Steam partial pressure swing calcination process for producing high-purity CO2 from CaCO3

Abstract

The present invention discloses a calcination process for producing high-purity CO2 from solids containing CaCO3, which operates cyclically and continuously on solids arranged in a packed bed or moving bed. Each cycle includes a first step in which fuel combustion with a bed of CaCO3-containing solids heats them to an average temperature of 800-950°C, and a second step in which a CO2 partial pressure swing is used to extract pure CO2 from the CaCO3-containing solids by supplying steam to the bed. The combustion can be carried out directly with air, oxygen-enriched air, or an O2 / CO2 mixture when the process is applied to the calcination of a continuous flow of limestone in a moving-bed shaft kiln. The process can also be used to calcinate CaCO3 formed in a reversible calcium looping process, which includes a carbonation reaction step to form CaCO3 from CaO.

Description

[Technical Field]

[0001] Object of the invention The present invention discloses a calcination process for producing high-purity CO2 from solids containing CaCO3. The process operates cyclically and continuously on solids arranged in a packed bed or moving bed. Each cycle of the process comprises a first step in which fuel is combusted with a bed of CaCO3-containing solids, heating them to an average temperature of 800-950°C, and a second step in which steam is supplied to the bed of CaCO3-containing solids to apply a CO2 partial pressure swing, calcining them, and causing adiabatic cooling, preferably between 30 and 200°C. The combustion can be carried out directly using air, oxygen-enriched air, or an O2 / CO2 mixture when the process is applied to the calcination of a continuous flow of limestone in a moving-bed shaft kiln. The process can also be used to calcinate CaCO3 formed in a reversible calcium looping process, which involves a carbonation reaction step to form CaCO3 from CaO. In these cases, the preferred method of combustion of the fuel used to heat the bed of solids containing CaC0 is a chemical looping process using a second solid containing a metal oxide (preferably NiO or CuO) as an oxygen carrier, mixed into the bed of solids containing CaC0.

[0002] Technical Field This is a process for recovering pure CO2 during the calcination of CaCO3. When the source of CaCO3 is natural limestone and the product is CaO, the process is applicable to shaft kilns. When the source of CaCO3 is a process that includes a carbonation step of CaO, the process is applicable to calcium looping processes for recovering CO2 from flue gases, or adsorption-enhanced reforming or adsorption-enhanced water-gas shift processes in which carbonaceous fuels are converted to hydrogen while CaCO3 is formed from the carbonation of CaO. [Background technology]

[0003] The capture of CO2 in a highly pure form (e.g., CO2 having a concentration greater than 99% v) from fuel and flue gases in industrial processes is of interest in terms of climate change mitigation efforts. The original source of carbon in the gas can be from the carbonaceous fuel used to power the process or from the decomposition of natural sources of CaCO3 (i.e., when making lime and cement).

[0004] Alternatively, there are high-temperature calcium looping CO2 capture systems (typically at temperatures of 600-700 °C) that use CaO as a CO2 adsorbent to form CaCO3. CO2 can be captured from combustion flue gases or during adsorption-enhanced reforming processes or calcium-enhanced water-gas shift processes to produce hydrogen from fuel gases while capturing CO2 at high temperatures using CaO. These high-temperature operations facilitate the recovery of the equivalent energy consumed during the exothermic carbonation of CaO (-170 kJ / mol CaO at 650 °C) as during the calcination of CaCO3, which typically occurs at temperatures above 900 °C (depending on the ambient CO2 concentration).

[0005] Calcination of CaC03-containing solids using vacuum calcination conditions has been described in the state of the art. The CaC03-containing solids may be natural minerals such as limestone or carbonated solids from calcium looping CO2 capture processes.

