inlet of the first stage, from which the hydrogen-enriched recycle gas will be delivered as low pressure permeate. The permeate of the second stage will be recompressed to join the feed at the inlet of the first stage, while the permeate of the third stage will be recompressed to the inlet of the second stage.

(12) Alternatively, pressure swing adsorption may be used as gas separation system 10 for purification of product methane and separation of hydrogen-rich gas for recycle to hydrogasification.

(13) Hydrogasification reactor 4 and regeneration reactor 7 comprise a coupled reactor pair for the working hydrogasification reaction and regeneration steps. Each of reactors 4 and 7 comprises at least one bed containing a CO.sub.2 sorbent, and optionally also another solid component having catalytic and favourable heat transfer media characteristics. It is contemplated that the beds will cycle or be switched between the hydrogasification and regeneration reaction zones.

(14) Pressure and temperature conditions in the hydrogasification reactor 4 will be selected to be favourable for combined methanation and steam reforming reactions such that the process achieves self-sufficiency in producing the amount of hydrogen needed for the methanation reaction, while also favourable for the carbonation reaction binding CO.sub.2 to the sorbent. The regeneration reactor 7 will typically be operated at relatively higher temperature in order to release the CO.sub.2 from the sorbent, and may be operated at substantially the same pressure or at lower pressure to facilitate regeneration and reduce air compression requirements for regeneration.

(15) A preferred CO.sub.2 sorbent is CaO, alternatively provided by calcining of natural limestone or dolomite, or by synthesis to achieve enhanced mesoporosity and stability for extended cycling without rapid deactivation due to sintering and pore blockage. It has been found in the art that composite structures comprising CaO supported on or encapsulated in a mesoporous ceramic (e.g. alumina) may achieve superior durability against deactivation, while also providing a hardened external shell for superior attrition resistance. Such mesoporous composite structures may be usefully applied in fixed bed monoliths as well as in granular media for fluidized bed operations. The present invention contemplates the use of zirconia and alumina for such mesoporous composites with CaO.

(16) At relatively higher pressures around 30 bara and with high steam concentrations during sorbent processing, Ca(OH).sub.2 may be a useful intermediate between CaO and CaCO.sub.3. It is known that the sintering longevity problems in cycling between CaO and CaCO.sub.3 may be largely avoided by cycling between oxide, hydroxide and carbonate compounds of the sorbent.

(17) The regeneration gas for regeneration reactor 7 provides heat for calcining the sorbent, and oxygen and/or steam for decoking the catalyst, while serving as sweep gas to purge CO.sub.2 released in calcining and decoking functions. A preheater 15 is provided to preheat air or enriched oxygen provided for oxidation and sweep gas functions, while superheating any steam provided to assist decoking and as sweep gas. A feed air compression unit 16 is provided, which would usually be an air compressor. Feed air compression unit 16 may alternatively include a feed air blower, a pressure swing adsorption oxygen enrichment unit, and an oxygen compressor delivering oxygen to preheater 15. Oxygen enrichment may be desired to reduce air compression loads in higher working pressure embodiments of the process, or to facilitate capture of concentrated CO.sub.2 from the exhaust. A heat recovery section 18 is provided to recover heat from the CO.sub.2 and sweep gas discharged to exhaust.

(18) Heat recovery sections 8 and 18 are here contemplated to be steam generators, providing steam by conduit 19 to hydrogasification reactor 4 and optionally in some embodiments also by conduit 20 to preheater 15 and regeneration reactor 7.

(19) Feed preparation section 2 includes steps of sizing and drying as necessary. Feed pressurization section 3 includes a lock hopper system or a pressure feeder device to introduce the feed biomass into the pyrolysis and gasification process at a working pressure of preferably about 5 bara to about 50 bara, and more preferably about 10 bara to about 20 bara.

(20) Solids are removed from the effluent pyrolysis gas exiting hydrogasification reactor 4 by a solids removal section 5 including one or multiple cyclones, and optionally also high temperature filters such as metallic or ceramic candle filters. A desulfurization reactor (e.g. using zinc oxide for H.sub.2S removal), or sorbent beds for removal of alkalis or chlorides, may be included here for protection of any downstream catalysts.

(21) FIG. 2 shows a simplified schematic of an alternative apparatus 40 with the regeneration reactor 7 embedded in a gas turbine 41 to enable the efficient use of pressurized air for regeneration. Gas turbine 41 includes a compressor 42 coupled to a high temperature expander 43 and a mechanical load 44 which may be an electrical generator. The expanded exhaust from expander 43 includes vitiated air and CO.sub.2 released from the sorbent, and is subjected to heat recovery in heat exchanger 47 before generating steam in downstream heat recovery section 18. Heat exchanger 47 provides heat to preheater 15 in this embodiment, and may be directly integrated with preheater 15 in well known recuperator or rotary regenerator embodiments well known in the gas turbine art.

