Patent Description:
This is a process to manufacture a syngas from CaCO<NUM> and H<NUM>. When the source of CaCO<NUM> is a natural limestone or other mineral containing CaCO<NUM> at ambient temperature, the process is applicable to shaft kilns, flash calciners or circulating fluidised bed calciners. When the source of CaCO<NUM> is a process involving a step of carbonation of CaO, the process is applicable to calcium looping processes to capture CO<NUM> from gases.

The future deployment of renewable electricity at large scale and low cost will make electrolytic H<NUM> also available at low cost. In this context, there is growing interest in the development of industrial processes to manufacture C-based synthetic fuels from CO<NUM> and H<NUM>, "Power to Liquid" or PtL processes, preferably using renewable CO<NUM> (i.e., captured from the atmosphere or from biogenic sources) to have a "close to zero" carbon footprint. A recent review of current commercial PtL processes is given by <NPL>). Current commercial PtL processes involve a step of reverse water gas shift, RWGS, to obtain a syngas containing CO, H<NUM>, H<NUM>O(v) and some unconverted CO<NUM>. Solid catalyst containing Cu or Fe are used to accelerate the RWGS reaction. There are strict purity requirements for the CO<NUM> feedstock fed to these PtL processes (see for example Table <NUM> in Dieterich et al. ), usually because of the need to preserve the activity of the RWGS catalyst.

To overcome the need of high purity levels in the CO<NUM> feed to the PtL process, and to achieve more energy efficient and intensive processes, there is also substantial R&D in the development of direct routes of hydrogenation of metal carbonates (<NPL>). Background on the direct hydrogenation of CaCO<NUM> goes back to <CIT>) that disclosed a batch hydrogenation method of CaCO<NUM>, involving the heating up of CaCO<NUM> (to at least 200ºC) in a pressurised vessel to contact H<NUM> (at pressures up to <NUM> atm) during <NUM>-<NUM> hours and produce fuel gases (methane, ethane, propane etc). In a different publication, co-workers of the author (<NPL>) reported the formation of CO in similar experiments at 660ºC, with an overwhelming presence of H<NUM> (><NUM>%v) in the gas mixture and a complete conversion of the small sample of CaCO<NUM> used in all the experiments. They also refer to additional experiments, not reported in <CIT>, where the hydrogenation process of CaCO<NUM> was satisfactorily "carried out with hydrogen pressures only slightly above ambient. Thus, the hydrogen needs only to be at a pressure sufficient to insure intimate contact of the hydrogen gas with" CaCO<NUM>. More recently, Lux et al stated that, "the admixture of transition metals to main-group metal carbonates opens up a new pathway in metal carbonate hydrogenation, because many transition metals are known to catalyze hydrogenation reactions, CO<NUM> evolved from the carbonate is converted into CO, CH<NUM>, or directly to higher hydrocarbons CxHy and CxHyOz". <NPL>) first reported hydrogenation test with mixtures of CaCO<NUM> and transition metals, that seemed to act as catalyst of the hydrogenation reactions of CaCO<NUM> (with CO and CH<NUM> being the main hydrocarbons product). They observed that the calcination of CaCO<NUM> was taking place at comparably low temperatures, with this effect enhanced by the admixture at close to atomic level of transition metals with the CaCO<NUM>. In a recent review of this phenomena by <NPL>) a wide range of dual functional materials, combining CaO as CO<NUM> sorbent and a transition metal as hydrogenation catalyst, are referred. Such DFM have promising properties to operate in a wide range of process concepts, as reviewed by S. (<NUM>) noted above. These processes usually involve means to circulate solids between a carbonator (where CO<NUM> contained in a gas reacts with CaO to form CaCO<NUM>) to a hydrogenator (where the CaCO<NUM> will react with H<NUM> to form CaO and a gas containing hydrocarbons). Alternatively, these processes have been proposed to operate following principles of pressure and/or temperature swing adsorption, alternating between carbonation and hydrogenation conditions in the same vessel (see for example <FIG> in <NPL>).

Relevant for this invention is to note that Cu-based catalyst and Fe-based catalyst are within the list materials with catalytic properties for reverse water gas shift reactions. Also, that these Cu-based and Fe-based materials are known to have excellent properties as oxygen carrier materials in chemical looping combustion applications, where the metal oxides (CuO or FeO/Fe<NUM>O<NUM>/Fe<NUM>O<NUM>) can burn a fuel gas into CO<NUM> and H<NUM>O while reducing to Cu or Fe or a lower oxidation form of the iron oxide. In a subsequent step, the metals or reduced forms of the iron oxides can be exothermically oxidised again in contact with air or other gas containing O<NUM>. Also relevant for this invention are known systems that combine chemical looping combustion steps to provide the energy needed to drive CaCO<NUM> calcination steps in calcium looping processes, where carbonation of CaO occurs in contact with a gas containing CO<NUM> followed by the calcination of CaCO<NUM> to regenerate CaO and obtain pure CO<NUM>. <CIT> discloses one of such systems, where the generation of CaO in the bed during the calcination step allows for a subsequent step of CaO-based Sorption Enhanced Reforming and CaO-based Sorption Enhanced Water Gas Shift reactions, to produce H<NUM> from C-fuels. Note that these enhanced reaction conditions as such as to promote the removal of CO<NUM> from the gas phase by carbonation of CaO to form CaCO<NUM> while undertaking reactions that generate water gas shift gases. Following Le Chatelier's principle, by removing CO<NUM> from the gas phase, the equilibrium of the water gas shift reaction of CO and water vapour, H<NUM>Ov, is displaced towards the formation of H<NUM>. A more recent example of these system is the so called CASOH process, where a bed of solids containing CaO first removes the CO<NUM> contained in the BFG and promotes the enhanced water gas shift of the CO by capturing the resulting CO<NUM> as CaCO<NUM>. In a subsequent step, CaO must be regenerated, burning a fuel gas with a chemical loop of Cu/CuO to produce CO<NUM> and water vapour (<NPL>).

Most relevant for this invention, <NPL>) have recently reported the catalytic effect of CaO particles in RWGS experiments ("the conversion approaches equilibrium for many experimental conditions in the presence of CaO"). They provide experimental evidence (see for example Figure A6 in their supplementary information material) of a sharp increase in CO concentration in the product gas when the CaCO<NUM> containing solids exceeded the calcination temperature (given by the equilibrium of CO<NUM> on CaO at the CO<NUM> partial pressure of the feed gas), consistent with a mechanism involving R1+R3 in their Scheme <NUM>, at the temperatures tested. Their equation <NUM>, using the parameters listed in their Table <NUM> as "CaO 0W" provides a useful kinetic equation to fit their experimental results of RWGS reaction in the presence of CaO as catalyst, when combined with known expressions in the state of the art for the equilibrium constant of CO<NUM> on CaO (for example with the equilibrium equation of Hills) or for the water gas shift, WGS, equilibrium (for example the equilibrium WGS equation of Twigg).