[0006] U.S. Patent Application Publication No. 2017 / 0096335 claims greater energy efficiency by operating an adiabatic calcium-looping packed-bed reactor at a sufficiently high pressure during the exothermic carbonation step so that the carbonated solids reach a high enough temperature to initiate a reversible vacuum-driven calcination step after each carbonation step. The CaO solids are heated to a maximum of 150°C at the end of the exothermic carbonation period. It is claimed that this temperature increase is sufficient to sustain a subsequent vacuum calcination step, which causes an equivalent temperature decrease as the CaCO3 is calcined. Note that the system is assumed to be adiabatic, so no external heat is applied to the packed bed of sorbent. This means that the sorbent is heated during carbonation and cooled to its original temperature during calcination, endlessly during successive carbonation / calcination cycles. However, such a process faces practical limitations that may make it impractical. On the other hand, real systems are not completely adiabatic. Heat transfer from external sources or from gaseous reactants supplied at higher temperatures than the gaseous product is necessary to satisfy the overall heat balance. However, there are constraints imposed on the temperature of the gaseous reactants from the additional heat transfer equipment required because, in some cases, preheating such reactants is not feasible. For example, in an adsorption-enhanced water-gas shift process using CaO as a CO2 adsorbent, the synthesis gas (containing CO) must be supplied at a temperature below 500 °C to avoid carbon deposition. In this regard, metal dusting in heat exchangers used as gas preheaters must be prevented by supplying the gas at temperatures as low as 300 °C. However, the product gas tends to exit the reactor at higher temperatures (i.e., between 650 and 850 °C, depending on the temperature reached during CaO carbonation).

[0007] The problems associated with the need for a heat supply to maintain calcium looping system operation are exacerbated when using packed beds and when considering the typically low molar carbonation conversion rates of real CaO materials when reacting with CO2 over many cycles. Despite intense research efforts over the past several decades, there are still no suitable, widely available CaO sorbents (i.e., with demonstrated mechanical stability and reactivity to rapidly and reversibly react with CO2 under realistic reactor conditions) that can maintain molar carbonation conversion rates of greater than 3–10% from CaO to CaCO3 over hundreds or thousands of cycles, as is the case for CaO derived from calcination of natural limestone. This means that active CaO in calcium looping processes using currently known CaO materials is typically accompanied by a very large portion of material inactive to carbonation (95–99% w, including other solids in the system, such as catalysts and their supports). Therefore, despite the exothermic nature of CaO carbonation, the achievable temperature rise during carbonation does not exceed 50–150°C (depending on the carbonation conversion rate and the proportion of inert materials), which does not allow for the operational period for effective CO2 capture by carbonation (600–700°C) to be connected to the operational period for effective calcination of CaCO3 in pure CO2 (the decomposition temperature of CaCO3 at atmospheric pressure in pure CO2 is approximately 900°C). Therefore, when operating with known CaO adsorbent properties (i.e., maximum CO2 loading capacities of 0.05 and 0.1) and known operational periods for carbonation and calcination, it is known that additional heat requirements during calcination are necessary. Furthermore, U.S. Patent Application Publication No. 2017 / 0096335 does not mention the need for an external heat supply to their calcium looping system.

[0008] U.S. Patent No. 9,505,998 discloses a solution to the problem of heat supply to a packed-bed calciner by locating a combustion chamber for providing additional thermal energy to the calciner through a metal wall separating an area containing a bed of solids containing CaCO from the combustion chamber. However, this solution has limited applicability because it requires a large heat transfer area of ​​expensive alloys to enable the necessary heat transfer between the combustion chamber and the calciner at temperatures above 900°C, especially considering that the two may need to operate under a significant pressure differential.