(22) In order to protect the turbine blades of expander 43 from erosion, corrosion or fouling damage, it is necessary to provide a hot gas clean-up section 50 to remove solid particulate, alkalis and any chloride or sulphur compounds that have not been retained by the sorbent under regeneration conditions. The hot gas clean-up section 50 will include cyclones, filters (metal fabric, ceramic candles or precoat filters), and chemical sorbents as necessary to capture the alkalis and any other detrimental components. Captured solids and spent sorbents will be released from discharge conduit 51.

(23) FIGS. 3 and 4 show an embodiment 60 with rotary fixed bed reactors that will be particularly useful for smaller scale applications of the invention, e.g. for local supply of methane fuel from forestry operations in remote areas.

(24) Reactors 4 and 7 are combined in a cyclic rotary reactor 61. A plurality of fixed beds 62 are mounted in a rotor 63 rotating about rotary axis 64 between a first rotary valve face 65 and a second rotary valve face 66. First rotary valve face 65 engages sealingly with a first valve stator face 67, and second rotary valve face 66 engages sealingly with a second valve stator face 68. Fluid connection ports 75, 76, 77 and 78 are provided in first valve stator face 67, while fluid connection ports 85, 86, 87 and 88 are provided in second valve stator face 68. FIG. 4 shows the ports for the first valve stator face, with bracketed reference numerals for the corresponding ports of the second valve stator face.

(25) The biomass feed is pressurized and decomposed by pyrolysis reactor 70 before admission to port 75. Conduits 12 and 19 respectively provide recycle hydrogen and steam to port 75. Raw product methane gas is delivered from port 85 to heat recovery, clean-up and purifications steps.

(26) Preheated regeneration air is introduced to port 77, while the CO.sub.2 containing exhaust is discharged from port 87 to heat recovery. Cocurrent regeneration as shown in FIG. 3 may be replaced with countercurrent regeneration by introducing the preheated regeneration air to port 87, and discharging the exhaust from port 77.

(27) Intermediate ports 76 and 78 in the first stator, and intermediate ports 86 and 88 in the second stator, are provided to enable buffer purge steps with steam or other inert gas between the hydrogasification and regenerations steps, so as to avoid hazardous direct contact of undiluted air with high concentration fuel gas. The intermediate ports may also be used for pressure equalization steps between the hydrogasification step performed at elevated pressure and the regeneration step performed at lower pressure or substantially atmospheric pressure.

(28) FIG. 5 shows a circulated fluidized twin bed embodiment 100 of the invention that would be particularly applicable to larger scale installations. Fluidized bed loop 101 includes a bubbling bed hydrogasification reactor 4 energized by recycle hydrogen from conduit 12 and steam from conduit 19, and a circulating fluidized bed regenerator reactor 7 energized by feed air from compressor 42.

(29) The fluidized bed solid media includes CaO sorbent, preferably formed in composite mesoporous ceramic pellets. The ceramic (preferably also impregnated with transition group metal catalysts such as nickel, or noble metal catalysts such as rhodium with ceria) may itself have catalytic properties for reforming of pyrolysis gas and tars, and for methanation of syngas.

(30) The granular media may be a mixture of sorbent, catalyst and heat exchange particles. The media should have high heat capacity, thermal conductivity and attrition resistance. Olivine sand is recognized as having excellent properties as heat transfer media in biomass gasification, including moderate catalytic properties for reforming tar constituents. Magnetite may also be useful as heat transfer media, with the potential advantage of downstream magnetic separation between the heat transfer media and char.

(31) Carbonated sorbent, char and coked catalyst are transferred from hydrogasification reactor 4 to regenerator reactor 7 via siphon 102. Calcined sorbent and regenerated catalyst are transferred back from regenerator reactor 7 to hydrogasification reactor 4 via cyclone 103 and siphon 104. Ash may be released from the bottom of regenerator 7 by a lock hopper or an intermittently operated valve.