Despite the progress in the science of direct hydrogenation of CaCO<NUM> containing materials to CO and H<NUM>Ov, the recent confirmation of the role of CaO as a reverse water gas shift catalyst, and the availability of dual functional materials containing CaCO<NUM> and transition metals where the direct hydrogenation of CaCO<NUM> is shown to occur at lower temperatures than the calcination temperatures, there is lack of methods to undertake these processes at industrial scale. In particular, there is a need to solve with viable methods, and without emitting CO<NUM>, the large energy requirements for the endothermic hydrogenation of CaCO<NUM> containing solids (note that considering an enthalpy of <NUM> kJ/mol for the endothermic calcination of CaCO<NUM> and a RWGS reaction enthalpy of <NUM> kJ/mol, the enthalpy of hydrogenation of CaCO<NUM> to CO and H<NUM>O is about <NUM> kJ/mol). All reported hydrogenation experiments of CaCO<NUM> use small non-adiabatic set ups, where the sample solids are heated from a surrounding oven (usually an electric oven). This is not a practical solution for large scale systems designed to supply syngas to PtL processes because heat transfer by thermal conduction in a packed or moving bed of solids is very inefficient due to the low thermal conductivities typical in packed beds of solids. Indeed, combustion of a fuel within a bed of solids is the common practice to supply similar flows of thermal energy for similarly endothermic reactions (such as calcination of CaCO<NUM> containing solids). However, such methods of combustion would translate into unacceptable levels of flue gas emissions with CO<NUM> and the decomposition and loss of the CaCO<NUM> intended for use in direct hydrogenation reaction. It is therefore important for the development of hydrogenation processes of CaCO<NUM> at large scale to conceive methods to transfer the necessary heat from a combustion reaction to the hydrogenating CaCO<NUM> solids, while limiting, or virtually avoiding, the emissions of CO<NUM> to the atmosphere from the combustion or from the decomposition of CaCO<NUM> before hydrogenation takes place.

On the other hand, in view of the growing interest and demand for syngas with H<NUM>/CO molar ratios close to <NUM> for commercial PtL processes, it is important to define methods that target the energy efficient production of such targeted syn-gas H<NUM>/CO ratio close to <NUM>, while achieving maximum product yields of CO and minimum contents of the CO<NUM> in the product gas (i.e., maximum CO<NUM> conversion in the reverse water gas shift reaction).

All referred papers and patents in the field of direct hydrogenation of CaCO<NUM> are silent on the previous important problems. As a result, no full-scale method for direct hydrogenation of CaCO<NUM> (i.e., providing a practical solution to the energy supply for the hydrogenation reaction while maintaining low or "close to zero" emissions of CO<NUM>) has been disclosed. Indeed, large scale integrated processes for PtL using CaCO<NUM> as a source of carbon, still rely on the use of the concentrated CO<NUM> gas resulting from a prior calcination step of CaCO<NUM>. For example, <CIT> relates to a system and process for producing liquid hydrocarbons from a calcium carbonate feed-stock involving in its first step the liberation of carbon dioxide gas by supplying heat to a CaCO<NUM> calciner. The high energy requirements in such step of endothermic calcination of CaCO<NUM> to produce CaO and concentrated, or even pure CO<NUM>, has generated a variety of patented calcination methods in the state of the art. Several methods for such CaCO<NUM> calcination step involve the burning of a fuel gas in the vicinity of the calcining solids, using comburent mixtures of air enriched in O<NUM> or even N<NUM>-free CO<NUM>/O<NUM> mixtures. Such oxy-combustion CaCO<NUM> calcination systems have been developed and integrated in a range of cement manufacturing processes, lime shaft kilns and calcium looping processes (including post-combustion CO<NUM> capture processes and sorption enhanced reforming and water gas shift processes), involving moving bed, fluidised bed or entrained bed calcination technologies, but none of them have been conceived or are applicable under the reducing conditions in a reactor where hydrogenation of CaCO<NUM> with H<NUM> is taking place.

From the previous review of the state of the art, it can be concluded that there is still a need of a method to obtain syngas from CaCO<NUM> and H<NUM> solving the above-cited drawbacks.

The main object of the present invention is to provide a method of hydrogenation and calcination of CaCO<NUM> with H<NUM>, to produce CaO and a gas containing at least H<NUM>, CO, CO<NUM> and H<NUM>O(v), wherein the CO<NUM> is released from the calcination of solids containing CaCO<NUM> when reacted with H<NUM>. The syngas product will preferably have a H<NUM>/CO molar ratio close to <NUM> to facilitate its use downstream in synthesis processes.

The method is applicable to processes for calcination of limestone that use thermally insulated vessels such as moving bed shaft kilns, fluidised beds or flash calciners with a continuous feed of limestone and a continuous discharge of lime. These thermally insulated vessel have at least one calcination section of the vessel, where the solids reach average calcination temperatures (typically higher for example than 900ºC in the calcination section of a lime shaft kiln). Preheating of the limestone before reaching the calcination zone and cooling of the lime before discharging it from the calciner are part of the state of the art to minimise energy requirements in calciners. The method is also applicable to calcium looping processes using thermally insulated vessels containing CaCO<NUM>. In this case the CaCO<NUM> is formed during the capture of CO<NUM> by carbonation of CaO in contact with a combustion flue gas or with any other gas containing CO<NUM>.

The present invention discloses a method of producing a syngas from CaCO<NUM> and H<NUM> wherein the calcination of CaCO<NUM> in carried out in a thermally insulated vessel by reacting it with H<NUM>, to produce CaO and a syngas containing at least H<NUM>, CO, H<NUM>O and CO<NUM>, characterised by a sequence of at least the following consecutive steps carried out at least twice:.

Once that Tmax has been reached, the second step in the method of this invention ii) is a step of reacting with H<NUM> the heated bed of solids containing at least CaO and CaCO<NUM>, while allowing their adiabatic cooling by <NUM> to 250ºC. As shown in dedicated experiments described below, reverse water gas shift of the CO<NUM> evolved from calcination of CaCO<NUM> will take place in these conditions, catalysed by the CaO that is generated from CaCO<NUM> calcination and/or the presence of a RWGS catalyst mixed with the CaCO<NUM> containing solids, such as Cu or Fe-based catalyst. Since the vessel has been chosen to be thermally insulated (i.e., close to adiabatic when the reactor has the scale of a lime kiln), the only energy available to drive the hydrogenation of CaCO<NUM> (requiring <NUM> kJ/mol) is the sensitive heat contained in the CaCO<NUM> containing solids, in the CaO and in any other solids (i.e., the RWGS catalyst) that may be in their proximity, after their heating up during step i). Adiabatic cooling of such solids will therefore take place, as the conversion of CaCO<NUM> to CaO progresses with time. The interval of temperatures allowed for this adiabatic cooling is between <NUM> to 250ºC. This has been defined after modelling the phenomena of "Desorption Enhanced Reverse Water Gas Shift", DERWaGS, which is the reverse of the known Ca-based Sorption Enhanced Water Gas Shift, as described in the detailed description below that follows Le Chatelier principles. As will be shown below, under the chosen reaction conditions, the DERWaGS phenomena cause a continuous fast release of CO<NUM> from the calcining CaCO<NUM> containing solids ensuring a steady supply of CO<NUM> to the local gas atmosphere of the calciner, where RWGS reaction is taking place. Such a steady concentration of CO<NUM> will be maintained by the excess of CaCO<NUM> in the vessel during this step, enhancing the formation of CO in the gas phase. Critically, as long as there is some excess of CaCO<NUM> in the vessel during step ii), certain operating windows disclosed in this patent application under DERWaGS conditions lead to higher concentration of CO than what would be expected from the RWGS equilibrium under the same temperature when fed with a H<NUM>-rich gas with the same CO<NUM> partial pressure at the inlet that the equilibrium partial pressure of CO<NUM> on CaO at the same temperature. No other works in the state of the art have described the DERWaGS phenomena.