[0009] U.S. Patent No. 8,506,915 discloses another solution to the problem of heat requirements for calcination in a bed of CaCO3-containing solids by introducing a second solid into the Cu / CuO-containing bed. This is a type of oxygen carrier material being developed and tested in the state of the art for chemical looping combustion or indirect combustion applications. In the chemical looping combustion process, metals (such as Ni, Cu, and Fe) are oxidized at high temperatures in an "air reactor" in a highly exothermic process. In the next step, the metal oxide or "oxygen carrier" (such as NiO, CuO, or Fe2O3) is reduced by fuel gas to produce N2-free flue gas and regenerate the metal or "metal fuel." However, the use of a metal fuel oxidation reaction with air to provide energy for CaCO3 calcination has the problem of undesirable dilution of CO2 generated during calcination by the nitrogen contained in the air. To address this issue, a key feature of U.S. Patent No. 8,506,915 is the selection of a Cu / CuO chemical loop, in which CuO undergoes an exothermic reduction reaction to Cu using fuel gas (primarily combusted to CO and HO), providing the solid bed with the energy necessary to drive the calcination of CaCO during the CuO to Cu reduction step. This process has been proposed for the adsorption-enhanced reforming of methane and steam to produce a hydrogen-rich stream and high-purity CO, using a bed of solids containing a nickel reforming catalyst as a third, additional solid to the CaO adsorbent and Cu oxygen carrier (Fernandez et al., Conceptual design of hydrogen production process from natural gas with CO capture using Ca-Cu chemical loop, International Journal of Greenhouse Gas Control 2012, |Volume 6|, pp. 126-141).Another variation of this process has recently been proposed, treating steel mill off-gases such as blast furnace gas (BFG) in the so-called CASOH process. A bed of CaO-containing solids first removes the CO contained in the BFG and then recovers the resulting CO as CaCO, promoting an enhanced water-gas shift of CO. In a subsequent step, reduction of CuO with a fuel gas such as CH can regenerate CaO while producing CO and water vapor (Fernandez et al., Advanced Packed-Bed Ca-Cu Looping Process for the CO Capture From Steel Mill Off-Gases, Frontiers in Energy Research 2020, July 2020 | Volume 8 | Article 146). The design of these combined calcium looping and chemical looping processes shows that they can be highly energy-efficient if, and only if, effective heat recovery for energy generation is included in the system. This is because these processes must function with an additional intermediate reaction step, involving the oxidation of Cu with air, before the desired reduction-calcination step. Such oxidation poses the demand for technically challenging operation at high pressures, necessary to minimize calcination of CaCO3 during the oxidation of Cu. Maximizing H2 yield (as this is the targeted product from such processes) is technically more attractive than maximizing thermal energy output from the process (even in the form of a hot gaseous product stream), and the need for additional energy generation equipment is a weakness in practical applications targeting maximum H2 production yield from adsorption-enhanced reforming or calcium-enhanced water-gas shift processes.

[0010] Also related to this invention is a patent relating to the application of vacuum pressure to the limestone calcination process so that the CO2 generated from the calcination of CaCO3 can be obtained separately, or even in pure form.

[0011] U.S. Patent No. 4,748,010 provides a detailed review of the prior art for vacuum calcination of batches of limestone to produce lime. It discloses a semi-continuous CaCO3 calcination process that attempts to take advantage of the favorable calcination conditions achieved when a vacuum is applied to preheated stone containing CaO and CaCO3. In the first step, the limestone is partially calcined by heating it with air combustion until 50-85% calcination conversion is achieved (i.e., % of moles of CaCO3 converted to CaO), resulting in a hot mass with sufficient stored heat to complete the calcination of the remaining unconverted limestone in one final cycle. To accomplish this, the solid hot mass is transferred to a vacuum chamber zone, where the heat retained in the solid is used to achieve complete conversion of the limestone, while the partially calcined hot mass is subjected to a sufficiently strong vacuum for the time required to achieve complete conversion of the limestone. It should be noted that U.S. Patent No. 4,748,010 is not intended to be a CO2 capture process, as the majority of the CO2 generated during calcination is released in diluted form in the flue gases released during the first heating step and partial calcination of the limestone under standard combustion conditions in a kiln. Furthermore, it is important to note that the proportion of pure CO2 achievable by the method disclosed in U.S. Patent No. 4,748,010 does not in practice exceed a molar conversion of 10% of the initial mass of limestone due to known thermodynamic and kinetic constraints, which are briefly discussed below.

[0012] The latest advances in the calcination of CaCO3 provide further experimental support for the phenomenon of adiabatic cooling of CaCO3-containing solids preheated to their calcination temperature when subjected to rapid changes in the partial pressure of CO2 in their gaseous environment, or pressure swings. Hills, A.W.D., "The mechanism of the thermal decomposition of calcium carbonate," Chemical Engineering Science, 1968, Vol. 23, pp. 297–320, investigated how a decrease in the partial pressure of CO2 around a calcined stone (or a stone containing a CaCO3 core) initially results in a rapid calcination rate and a decrease in the temperature of the carbonated core. For example, Hills's Figure 4 shows the decrease in temperature inside a 1-cm-diameter calcined particle after undergoing a step change in the partial pressure of CO2 around the particle, which was held at atmospheric pressure. The temperature decrease recorded inside is a manifestation of rapid heat transfer from the interior to the exterior of the CaCO3 stone surface, partially driving the calcination. After this initial temperature drop (about 50°C in less than 5 minutes), the interior portion of the stone must reach the exterior temperature, and the slow heat transfer from the exterior CaO portion of the partially fired stone to the shrinking carbonated core becomes the dominant process in the overall firing rate of the stone. Note that in this example in Figure 4, full firing of a 1 cm diameter particle is achieved after 70 minutes in a furnace environment of 867°C.