(32) FIG. 6 shows a two stage fluidized bed embodiment 150, with the first stage 101 using relatively robust but less catalytically active media such as olivine or nickel impregnated olivine or glass-ceramic transition metal (e.g. Ni, Mo, W and combinations thereof) catalysts such as developed by the Gas Technology Institute, and the second stage 201 using more active and more delicate catalysts and lime sorbent for improved conversion of tars and higher yield of methane. In this embodiment, the first stage 101 is a fluidized bed reactor loop achieving hydropyrolysis and partial conversion to methane and hydrogen, while the second stage 201 performs more complete hydroconversion.

(33) The solids separation section 5 may here include means to remove catalyst poisons (e.g. sulphur, chlorides, alkalis, etc.). Second stage fluidized bed loop 201 includes the second stage hydrogasification reactor 204, the second stage regeneration reactor 207. Deactivated catalyst and carbonate sorbent are transferred from hydrogasification reactor 204 to regenerator reactor 207 via siphon 212. Calcined sorbent and regenerated catalyst are transferred back from regenerator reactor 207 to hydrogasification reactor 204 via siphon 214.

(34) The hot gas effluent from regenerator reactor 207 is delivered through cyclone 213 and hot-gas cleanup section 50 to the inlet of gas turbine expander 43.

(35) Fluid control means 215 is provided to control flows of recycle hydrogen from conduit 12 and steam from conduit 19 to energize fluidized beds in first stage reactor 101 and second stage reactor 201. Fluid control means 215 may include control valves, expanders or compressors as needed to control flows and regulate pressures.

(36) FIG. 7 shows a two stage fluidized bed embodiment 250, with the first stage 101 a fluidized bed hydropyrolysis reactor as in embodiments 100 or 150, and the second stage a methanation or hydroconversion reactor 61 with a cyclic rotary switching mechanism for rotary fixed beds 62 as shown in embodiment 60. Beds 62 are switched cyclically between methanation or hydroconversion steps as reactor 254, and regeneration steps as reactor 257. Hot gas cleanup can be performed in cleanup section 250 with filtered solids and spent sorbents removed by discharge conduit 251. Heat recovery steam generator 258 may be provided to recover heat between first stage hydropyrolysis reactor 4 and second stage methanation or hydroconversion reactor 254.

(37) FIG. 8 shows an embodiment 300 for coproduction of methane and liquid hydrocarbons with a three stage reactor system. A hydropyrolysis reactor 304 is provided upstream of a hydroconversion reactor 310 for production of liquid hydrocarbons, and itself upstream of a sorption enhanced reactor 314 for production of hydrogen and methane. The hydrogen from reactor is delivered to hydropyrolysis reactor 304. Reactor 304 provides rapid heating of the biomass particles to a pyrolysis temperature in the range from about 300 C. to about 500 C., in order to decompose the biomass into pyrolysis gas (including light hydrocarbons, some syngas, and tar vapours) and char. Heating may be achieved by mixing the biomass particles with a granular heat transfer media in a mechanical or fluidized bed contacting system. Various mechanisms are well known for fast pyrolysis reactors (e.g. auger reactors and circulating fluidized beds), and may be used in a mechanical contacting system in reactor 304. The granular heat transfer media should have high heat capacity, thermal conductivity and attrition resistance. Olivine sand is recognized as having excellent properties as heat transfer media in biomass gasification, including moderate catalytic properties for reforming tar constituents. Nickel-impregnated olivine has improved catalytic properties. The glass-ceramic catalysts developed by the Gas Technology Institute are believed to be superior for hydropyrolysis applications. Magnetite may also be useful as heat transfer media, with the potential advantage of downstream magnetic separation between the heat transfer media and char.

(38) The heat exchange media is circulated between reactor 304 and a media heater 315, with pyrolytic char being discharged from reactor 304 with spent heat exchange media returning to the media heater 315. Combustion of char in media heater 315 may conveniently provide heat required for heating the feed biomass to reaction temperature and for the endothermic pyrolysis and initial gasification reactions. Ash is discharged from media heater 315.

(39) A portion of the char exiting reactor 304 may be separated from the heat exchange media by char separator 316 as the feedstock for an auxiliary oxygen or steam gasification method to generate syngas. After water gas shift and CO.sub.2 removal from the syngas, supplemental hydrogen may thereby be provided for the subsequent hydrogasification reaction. Alternatively a portion of the char separated by char separator 316 may be diverted to other external uses, including sale of charcoal as a solid fuel, or as a bio-char soil amendment for agriculture or forestry uses with an important purpose of carbon sequestration in the soil. Ash may also be a useful byproduct for soil enhancement and recycle of nutrients for overall sustainability of biomass cultivation, harvesting and utilization.