An adiabatic heat balance on the step ii), in the absence of input/outputs of energy other than those linked to the endothermic hydrogenation of CaCO<NUM> indicates that the cooling of <NUM>-250ºC taking place during this CaCO<NUM> hydrogenation step will typically be linked to a decrease of <NUM> to <NUM> in the CaCO<NUM> weight fraction in the solids during the step ii) described above. Therefore, in order to make the process continuous, it is necessary to complete the method with a final step iii) of replenishing or regenerating the content of CaCO<NUM> in the solids resulting from step ii).

The equilibrium model solved to quantify the phenomena of DERWaGS reveals optimum operating conditions that further define the preferred conditions of operation during the hydrogenation step ii) to produce a syngas with a volume fraction of between <NUM>-<NUM> of CO, wherein the step ii) is carried at a operating pressure, in atm, which is within the interval resulting from multiplying <NUM>-<NUM> by the equilibrium partial pressure of pure CO<NUM> on CaO by the equilibrium constant of the reverse water gas shift reaction, with the equilibrium constants calculated at any temperature in the interval of temperatures of the calcining solids during their adiabatic cooling in step ii). Such optimum conditions are chosen from the equilibrium diagrams developed for DERWaGS, so that H<NUM>/CO ratios close to <NUM> are reached in the product gas. This will facilitate the use of the syngas in downstream processes of hydrocarbon synthesis (i.e., Fischer Tropsch, Methanol, DME etc). Since the object of this invention assumes availability of low cost H<NUM>, optimum conditions can also be defined as those leading to H<NUM>/CO ratios lower than <NUM>, as long as volume fractions of CO<NUM> in the same product gas is below <NUM>. This is because it would be then easy to correct the H<NUM>/CO ratio to the desired level by simply feeding additional H<NUM> to the product gas coming out from step ii).

The steps i), ii) and iii) are always decoupled: they can take place in the same vessel but in different times, or in separated vessels with a certain transport method of solids between them. When they take place in the same vessel at different times, several reactors (at least two) will operate in parallel, as in state-of-the-art temperature and pressure swing gas separation processes or chemical looping and calcium looping processes using packed beds. When steps i), ii) and iii) take place in separate vessels with a communication between them to allow solid transport between them, the method can use interconnected fluidised bed reactor systems available in the state of the art of chemical looping and calcium looping processes.

In order to reduce the net CO<NUM> emissions and cost of the method, it is preferred to use a range of opportunity fuels and comburent gas mixtures during the energy demanding step i). Opportunity fuels means in this context that they are low cost in the location of the plant and/or that the energy and environmentally efficiency for the combustion process in step i) is the highest. For example, by targeting the production of a flue gas during step i) free of CO<NUM> (i.e., when burning H<NUM> in the step i), or by burning the fuel gas with pure O<NUM> or O<NUM>/CO<NUM> comburent mixtures to produce a gas concentrated in CO<NUM> suitable for permanent geological storage or use. Therefore, preferred embodiments of the previous method can be defined wherein the fuel in step i) is a choice of H<NUM> or a C-fuel, and the comburent a choice between air, enriched air or pure O<NUM>. In such methods, direct combustion will take place in the proximity of the CaCO<NUM> containing solids and heat transfer from the flames or hot gases generated in said combustion will ensure the fast heating up of the CaCO<NUM> containing solids. However, some calcination of CaCO<NUM> will be unavoidable, due to the local hot spots in gas flames known to be present when combustion takes place in a packed bed (or moving bed of packed solids, such as in lime shaft kiln).

To minimise inefficiencies linked to hot spots produced by gas flames, and the associated CO<NUM> losses by calcination during step i), the combustion in step i) can exploit state of the art techniques for chemical looping combustion in packed or moving beds. To this end, a preferred embodiment is wherein the bed of solids containing CaO and CaCO<NUM> also contains a metal-containing solid such as Cu and Fe, with catalytic activity for reverse water gas shift reaction and with oxygen carrying capacity when oxidised to CuO or iron oxides, so that a chemical looping combustion process using CuO/Cu or iron oxides/Fe is used to burn the fuel in step i). The choice of Cu-containing or Fe-containing solids is justified as a way to further enhanced the kinetics of the RWGS reaction in step ii) while allowing the combustion of the fuel and heating of the bed of the solids in step i) under the characteristically moderate and well controlled temperatures in chemical looping combustion applications.

As described above, the method is applicable in continuous mode by feeding natural limestone (or other minerals containing CaCO<NUM>) in step iii) to the adiabatic vessel where step i) takes place, and extracting solids from the reactor where the step ii) takes place. As a result of hydrogenation of CaCO<NUM>, CaO will be a component in said stream of extracted solids. Depending on operating pressures and temperatures in step ii), the catalytic activity of CaO for RWGS will be sufficient to achieve the desired conversions in the RWGS reactions. If CaO is mixed with a Cu or Fe catalyst to enhanced RWGS reaction rates, these catalysts will be extracted with the calcined solids and a segregation or other means of mechanical separation by density or particle size will allow the recycling of the Cu-catalyst or Fe-catalyst to the reactor system, while exporting the resulting CaO product. Fe-catalyst will typically be preferred in these applications respect to Cu-catalyst, because one of the main markets for CaO products is the iron and steel industry, where impurities of Fe in the CaO do not represent a drawback in the quality of the CaO product.

Alternatively, the step iii) of regenerating the content of CaCO<NUM> in the solids resulting from step ii) is achieved by carbonation of CaO in contact with a gas containing CO<NUM>. This means that the method of this invention is part of a larger method or system, involving a capture of CO<NUM> from a gas, as it is the case in any of the variety of methods of calcium looping for CO<NUM> capture.

The practical application of the previous methods requires of devices with preferred features, in particular reactors and reactor systems to allow a continuous operation of the different steps involved in the method. For simplicity, the use of additional devices to preheat reactants, recover heat from gas and solid products are omitted in the descriptions of preferred embodiments below, as they will typically be similar to those described in the state of the art of lime kilns or in the state of the art of CaCO<NUM> calcination and carbonation reactors in calcium looping systems, where heat recovery and preheating of gas and solid reactants are standard practises.

The methods of this invention can be carried out using packed bed reactors containing solids with CaCO<NUM>, CaO and optionally a Cu-based or Fe-based material with reverse water gas shift catalytic properties as well as oxygen carrying capacity for chemical looping combustion of a fuel gas. In a preferred embodiment the method is carried out continuously using at least three packed bed refractory-lined reactors containing at least CaO and CaCO<NUM> operating in parallel, and allowing the switching of gas inlets and outlets to accommodate pressure and temperature swings, further characterised by a synchronised cycling between the following operating modes in each reactor:.