[0013] In summary, there remains a need to develop an efficient method for calcining CaC0 to produce pure streams of CaO and CO2 for a wide range of applications, such as lime kilns or calcium looping systems for CO2 capture from gas, which involve repeated cycles of CaC03 calcination and CaO carbonation steps. Summary of the Invention

[0014] The primary objective of the present invention is to provide a process for recovering, in high-purity form, a portion of the CO2 released from the calcination of solids containing CaC03. The process is applicable to limestone calcination processes in shaft kilns, which are insulated vessels having a moving bed of solids containing CaC03 in the calcination zone of the shaft kiln. The method is also applicable to calcium looping processes using an insulated vessel containing a packed bed of solids containing CaC03 as formed during CO2 recovery by CaO in contact with combustion flue gas or any other gas containing CO2, or during adsorption-enhanced reforming or adsorption-enhanced water-gas shift process steps that convert fuel gas to hydrogen when CaC03 is formed in the presence of CaO and other gases.

[0015] The present invention provides a calcination process for recovering CO2 in a high purity form after separation of CO2 from steam, wherein CO2 is released from the calcination of a solid containing CaCO3, said solid being preheated to a temperature of 650-800°C and placed in an insulated vessel (1), the process comprising at least the following series of consecutive steps: i) a heating step in which a fuel is combusted within the bed of solids to heat the bed of solids containing CaCO to an average temperature of 800-950°C; and ii) CO2 partial pressure swing process by feeding steam to a bed of solids; It is characterized by at least two consecutive repetitions of Preferably, the heating step lasts from 5 to 20 minutes and / or the CO2 partial pressure swing step lasts from 5 to 20 minutes.

[0016] Thus, the process yields CO2 and CaO as separate reaction products after steam condensation. Calcination under reduced partial pressure of CO2 around the solids, imposed by injection of steam heated at temperatures above 550°C (to avoid the risk of CaO hydration), causes a temperature drop of 30-200°C in the bed of solids. The temperature drop is the result of an increase in the molar conversion of CaCO3 to CaO of 0.01 to 0.07 under the near-adiabatic conditions encountered in the insulated vessel in which the calcination is taking place. Preferably, the combustion of the fuel is carried out using air, oxygen-enriched air, or an O2-CO2 combustion-sustaining mixture. Even more preferably, the fuel is a hydrocarbon gas or hydrogen.

[0017] The intent of the preceding series of steps is to take advantage of the large difference between heating and calcination times for the typically large particle size (10-150 mm) of limestone processed in shaft kiln calciners, or for particles or pellets (2-10 mm) used in packed-bed calcium looping processes. As quantified below in several examples and experiments, and as known in the state of the art of shaft kiln operation and calcium looping, the rapid heating rate of the particles in step i) relative to their calcination rate in the same step i) allows for the introduction of a certain amount of thermal energy into the bed of particles to heat them while limiting the progress of calcination of the CaCO3. However, this step inevitably causes some calcination of the CaCO3 and releases some of the CO2 generated in such calcination as diluted gas in the flue gas generated by the combustion of fuel in the bed. At the end of step i), the solids are heated to an average temperature of 800-950°C from the low initial temperature (30-200°C lower) resulting from the injection of steam in the previous step ii). A higher average temperature after heating step i) leads to a greater loss of CO2 during heating step i) due to increased calcination of CaCO3. To minimize such losses, heating step i) must be performed at high pressure, increasing energy consumption and costs. On the other hand, an average temperature below 800°C at the end of heating step i) requires a longer period of step ii) because the rate of calcination slows during the adiabatic cooling imposed in step ii). Using oxygen-enriched air or a mixture of O2 and CO2 as the combustion medium during the combustion of fuel in step i) instead of air has a beneficial effect in minimizing the decomposition of CaCO3 during combustion step i). This is because combustion with a higher O2 content allows for a higher concentration of CO2 in the flue gas, which is known to reduce the calcination rate. For example, if the CO2 content of the CO2-enriched flue gas is 50% v at atmospheric pressure, the calcination rate of CaCO3 is very slow compared to the heating rate of the stone until the stone reaches a value close to 850°C (see other examples below). This is because the equilibrium partial pressure of CO2 over CaO at 850°C is close to 0.5 atm.The use of the combustion product with a high O content during step i) also facilitates the introduction of the energy required for the firing of step i) in a shorter time. This is beneficial because it allows step ii) to be initiated in a process where the overall / average heat input to the kiln remains unchanged over an extended period of time. In other words, initiating step ii) without increasing the energy input during step i) necessarily leads to a decrease in the average production rate of CaO. This directly affects the rate of lime production in the shaft kiln or the speed or capacity of a calcium looping recovery system that uses such CaO to recover CO.