(40) Solids are removed from the effluent pyrolysis gas exiting pyrolysis reactor 304 by a solids removal section 317 including one or multiple cyclones, and optionally also high temperature filters such as metallic or ceramic candle filters. A catalyst poison removal section 318 (including a desulfurization reactor using zinc oxide for H.sub.2S removal, and optionally including other sorbent beds for removal of chlorides and/or alkalis) may be included here for protection of downstream catalysts. The pyrolysis gas is also cooled by a heat recovery steam generator 319, either upstream or downstream of the catalyst poison removal sorbent beds.

(41) The cooled pyrolysis gas is introduced to catalytic hydroconversion reactor 310, together with hydrogen (or hydrogen-rich gas) and optionally also with steam. Hydrogen reactively deoxygenates the pyrolysis gas components to generate a mixture of lighter and heavier hydrocarbons by hydrodeoxygenation and decarboxylation reactions. Hydrogen and steam act to crack larger molecules, and to inhibit coking. The reactor effluent is provided to a first separator 321 from which a liquid fraction of heavier hydrocarbons is delivered by conduit 322 as a first liquid hydrocarbon product for further processing and use as desired.

(42) The overhead fraction from first separator 321 is cooled by heat recovery unit 330 to generate steam or preheat water upstream of second separator 331 in which water is condensed and separated from a liquid fraction of gasoline range hydrocarbons which is delivered as a second liquid hydrocarbon product by conduit 332 for further processing and use as desired. The overhead fraction of gases and vapours from second separator 331 contains H.sub.2, CO, CO.sub.2, CH.sub.4 and other light hydrocarbons along with some water vapour. This fraction is reheated and admitted to sorption enhanced reactor 314, optionally together with a portion of the cleaned pyrolysis gas from catalyst poison removal section 318 as controlled by valve 340 in conduit 341.

(43) Carbonation of CaO in sorption enhanced reactor 314 removes CO.sub.2 and also CO by water gas shift. Light hydrocarbons are preferentially prereformed so that the product of sorption enhanced reactor 314 will be mostly hydrogen with methane as the main residual carbon-containing compound. After clean-up and cooling of this product gas mixture from reactor 314, gas separation system 10 separates substantially purified methane into SNG product delivery conduit 11, and a hydrogen-enriched recycle stream into conduit 12 and back to the hydropyrolysis reactor 304.

(44) The gas separation system 10 may be operated to deliver a sufficient amount of hydrogen to the hydropyrolysis reactor and a significant amount of methane, with more methane production feasible if less liquid hydrocarbons are produced. In the absence of supplemental hydrogen imported to the process, approximately half the carbon in the feed biomass may be converted to product hydrocarbons (including heavier liquid hydrocarbons in the first product, gasoline range hydrocarbons in the second product, and product methane as a third product). The balance of the carbon is discharged as char or CO.sub.2. Higher carbon conversion can be achieved with the addition of imported hydrogen.

(45) The product split between methane and liquid hydrocarbons can be varied operationally by adjusting gas separation unit 10 so that more or less methane is delivered. With lower production of methane, more recycle hydrogen is available to the hydropyrolysis and hydroconversion reactors so that more liquid hydrocarbons are produced.

(46) Production of liquid hydrocarbons can be maximized by turning off the delivery of product methane. In one embodiment of the invention, the gas separation system 10 is removed so that the entire product effluent of sorption enhanced reactor 314 is delivered to hydropyrolysis reactor 304, while liquid hydrocarbon delivery of the first and second products is accordingly augmented. Reactor 314 is then operating as a sorption enhanced steam reformer, with several advantages including (1) lower temperature operation than a conventional steam reformer, (2) convenient direct use of char combustion for regenerating the sorbent, (3) integrated water gas shift and CO.sub.2 removal, and (4) scavenging of any alkali and chloride impurities in the recycle gas by the lime sorbent.

(47) Regenerator reactor 7 and sorption enhanced reactor 314 comprise a coupled reactor pair for the working reaction and for regeneration of the sorbent and catalyst. A catalyst regeneration reactor 330 is also provided for decoking catalyst from hydroconversion reactor 310.

(48) A portion of compressed air from compressor 42 and preheater 15 is provided to pyrolysis media heater 315 for combustion of char to heat the media, with flue gas heat being recovered in steam. The remainder of the compressed air from compressor 42 and preheater 15 is provided optionally with steam to regeneration reactors 7 and 330 to burn coke off the catalysts and decarbonate the CaCO.sub.3 formed in sorption enhanced reactor 314. Heat recovery steam generators 8, 319 and 350 deliver steam to sorption enhanced reactor 314, or to regeneration reactors 330 and 7 as needed.