Examples and devices for methods of chemical looping combustion in stationary packed beds are also relevant for a preferred embodiment of the method of this invention wherein at least a fourth reactor is used to allow the combustion of the fuel gas in step i) in two chemical looping combustion sub-steps: a first sub-step wherein the solids containing CuO, are reduced with the fuel gas and a second sub-step where the reduced solids are oxidised with a comburent gas. Such arrangement allows the capture of the CO<NUM> released during the combustion step i), thereby increasing the overall CO<NUM> capture efficiency in the system.

Examples for calcination processes for CaCO<NUM> in moving bed shaft kilns with a single shaft are also relevant for a preferred embodiment of the method of this invention wherein the method is carried out in a single moving bed shaft kiln of limestone, further characterised by a step iii) wherein there is feeding of limestone to the top of the shaft while discharging CaO from the bottom of the shaft, so that the total bed of solids in the shaft move downwards during step iii) before a new cycle of combustion (step i)) and hydrogenation (step ii)) takes place.

Examples of calcination processes for CaCO<NUM> in double shaft kilns are also relevant for this invention because of their inherent high thermal efficiency properties when it comes to use the sensitive heat of all gas and solid products of reaction to preheat solid and gas reactants.

When the methods described above require of a continuous movement of solids containing CaCO<NUM> and CaO in particulate form between reactors, it is also disclosed a device to carry out the methods comprising:.

a solid supply line of CaCO<NUM> connected to any point in the solid circulation loop and a solid purge line connected to the first reactor.

When the methods require of a continuous movement of solids of CaO, CaCO<NUM>, and a Cu-based or Fe-based reverse water gas shift catalyst, it is also disclosed a device to carry out the method comprising:.

The invention also relates to a device to carry out the method involving a regeneration of CaCO<NUM> containing solids by reacting CO<NUM> with CaO, the device further comprising a third fluidised bed reactor acting as carbonator, receiving circulating solids from the second reactor and supplying solids to any point in the solid circulation loop between the two reactors where combustion (step i)) and hydrogenation (step ii)) are taking place.

Reference is made from this point to the accompanying drawings to define preferred embodiments of the invention and provide equilibrium modelling calculations and experimental evidence to support the claims. <FIG> represents the sequence of process steps of the invention. All other embodiments disclosed below contain the steps disclosed in <FIG>. The target of the method is the hydrogenation and calcination of CaCO<NUM> with H<NUM>, to produce CaO and a gas containing at least H<NUM>, CO, CO<NUM> and H<NUM>O(v). The endothermic enthalpy of hydrogenation is taken as ΔH=<NUM> kJ per mol of CaCO<NUM> reacted to CO and H<NUM>O; and ΔH=<NUM> kJ per mole of CaCO<NUM> calcined to CaO and CO<NUM>, with the difference between two enthalpies (<NUM> kJ/molCO<NUM>) being the enthalpy of the reverse water gas shift reaction, RWGS; all enthalpies are assumed independent of temperature, for simplicity in the calculations. The hydrogenation and calcination method of this invention is achieved by repeating three consecutive steps:.

All external energy supplied to drive the endothermic hydrogenation of CaCO<NUM> and the endothermic calcination of CaCO<NUM> is supplied during step i) and iii). The adiabatic cooling marked as ΔT=<NUM>-250ºC in <FIG>, will be proportional to the molar conversion of CaCO<NUM> to CaO during step ii), that comes from the addition of the molar conversions of hydrogenation of CaCO<NUM> to CO and H<NUM>O and the calcination of CaCO<NUM> to CO<NUM>, all multiplied by their respective enthalpies of hydrogenation and calcination. Typically, depending on reactor characteristics to carry out the method (as discussed below) the CaCO<NUM> containing solids will start step ii) at the temperature of Tmax attained at the end of step i). They will progressively cool down as the CaCO<NUM> conversion to CaO increases during step ii). To estimate the range of expected ΔT during adiabatic cooling, it is assumed for simplicity that all specific heat capacities of all solids are 1kJ/kgºC and that gas reactants are preheated to close to the gas product temperatures. CaCO<NUM> has <NUM> weight fraction of CO<NUM> on CaO. Therefore, the release of just <NUM>%w of CO<NUM> from a solid containing CaCO<NUM> will translate into an adiabatic cooling of 48ºC when the CO<NUM> is fully converted to CO by RWGS (the cooling would be of 39ºC if the only products of calcination are CaO and CO<NUM>). Common commercial processes for gas separation using solid sorbents can capture just about <NUM>-<NUM>%w of the sorbent's weight as CO<NUM>. This includes calcium looping CO<NUM> capture systems, where residual CO<NUM> carrying capacities (molar conversion of CaO to CaCO<NUM>) of just around <NUM>-<NUM> are common after just <NUM> carbonation-calcination cycles. In addition, targeting values of molar conversion of CaCO<NUM> to CaO during step ii) higher than <NUM>, characteristic of advanced Ca-sorbent materials, would be counterproductive for the method of this invention, because the associated adiabatic cooling would bring the solids to temperatures where the kinetics of calcination become too slow (i.e., < 700ºC) to design process intensive operations. Therefore, adiabatic cooling between <NUM>-250ºC is the preferred range of temperature change for step ii).

As shown in <FIG>, the combustion step i) is carried out by burning in the proximity of the CaCO<NUM> containing solids a fuel gas with a comburent (for example using combustion methods such as those used in lime kilns, which are the most efficient methods to supply heat directly to CaCO<NUM> containing solids). If the comburent is air (i.e., conventional combustion) and the gas fuel contains carbon, there will be emissions of CO<NUM> diluted in the flue gases coming from the energy intensive combustion step i). To avoid such CO<NUM> emissions, O<NUM> or mixtures of O<NUM>/CO<NUM>, can be used as comburent instead of air, using techniques of oxy-combustion in lime kilns. Alternatively, H<NUM> can be used as fuel gas in this combustion step i), so that there is virtually no CO<NUM> emitted in the flue gases from step i).