[0018] Therefore, preferred embodiments of the present invention further include fuel combustion using oxygen-enriched air or an O2-CO2 combustion-sustaining mixture. As illustrated in the example lime kiln below, making the duration of step i) identical to the duration of step ii) would require doubling the energy input during step i) to maintain the same overall lime production capacity relative to prior art processes without recovery. Therefore, the use of a combustion-sustaining mixture with approximately 40% vol O2, which is part of the state-of-the-art in oxy-combustion systems, facilitates this doubling of energy input while maintaining a similar flow rate of gases within the kiln, which is important for maintaining heat transfer characteristics, bed pressure drop, energy efficiency, etc., relative to those of modern kilns without CO2 recovery.

[0019] Once the stone or carbonated sorbent is heated to a temperature of 800-950°C at the end of step i), the onset of step ii) promotes rapid partial calcination of the CaCO3-containing solid, which is associated with an equally rapid drop in temperature within the solid. In an industrial-scale packed bed with many stones or carbonated sorbents simultaneously calcining under low partial pressures of CO2, CO2 release is limited at this point by the maximum capacity of downstream equipment to process the CO2 and steam streams emerging from the bed. However, as calcination progresses, the calcination rate eventually decreases over time, reducing the availability of heat within the stone and creating a temperature profile within the stone or sorbent. The temperature within the stone drops to a value that makes the intrinsic calcination rate too slow for the conditions of the calcination reaction front temperature and partial pressure of CO2 (which can be assumed to be in equilibrium at temperatures above 700°C). Considering the state-of-the-art experimental and modeling results, and as discussed below, after 5-20 minutes of calcination under reduced CO partial pressure in step ii), a decrease in the average temperature of the solid bed (30-200°C) occurs. This is associated with an increase in the molar conversion of CaCO to CaO, or a decrease in the CaCO mole fraction, of 0.01 to 0.07 under adiabatic conditions. At this point, a return path must be initiated to reheat the solid bed. This is the purpose of step i), which involves a combustion step within the solid bed to facilitate rapid heat transfer of combustion energy to the solids that were just cooled during step ii).

[0020] If the insulated vessel is a lime shaft kiln, the above cycle is repeated until each stone of limestone initially fed into the kiln is fully calcined in the shaft kiln. Such a kiln can be operated in a continuous mode, preferably in a moving bed configuration, by introducing a batch of solid particles containing CaCO3 into the insulated vessel, preferably at the top of the kiln, so that the solids move downward with each combustion step, while a batch of particles containing CaO is extracted from the bottom of the kiln. In such a shaft kiln, steam is added to the bottom of the calcination zone to avoid contact of the CaO product with the steam in the cooling section of the kiln at temperatures below 550°C (thus avoiding hydration of the CaO, which could result in CaO expansion, destruction of the destoner, and blockage of the bed).

[0021] In a preferred embodiment, the insulated vessel of the process of the present invention is the calcination section of a shaft kiln, which further includes a limestone preheating zone at the top of the kiln and a CaO cooling section which also preheats the combustion air (oxygen-enriched air or an O2 / CO2 mixture, if a combustion continuum is preferred to maintain higher thermal energy and higher CO2 partial pressure during combustion or heating in step i).