(49) FIG. 9 shows an embodiment 400, in which a hydrogasification system similar to embodiment 1 is coupled to a solid oxide fuel cell (SOFC) 410 for generation of electricity. The SOFC has a solid oxide electrolyte 412 between a cathode chamber 413 and an anode chamber 414. The cathode chamber has inlet port 415 and outlet port 416, and the anode chamber has inlet port 417 and outlet port 418. The anode chamber includes an internal reforming catalyst (which may be comprised within the anode itself) to convert methane and steam to hydrogen and carbon oxides CO and CO.sub.2 under SOFC operating conditions at elevated temperatures in the range of about 600 C. to about 800 C.

(50) A portion of the compressed air (or compressed oxygen-enriched air) from air compression unit 16 is heated in thermal recuperator 422, and fed to cathode inlet port 415. Vitiated cathode gas is discharged from port 416 and exhausted through recuperator 420 which recovers sensible heat from this gas.

(51) A hot gas clean-up section 430 is provided to remove solid particulate, alkalis and any chloride or sulphur compounds that have not been captured by the sorbent. The hot gas clean-up section 430 may include cyclones, filters (metal fabric, ceramic candles or precoat filters), and chemical sorbents as necessary to capture the alkalis and any other detrimental components. Captured solids and spent sorbents will be released from discharge conduit 431.

(52) FIG. 10 is a graph of methane conversion, methane concentration, hydrogen concentration, and the ratio of hydrogen output to hydrogen input (H.sub.2out/H.sub.2in) from the hydrogasification reactor of the invention, versus the ratio of carbon carbonated on the sorbent to carbon content of the biomass feed to the process.

(53) These correlations were derived for the case of an AEM process operating with CaO sorbent at a temperature of 600 C. and a pressure of 10 bara, with 1.0 molecules of hydrogen and 0.6 molecules of water vapour (including initial water content within the biomass) provided per atom of carbon in the original woody biomass feed. Thermodynamic equilibrium was assumed for the water gas shift, steam reforming and methanation reactions, with methane the only hydrocarbon molecule participating in these post-pyrolysis reactions. Char and coke deposition was assumed to consume about 21% of feed biomass carbon.

(54) Modeling runs were performed for different values of sorption uptake of carbon, shown as fractional carbonation of the original biomass carbon. Maximum fractional carbonation was found to be 0.4491, at which condition the partial pressure of CO.sub.2 is at equilibrium with a mixture of CaO and CaCO.sub.3. At a fractional carbonation of 0.370, the hydrogen output is equal to the amount of hydrogen input, thus defining an ideal self-sustaining condition without excess hydrogen generated or any external supplemental supply of hydrogen. With allowance for imperfect separation of product methane and recycle hydrogen, the practicable operating condition for self-sustaining hydrogen generation (without supplemental hydrogen supply from any external source) in this example will require fractional carbonation greater than 0.37, and of the order of 0.4.

(55) The above example shows that high purity methane can be produced by the sorption enhanced methanation method according to the invention, with methane the predominantly surviving carbon compound after nearly complete removal of CO.sub.2 and CO. This example contrasts dramatically with the related and well known process of sorption enhanced steam reforming of methane, where methane is nearly completely extinguished along with CO.sub.2 and CO in order to achieve hydrogen production with highest possible conversion.

INDUSTRIAL APPLICABILITY

(56) Disclosed embodiments of the method and system are useful for high efficiency conversion of biomass, including forestry residues (including those generated by logging, thinning, and wildfire prevention fuel load reduction activities) and sawmill waste into SNG, either as a fuel commodity or for high efficiency generation of electrical power. Disclosed embodiments provide advantageous integrations with gas turbines and/or solid oxide fuel cells. A portion of the biomass may also be converted into heavier and lighter hydrocarbon liquids.

(57) Disclosed embodiments of the system may be used at industrial scale limited only by transportation distances for collection of biomass feedstock, or at smaller scale in rural or remote areas for combined generation of heat, high heating value fuel gas and electricity. At the smallest scale, the system may be used for residential heating, methane fuel production and electrical power generation through a solid oxide fuel cell or other energy converter consuming a portion of the product methane.

(58) In view of the many possible embodiments to which the principles of the disclosed invention may be applied, it should be recognized that the illustrated embodiments are only preferred examples of the invention and should not be taken as limiting the scope of the invention. Rather, the scope of the invention is defined by the following claims. We therefore claim as our invention all that comes within the scope and spirit of these claims.