Furthermore, to maximise syngas product yields, it is critical for this method that the energy released during the combustion step i) is mainly used to heat up the CaCO<NUM> containing solids and their accompanying solids (that will be at least CaO, resulting from a partial calcination of CaCO<NUM> and a RWGS catalyst, if any) and not to calcine them during step i). When extensive calcination takes place during step i), there is a loss of CaCO<NUM> reactant for the hydrogenation and syngas production step ii) and an unwanted emission of CO<NUM> during step i). Heating CaCO<NUM> containing solids without calcining them (or keeping low their calcination conversion during such heating) is achievable by operating the combustion step i) at conditions close to the equilibrium of CO<NUM> on CaO, so as to maximise the heating rate of solids while minimising their rate of calcination. For example, in lime shaft kilns, operating at close to atmospheric pressure (Pcomb=<NUM>), stones of CaCO<NUM> are heated up to a Tmax of about 900º±50ºC by the combustion flames in the kiln in just a few minutes, well before they can calcine (note that full calcination of a stone in the kiln usually takes several hours). This is consistent with the fact that the equilibrium partial pressure of CO<NUM> on CaO reaches <NUM> atm around 900ºC (see <FIG>) and calcination is controlled by the heat transfer to the calcination fronts that develop in the CaCO<NUM> containing solids at kiln conditions. However, some CO<NUM> emissions from the decomposition of CaCO<NUM> containing solids to CaO will be unavoidable during the combustion step i), due to local temperature profiles when flames have developed in the bed of CaCO<NUM> containing solids of step i). To further reduce these unwanted losses of CO<NUM> during the combustion step i), it is preferred to carry out the combustion step using a chemical looping combustion method to burn the fuel gas used in step i) without contacting the air (or the O<NUM>) used as comburent. As schematically shown in <FIG>, the CaCO<NUM> containing solids are in the proximity of a third solid, containing an oxygen carrier metal oxide (CuO in this case, although it could also be an iron oxide) that can be reduced to its metal form (Cu in this case, although it could also be Fe or a reduced form of the initial iron oxide) while burning the fuel gas. As will be shown below, Cu/CuO and Fe/iron oxides such as FeO, Fe<NUM>O<NUM> or Fe<NUM>O<NUM>, are preferred among other oxygen carriers in the state of the art of chemical looping combustion, because when in reduced form, Cu and Fe are known to have RWGS catalytic activity. Note that the small fraction of CaO produced in step i), or coming unconverted from previous reaction steps in the method, has been recently found to also have catalytic activity for RWGS (Giammaria et al, <NUM>). <FIG> plots the equilibrium constant of the RWGS reaction according to the equation of Twigg in the temperature interval <NUM>-1100ºC of interest for this invention.

The second step in the method of this invention (see <FIG> and <FIG>) is targeted to maximise the hydrogenation reaction of CaCO<NUM> containing solids when reacting them with H<NUM>. Since the CaCO<NUM> containing solids have been previously heated up to an average temperature of Tmax, they will decompose and/or react with H<NUM> when submitted to the conditions in step ii). This step ii) must be carried out under adiabatic conditions and without fully running out of CaCO<NUM> in the reaction environment. This is because, the intention is to use the supply of CO<NUM> from the CaCO<NUM> calcination reaction as a feedstock of pure CO<NUM> for the RWGS reaction. Indeed, as will be discussed below, several preferred embodiments of the method of this invention will operate under conditions that, to our knowledge, have not yet been described in the state of the art: conditions wherein the Desorption Enhanced RWGS phenomena, DERWaGS, appears.

The DERWaGS phenomena is the reverse of what is known in the state of the art as CaO-based Sorption Enhanced Water Gas Shift (Ca-SEWGS) reactions to produce H<NUM> from C-fuels. The DERWaGS equilibrium promotes the maximum formation of CO from CaCO<NUM> and H<NUM>, while minimising the presence of CO<NUM> in the product gas, by operating under conditions where the partial pressure of CO given by the RWGS equilibrium is higher than the partial pressure of CO<NUM> given by the equilibrium of CO<NUM> on CaO. Under DERWaGS, the only feed gas (or initial gas reactant) participating in the RWGS equilibrium is hydrogen and the only source of CO<NUM> in the gas phase is the pure CO<NUM> coming from decomposition of CaCO<NUM>. In these conditions, it is possible to demonstrate that by operating at pressure and temperature conditions that ensure fast calcination kinetics, as well as fast RWGS kinetics, it is possible to increase the productivity of CO in the CaCO<NUM> hydrogenation step, by generating a product gas with CO concentrations higher than those achievable when an equivalent mixture of H<NUM> and CO<NUM> (i.e., CO<NUM> at a concentration equal to the equilibrium concentration at the operating temperature) are fed to a reactor operating at the same temperature and pressure.

<FIG> illustrates the DERWaGS equilibrium in a region of temperatures of interest between <NUM>-900ºC, with a total pressure equal to P=<NUM> atm. The figure has been plotted assuming that both the equilibrium of CO<NUM> on CaO (<FIG>) and the RWGS equilibrium (<FIG>) are fulfilled, and that H<NUM> is the only gas fed to the inlet of a reactor containing CaCO<NUM> containing solids. The fact that CaCO<NUM> exist in the reactor in equilibrium, implies that PCO2 (in atm) in the gas phase is given by equation of <FIG>. The fact that pure H<NUM> is fed to the reactor, implies that PH2O=PCO according to the stoichiometry of the RWGS reaction. This leaves PCO as the only unknown in the resulting quadratic equation when substituting variables in the RWGS equilibrium equation. This allows the construction of DERWaGS equilibrium <FIG>. Surprisingly, preferred pressure and temperature windows of operation for the hydrogenation step ii) emerge from these thermodynamic equilibrium curves.

As can be seen in <FIG>, an optimum temperature for the hydrogenation of CaCO<NUM> exist at 771ºC, because at such temperature, the product gas contains a partial pressure of CO (PCO= <NUM> atm) that corresponds to a molar ratio H<NUM>/CO=<NUM> in the syngas, which is known to be optimum for downstream synthesis processes (i.e., Fischer-Tropsch). Increasing the temperature from that point, increases slightly the yield to CO, although values of H<NUM>/CO decrease. However, values of H<NUM>/CO lower than <NUM> are also preferred because they could be easily corrected to a value of <NUM> by adding H<NUM> (assumed to be a feedstock for this method) to the product gas exiting the reactor during step ii). Operating temperatures higher than 835ºC would be less productive when it comes to converting CaCO<NUM> into CO. This is because at temperatures higher than 835ºC, the CO<NUM> partial pressure, PCO2, in the product gas increases sharply. Therefore, an optimum temperature window of ΔT=<NUM> has been plotted in this figure for illustrative purpose when operating at close to atmospheric pressure. Remarkably, such temperature range is within what is feasible from the heat balance described above when discussing on the expected range of ΔT during adiabatic cooling.

<FIG> also includes a thick dashed curve that shows the equilibrium partial pressure of CO and water vapour that would be achievable at the exit of a RWGS reactor empty of CaCO<NUM> and fed with a gas containing CO<NUM> at the partial pressure given by the equilibrium curve PCO2 in the same figure (with the rest of the gas being H<NUM>). The DERWaGS phenomenon clearly yields higher PCO in the product gas until a temperature of 835ºC (<NUM>) is reached. From that point, calcining first the CaCO<NUM> to obtain a gas with pure CO<NUM> and then feeding H<NUM> to the resulting CO<NUM> stream before entering a RWGS reactor, would yield a RWGS product gas with slightly higher concentrations of CO. However, it is most likely that the method of this invention would still be economically preferable, as all reaction steps are taking place in a single reactor, while the alternative approach in the state of the art is to use two separated reactors (one for calcination and another one for RWGS reactions). In summary, when operating the method at close to atmospheric pressure, maximum temperatures Tmax as high as 850ºC are adequate to mark the end of the combustion step i) and minimum temperature of 750ºC (i.e., ΔT=<NUM>) are adequate to mark the end of the adiabatic cooling taking place during the hydrogenation step ii).