[0022] The present invention also relates to a shaft kiln suitable for carrying out the calcination process described above and for calcining CaCO3 in the form of limestone to produce a mixture of CO2 and steam and CaO, the shaft kiln comprising: First vertical shaft, means for carrying out the combustion of fuel with preheated air or other combustion-sustaining mixture in the firing zone;

[0023] inlet and outlet gas pipes and solids pipes with switching valves, one of the inlet pipes having a switching valve for supplying steam to a lower portion of the calcination zone, and an outlet pipe having a switching valve for allowing extraction of product gas containing CO2 and steam, all configured to allow a continuous repeating series of at least a first operating period and a second operating period;

[0024] means for feeding batches of limestone through a solid valve and extracting lime in another solid valve during a first operating period lasting from 5 to 20 minutes during which fuel is burned in the calcination zone in a running mode, air or other combustion-sustaining mixture is preheated by CaO, and limestone is preheated by combustion flue gases flowing from the calcination zone;

[0025] means for sealing the gas inlet / outlet and solids valve during a second operating period lasting 5 to 20 minutes, except for a gas inlet and a gas outlet for supplying steam to the bottom of the calcination zone to extract the product gas containing pure CO2 and steam; Equipped with.

[0026] Optionally, the shaft kiln further comprises a second vertical shaft connected to the first vertical shaft, the second vertical shaft including a further flow of air entering the second vertical shaft to transfer heat from the lime cooling zone and the hot flue gases to the limestone preheating zone.

[0027] Optionally, the shaft kiln comprises an insulated vessel, the insulated vessel being the calcination section of either the first vertical shaft or the second vertical shaft, further comprising a limestone preheating zone at the top of the kiln and a CaO cooling section for preheating the combustion air or other combustion-sustaining mixture.

[0028] Alternatively, a shaft kiln is provided with means for effectively combusting fuel in its calcination zone with preheated air, enriched air or O2 / CO2, said kiln having insulated walls, inlet and outlet gas pipes with change-over valves and solids pipes, at least one of the pipes having a change-over valve connected to a supply of steam, and characterized in that it periodically alternates between the following operating periods:

[0029] i. a first operating period lasting 5 to 20 minutes during which fuel is burned in the calcination zone, with the combustion mixture preheated by the CaO product and the limestone preheated by the combustion flue gases flowing from the calcination zone, said first operating period having means for feeding a batch of limestone through a solids valve and extracting limestone in another solids valve; and

[0030] ii. A second operating period lasting 5-20 minutes during which the gas inlet / outlet and solids valves are sealed except that the gas inlet supplies steam to the bottom of the calcination zone and the gas outlet allows the evacuation of steam and CO2.

[0031] A particular variation when the combustion sustaining mixture is O2 / CO2 involves staging separate streams of O2 (from the air separation unit) and CO2 (recycled) to prevent carbonation of the CaO exiting the kiln.

[0032] If a double shaft kiln is used, this operates step i) by combustion in only one shaft using some of the air preheated in the bed of the limestone preheating zone, while regulating the flow of further air into the second shaft to transfer heat from the lime cooling zone and hot flue gases to the limestone preheating zone and use such heat to preheat some of the air entering the kiln from the top.

[0033] According to the teachings of the state of the art of gas burners in kilns, the distribution of air and fuel in the calcination zone of the kiln during step i) can be adjusted in many ways to accommodate the maximum permissible temperature in the kiln, as required by CaO quality requirements. However, it is important to understand that, since kilns usually burn fuel continuously, the introduction of step ii) in the method of the present invention for operating a lime kiln necessarily implies a reduction in lime production capacity proportional to the ratio between the duration of the time of step ii) and the total duration of one cycle of the first and second time periods (t1 + t2) of steps i) and ii). Therefore, a preferred embodiment to solve this is to increase the thermal energy input during step i) so that the thermal energy of combustion during step i) is the result of multiplying the average energy input into an equivalent kiln without CO2 capture by the ratio of the sum of the durations of steps i) and ii) divided by the duration of step i).