The skilled in the art can reproduce the previous analysis to other operating pressures and temperatures, as shown in <FIG>. <FIG> shows that when applying the method using pressures under the atmospheric (i.e., vacuum pressures during step i) and ii)) it is possible to obtain gases with similar partial pressures of CO, that reach values as high as vCO=<NUM> (= PCO/P) due to the DERWaGS phenomena.

In contrast, <FIG> shows that the DERWaGS phenomena is negligible at the high pressure (P=<NUM> atm) used to build these equilibrium curves. This is due to the steep increase in PCO<NUM> as temperature increases. The method applied at these high pressures demands for temperatures Tmax= 1100ºC to be reached at the end of the step i). A preferred minimum temperature to mark the end of step ii) is 1000ºC, as the H<NUM>/CO ratio becomes lower than <NUM> at lower temperatures. The operation of steps i) and ii) at such extreme temperatures and pressures can be justified if limestone is used as a feedstock in step iii) to produce a syngas that can be fed to Fischer Tropsch or other PtL processes located downstream of the process of this invention, which are usually carried out at this or even higher pressures, with a molar ratio of H<NUM>/CO=<NUM>.

Using results like those in <FIG> for other pressures, allows to plot in <FIG> the volume fractions of CO (vCO=PCO/P) in the product gas of step ii) as a function of temperature and for different operating pressures during the step ii). The experimental points included in the <FIG> will be discussed later. Note that the curves refer to PCO but would be identical for PH2O. It becomes evident from these equilibrium curves that a volume fraction of CO between <NUM>-<NUM> is achievable under a wide range of temperatures if the pressure during the hydrogenation step ii) is adjusted correctly to the temperature of the solids. For example, a target gas composition with an H<NUM>/CO molar ratio of <NUM>, and having vCO=<NUM>, would have a vH2O=<NUM>, vH2=<NUM> and the remaining balance would be vCO2=<NUM>. To identify conditions of temperature and pressure at which such gas product would be produced at equilibrium, first note that PCO2=<NUM> demands for an equilibrium temperature of 766º C at atmospheric pressure. On the other hand, <FIG> shows the solution of the DERWaGS equilibrium when the molar ratio H<NUM>/CO=<NUM>. At the temperature of 766ºC (<NUM>), the PCO given by the DERWaGS equilibrium is PCO=<NUM> atm. Therefore, to maximise the partial pressure of CO to the target value of <NUM>, a total pressure P=<NUM> atm during hydrogenation step would be needed. <FIG> plots such equilibrium dependency of optimum total pressure during hydrogenation step to achieve a vCO=<NUM> in the product gas. The pressure swing between step i) and ii) in the example below is only <NUM> atm (=<NUM>-<NUM>) which is reachable at large scale even by the negative pressure produced by the blowers needed to force the contact of the reacting gases with the bed of solids containing CaCO<NUM>.

As it is derived from <FIG>, the composition of the product gas will change with time as the adiabatic cooling takes place during the hydrogenation reaction step ii), unless pressure changes with time during such step ii). Therefore, it can be operated the hydrogenation step ii) with a continuous pressure change (a decrease) in order to follow the equilibrium curves as shown in <FIG> and maintain a stable volume fraction of CO in the product gas, if this is needed for the processes taking place downstream to use the syngas produced by the method of this invention.

As discussed above, in each cycle of the method of this invention, the step ii) consumes a quantity of CaCO<NUM> that is in the order of between <NUM>-<NUM>%w of the total mass of solids initially present in the bed of solids containing CaCO<NUM>. This means that to run the process continuously, a feeding operation of said mass of CaCO<NUM>, and the extraction the molar equivalent quantity of CaO, needs to be accomplished in the CaCO<NUM> regeneration step iii), as represented in <FIG> ("optional" at the top of step iii)). This can be done using techniques available for lime kilns.

In other preferred embodiments of the method, the process is linked to a CO<NUM> capture step where the solids containing CaCO<NUM> are regenerated in step iii) by carbonation of CaO in contact with a gas containing CO<NUM>, as represented in <FIG> ("optional" at the bottom of step iii)). It is known in the state of the art of calcium looping to capture CO<NUM> with CaO, that the optimum temperature for carbonation is between <NUM>-700ºC for a wide range of flue and fuel gas compositions at close to atmospheric pressure. Such optimum range of carbonation temperatures increases with pressure. On the other hand, the carbonation reaction of CO<NUM> with CaO is exothermic, and would rise the temperature of the bed of solids containing CaCO<NUM> containing solids resulting at the end of a carbonation step iii). However, heat losses and the need to supply the energy required for the RWGS will still demand for a combustion step i), to heat up the solids to Tmax <NUM>-1100ºC, before switching a step ii).

To ensure that the kinetics of the RWGS are sufficiently fast in the temperature interval of the solids between the beginning and the end of step ii) corresponding to adiabatic cooling, the introduction of a known RWGS catalyst (i.e., Cu-based catalyst or Fe-based catalyst) is preferred. Said Cu or Fe catalyst are known to accelerate RWGS reactions. Furthermore, when oxidised to CuO or iron oxides and reduced back to Cu or Fe, they enable the chemical looping combustion (see <FIG>) required to heat the solids to Tmax in step i).

Regarding examples with experimental evidence supporting the previous thermodynamic analysis about the DERWaGS phenomena, the state of the art is silent. To experimentally demonstrate the DERWaGS phenomena, <FIG> show a first experiment conducted in a standard thermogravimetric apparatus testing the rate of calcination of CaCO<NUM> fine particles at atmospheric pressure and 750ºC. The sample of <NUM> of CaCO<NUM> is previously heated in an atmosphere of pure CO<NUM> at 750ºC to avoid decomposition, and then the gas contacting the sample is switched to have a pre-set vol fraction equal to <NUM> (which is only slightly below the equilibrium value of <NUM> of CO<NUM> vol at 750ºC according to <FIG>). Two curves of molar calcination conversion, Xcal, vs time are plotted under the same <NUM> value of volume fraction of CO<NUM>: when diluted in N<NUM> (dashed line) and when diluted in H<NUM> (solid line). As can be seen, in the case of N<NUM>-CO<NUM> the calcination rate is extremely slow. This is consistent with the state of the art of calcination kinetics of CaCO<NUM> in the proximity of the equilibrium of CO<NUM> on CaO. In sharp contrast, the calcination rate of CaCO<NUM> accelerates when the test is carried out at the same temperature and using a mixture of H<NUM>/CO<NUM> with identical CO<NUM> volume fraction than the previous N<NUM>/CO<NUM>. It is remarkable the acceleration observed in the conversion vs time curve, perhaps linked to the increasing availability of CaO in the sample of solids, enhancing through catalytic effects the RWGS reaction in the proximity of the calcining solids, as recently identified by Giammaria et al. Irrespective of the underlying mechanism, it becomes evident from <FIG> that the assumption of fast calcination rate during hydrogenation step ii), leading to <FIG>, is well supported by experimental evidence, because there is an acceleration in orders of magnitude of the rate of calcination during hydrogenation of CaCO<NUM> respect to the rate of calcination in a N<NUM>-rich atmosphere.

Regarding more direct experimental evidence for DERWaGS, <FIG> shows the experimental set up that has been used to gain such evidence. This includes a packed bed reactor (<NUM> internal diameter and <NUM> maximum height of bed of solids) filled up with CaCO<NUM> containing solids, together with CaO (optional) as RWGS catalyst and Cu/CuO containing solids (optional) as RWGS catalyst and as oxygen carrier. To facilitate experiments in such a small device, the step of heating by combustion is substituted by a step of preheating the solids with an electric oven to the pre-selected temperature in an atmosphere of pure CO<NUM> to avoid CaCO<NUM> decomposition during heating. A fast switching of valves allows the sudden feeding of H<NUM> to the preheated solids (for convenience, in order to trace gas concentrations in the existing analyser, a <NUM>%v of N<NUM> is used to dilute the H<NUM> fed to the reactor). As can be seen in <FIG>, the product gas has a composition fully consistent with the DERWaGS equilibrium during the first <NUM> seconds of the experiment. The packed bed configuration of the solids and the nature of the experimental set up does not allow for a precise control of local temperatures in different parts of the bed. Therefore, two dashed lines show the theoretical gas composition in a reasonable interval of temperatures in the bed of solids (see Figure caption for more details).

<FIG> shows for illustrative purposes a few points on the right hand side that correspond to a period when the hydrogenation of CaCO<NUM> has been completed and the concentration of H<NUM> at the outlet of the reactor equals the concentration at the inlet.

<FIG> confirm the results of <FIG> in experiments conducted at other different pressures and temperatures (see captions).

In experiments of <FIG>, there is a fraction of a commercial Cu-based catalyst that is known to be active for RWGS reactions. In contrast, <FIG> corresponds to an experiment with no such Cu-material in the bed. Therefore, the results of <FIG> indicate that CaO formed during calcination displays a substantial RWGS catalytic activity during the DERWaGS reactions, consistent with results of <FIG> and the state of the art given by Giammaria et al (<NUM>). Indeed, when using the kinetic model provided by Giammaria et al (<NUM>) for RWGS in the presence of CaO, and assuming plug flow of gases in the reactor used in <FIG>, it is possible to predict the compositions at the exit of the reactor and confirm similar observations.

Taking into account the results and discussions disclosed in previous paragraphs, it is possible to carry out the method of this invention using state of the art reactors and techniques in the fields of lime calcination, calcium looping for CO<NUM> capture and chemical looping combustion.

As a first example, <FIG> discloses a preferred device to carry out the method of this invention using lime kiln technology with a single shaft. The kiln includes a refractory lined (<NUM>) external wall (<NUM>) to sustain high temperature operations in the interior of the bed of stones and pressure swings as needed for the different embodiments of the method described above. The kiln also incoporates lock-hoppers connected to a solid feeding valve (<NUM>) and a solid extraction valve (<NUM>) of solids and a number of gas switching valves (<NUM>, <NUM>, <NUM>, <NUM>, <NUM>) installed in all inlet and outlet gas pipes. At least one of the pipes (usually the pipe of flue gases (<NUM>, <NUM>)) comprises a T pipe diversion, with a valve (<NUM>) to divert syngas product gas (<NUM>) during hydrogenation of CaCO<NUM>. Such kiln will typically contain a preheating zone (<NUM>) for limestone (<NUM>) fed at the top of the kiln and a gas preheating section (<NUM>) at the bottom of the kiln that cools the lime product (<NUM>) with air or other gases in the comburent mix (<NUM>) used during the combustion of a gas fuel (<NUM>) to heat up the solids in the hydrogenation and calcination zone (<NUM>) in step i). The process operates with a sequence of operation periods characterized by the positions of the switching valves as indicated in <FIG>: a period of step i), typically lasting a first time t1=<NUM>-<NUM> minutes as in existing lime kilns for limestone calcination, with a valve (<NUM>) opened for the emission of CO<NUM> produced during the fue combustion (<NUM>) mixed with other flue gases (<NUM>) from combustion, a valve (<NUM>) opened for the entry of a fuel (<NUM>) and a valve (<NUM>) opened for the entry of the comburent gas containing air, enriched air, or O<NUM>/CO<NUM> or O<NUM> (<NUM>); while the solid feeding valve (<NUM>) and the solid extraction valve (<NUM>) are closed as well as valves (<NUM>) and (<NUM>). After the period of step i), it follows a period of step ii) lasting a second time t2 between <NUM> and <NUM> minutes to allow syncronise operation of different shaft kilns, characterised by the addition of a H<NUM> supply line (<NUM>) by keeping switching valve (<NUM>) opened, keeping the valve (<NUM>) to divert product syngas (<NUM>) produced as the hydrogenation of CaCO<NUM> progresses. Valves (<NUM>), (<NUM>), (<NUM>), (<NUM>) and (<NUM>) are closed during step ii). For simplicity, the evolution of the calcination conversion of the original limestone particles (<NUM>) towards fully calcined lime (<NUM>), has been represented in <FIG> as concentric spheres of black (for limestone) and grey (for lime). For simplicity, in this configuration, the step iii) is taken a short period when all valves are closed except the feeding valve (<NUM>) and the extraction valve (<NUM>) of solids. It is possible to carry out small changes in this strategy to overlap the step iii) with a certain period during steps i) or ii). Other small variants of the process may refer to the feeding point of the comburent mix (<NUM>), for example by feeding (<NUM>) to the bottom of the calcination and hydrogenation region (<NUM>) to avoid the recarbonation of CaO when (<NUM>) is a comburent mix of O<NUM>/CO<NUM>, following the teachings of other oxy-fired shaft kilns. A vacuum pump (<NUM>) is optionally included in <FIG> for preferred embodiments when the shaft kiln is operated at pressures under the atmopheric pressure. For simplicity, the device of <FIG> is a single shaft kiln, but similar modifications would be also valid for double shaft lime kilns. Other variants of the process may refer to the addition of a material containing Fe mixed with the feed of limestone (<NUM>), to catalyse RWGS reactions if the CaO present in the hydrogenation and calcination region (<NUM>) does not provide sufficient catalytic activity.

<FIG> discloses a preferred device to carry out the method of this invention using packed bed reactors allowing pressure and temperature swings, wherein the method is carried out in a system composed of several adiabatic packed-bed reactors operating in parallel in different reaction steps, to allow the continuous treatment of gas streams in a synchronised manner as in temperature and pressure swing systems. In this particular example, the bed is composed of partially carbonated lime particles (equivalent to the partially calcined limestone particles represented in <FIG> as concentric spheres of CaCO<NUM> (<NUM>) and CaO (<NUM>) randomly mixed with a Cu-based material (represented as an empty circle when in reduced form and as a circle with a cross when in oxidized form). The example would be conceptually similar if Fe was used instead of Cu, and iron oxides were used instead of CuO. These materials are available in a wide range of compositions in the state of the art of chemical looping combustion and reverse water gas shift catalytic processes.

The process works cyclically, and the description can therefore be initiated from any step. Since the application of the method of this invention (steps i) and ii) as in <FIG>) requires a packed bed of solids containing CaCO<NUM>, it will be firstly described the step involving the generation of CaCO<NUM> in the bed. An illustrative example is when the CaO carbonation step represented as iii) in <FIG> is a CASOH step to decarbonize Blast Furnace Gases, BFG, (typically containing between <NUM>-<NUM>%vol of CO<NUM> and <NUM>-<NUM>% of CO with the remaining gas being mainly nitrogen). The CASOH step is a calcium-enhance water gas shift reaction where the valve (<NUM>) and the valve (<NUM>) are opened to allow the feed of BFG (<NUM>) and steam (<NUM>) to allow the enhanced water gas shift of the CO contained in the BFG. The valve (<NUM>) is also opened to allow the extraction of the CASOH H<NUM> rich product gas (<NUM>). A variety of temperature and pressure conditions known in the state of the art allow first a very fast water gas shift reaction of CO with H<NUM>O to form CO<NUM> and H<NUM> catalysed by the Cu-based solids. Then, the CO<NUM> reacts with active CaO to form CaCO<NUM> as soon as it is produced, which shifts the water gas shift equilibrium towards a higher production of H<NUM>. The reaction of CaO with the CO<NUM> originally contained in the BFG as well as the CO<NUM> produced in the water gas shift reaction of the CO originally contained in the BFG is represented in <FIG> as a sharp carbonation reaction front. The low content of active CaO in the bed (known to be typically lower than <NUM> of the total calcium in particles cycled more than <NUM> times in calcium looping systems) makes the carbonation front leaving behind the carbonated solids at higher temperature due to the exothermic nature of the carbonation and the water gas shift reactions. When the active bed of solids reaches close to total carbonation and the reactor is no longer capable of absorbing CO<NUM>, the step iii) must be terminated and step i) is initiated. In this example, step i) is a heating up step of the bed of solids containing CaCO<NUM> by combustion of a fuel (<NUM>) (it can be BFG or other fuel gas) to supply the necessary energy to the bed of solids that will drive the hydrogenation reactions during step ii). To this end, the combustion of a fuel gas (<NUM>) in step i), is divided into two chemical looping combustion sub-steps: j) and jj). In the sub-step j), the oxidation of the Cu-based solids to CuO can be carried out with a comburent (<NUM>) such as air, enriched air, O<NUM>/CO<NUM> or even pure O<NUM> if available at low cost from the electrolyser providing H<NUM> from for step ii). The use of pure O<NUM> (or a mixture of O<NUM>/CO<NUM>) avoids any losses of CO<NUM> by calcination of CaCO<NUM> during this sub-step j). Assuming the complete conversion of the oxygen, the outlet flow (<NUM>) during this stage is zero (it would be mainly composed of nitrogen if the oxidizing gas (<NUM>) was air or enriched air). Once the Cu has been totally oxidized, the next step jj) completes the chemical looping combustion of the fuel gas flow (<NUM>). This consists of reducing the CuO solids with the fuel gas (<NUM>) while producing H<NUM>O and CO<NUM> (when the fuel (<NUM>) contains carbon, if (<NUM>) is H<NUM> the only product of step jj) is H<NUM>Ov). During the reduction of CuO, the reaction front moves forward and the heat released during the CuO reduction is dedicated to further increase the temperature of the solids left behind the reduction reaction front. Once the CuO has been completely reduced to Cu, the packed bed is left at the target Tmax, to initiate the hydrogenation step ii). For simplicity, the fuel gas (<NUM>) is assumed to be at the same temperature as the solids at the end of step jj). When feeding gas reactants at lower temperatures to the packed bed reactor, a heat transfer front will develop (not shown in <FIG> for simplicity) giving as a result a colder region that grows from the entry point of the gas reactants and that is at the temperature of the gas reactants. The skilled in chemical looping combustion in packed beds can provide solutions to manage the appearance of such colder regions, for example by operating the packed beds with reverse flow of the inlet gases, by using regenerative heat exchangers to preheat feed gases, or by loading the packed bed with higher concentration of the metal (Cu or Fe) at the part of the reactor closest to the gas entrance, to compensate with additional heat (released in the combustion of a fuel gas in step i)) the cooling effect of the lower temperature feed gases.

Following the method of this invention, the next step involves step ii) in which DERWaGS pressure conditions described above are imposed on the preheated and carbonated solids to achieve the direct hydrogenation of CaCO<NUM> and promote the adiabatic cooling of the solids in the reactor. To this end, hydrogen (<NUM>) is fed to the bed of solids contained in (<NUM>) by opening valve (<NUM>), and to extract a syngas (<NUM>) from the previous heated solids containing CaCO<NUM>. This is allowed because the packed bed of solids preheated in step i) contains at the beginning of this step ii) sufficient sensitive heat to accomplish the calcination of the CaCO<NUM>, formed in step iii), leaving the bed at the assumed temperature to start a new CASOH step iii) in a subsequent cycle.

Note that the application of the method to capture CO<NUM> to other fuel gases containing CO<NUM> or to a combustion flue gas by carbonation of CaO in a packed bed, would not differ in any essential feature from the case described in the previous paragraph. The skilled in the art would prepare the bed of solids with a suitable molar ratio between active CaO and the metal used as oxygen carrier and as reverse water gas shift catalyst, so that the endothermic step ii) added to the heat losses expected in the system and the differences between inlet and outlet gas temperatures will be balanced by the heat released during the combustion step i), so that the method can be run with steady state repetitive conditions during cycles. In the device of <FIG>, the vacuum pump (<NUM>) is optional because it is only used when step ii) requires vacuum conditions to reach DERWaGS equilibrium conditions (see for example <FIG>). In applications when step ii) is carried out at high pressure, a valve (<NUM>) will be a pressure relief valve instead of an on/off valve and in applications where vacuum pressure is applied to step ii) the valve (<NUM>) will also have the role to control the vacuum pressure in the reactor.

All the previous devices and examples are using packed bed or moving bed technologies, which imply that the transitions between steps i), ii) and iii) are linked to on/off valve switching of gases after certain periods of operation (t1, t2) in parallel reactors. However, the transitions between steps i), ii) and iii) of <FIG> can be accomplished by circulating solids between different reactors. <FIG> shows the schematic of a device to carry out the method of the invention using interconnected fluidized beds, using the same numbering as in <FIG>.

Claim 1:
Method of hydrogenation and calcination of CaCO<NUM> in a thermally insulated vessel by reacting it with H<NUM>, to produce CaO and a syngas containing at least H<NUM>, CO, H<NUM>O and CO<NUM>, characterised by a sequence of at least the following consecutive steps carried out at least twice:
i) a step of combustion of a fuel to heat up a bed of solids containing at least CaO and CaCO<NUM> to a maximum heating temperature between <NUM>-1100ºC, at a pressure being essentially the equilibrium partial pressure of pure CO<NUM> on CaO at the maximum heating temperature;
ii) a step of reacting with H<NUM> the heated bed of solids containing at least CaO and CaCO<NUM>, while allowing their adiabatic cooling by <NUM> to 250ºC;
iii) a step of replenishing or regenerating the content of CaCO<NUM> in the solids resulting from step ii).