[0034] The lime kiln can also operate on solid fuel in step i) by mixing an appropriate proportion of solid fuel with the limestone feed. To minimize CO2 and other gas pollutant emissions during step i) involving the combustion of the solid fuel, a preferred embodiment uses a metal, such as Fe, Ni, or Cu, as the solid fuel, which is exothermically oxidized with air and contained in a second solid mixed with the CaCO3-containing solid. Vigorous premixing of the solid fuel with limestone results in a more controlled temperature profile and homogeneous combustion in the kiln's calcination zone. As shown in the examples below, the energy required for heating the solids (30 to 200 °C, as described above) is a small fraction of the energy required for their calcination, making it possible to maximize heating of the CaCO3-containing particles while minimizing calcination of the CaCO3 during the combustion step using metal fuel. Therefore, the amount of air required for oxidation of the metal fuel is small, and the maximum CO2 loss (caused by the equilibrium of CO2 on CaO at the maximum heating temperature) in such a small air mass is also small.

[0035] When the calcination process of the present invention is applied to a calcium looping process, the CaCO3-containing solids originate from the previous carbonation reaction step, which is characteristic of all calcium looping processes involving the carbonation of CaO to CaCO3. Because the calcium looping process is reversible, a continuous external supply of CaCO3 is no longer necessary. The CaCO3 formed during the carbonation step is primarily the CaCO3 that is calcined during the calcination step (note that there may also be a small percentage of CaCO3 from the purge or initial batch of CaCO3 or CaO). In these calcium looping systems, due to the limited reversible molar carbonation conversion ratios (0.03 to 0.1) of known CaO sorbents and the dilution of Ca-containing solids by catalysts and their supports, direct combustion of air and fuel within the solids bed is undesirable because localized flames and hot spots would compromise the integrity and activity of the catalyst and / or sorbent. Therefore, in these systems, it is important to perform heating step i) by indirect combustion of fuel, following the teachings of the packed-bed chemical looping combustion process.

[0036] Therefore, a preferred embodiment of the present invention relates to the application of the calcination method of the present invention to a calcium looping process using a packed bed, wherein a solid containing CaCO is formed from CaO and CO during a carbonation step in the presence of a second solid containing a metal such as Ni or Cu that acts as a catalyst and / or oxygen carrier, and wherein the combustion carried out in the heating step is a chemical looping combustion of a gaseous fuel using a combustion-sustaining gas by dividing the fuel combustion period of heating step i) in the following two sub-steps:

[0037] j) a first sub-step involving the oxidation of a second solid containing a metal such as Ni or Cu with a combustion-sustaining gas, the second solid being mixed into a bed of solids containing CaCO3; and

[0038] jj) a second sub-step comprising reducing the oxidized form of the metal fuel formed in sub-step j) by supplying a fuel gas to regenerate the metal fuel in the second solid while releasing concentrated CO2.

[0039] Carrying out substep j) at elevated pressure reduces the proportion of CO2 released in the flue gas and / or makes it possible to heat the bed of solids to a higher temperature, which is beneficial for step ii).

[0040] Thus, one such sub-step jj) is the regeneration of the metal catalyst, which must be in reduced form to maintain catalytic activity during the natural gas adsorption-enhancing reaction and steam or water gas shift that occur simultaneously with the CaO carbonation step.

[0041] In a preferred embodiment, the calcination process of the present invention is applied to a post-combustion CO2 capture process by calcium looping characterized by step iii) the additional carbonation of the CaO produced in step ii) by feeding combustion flue gas or any other gas containing CO2 to a packed bed of solids in order to increase the molar conversion of CaO to CaCO3 between 0.03 and 0.1 and to revive the CaCO3 present in the bed of solids at the start of step i).

[0042] In another preferred embodiment, the calcination process of the present invention is applied to a packed-bed adsorption-enhanced reforming process (SER). Such a process requires a second solid in the bed containing Ni, which is known to function as a metallic fuel (during step i)) and as a reforming catalyst during the SER process. Such an SER process includes step iii) the carbonation of the CaO produced in step ii) by feeding fuel gas and steam to a packed bed of solids containing CaO and Ni to catalyze the steam reforming of fuel gas to hydrogen, while increasing the molar conversion of CaO to CaCO between 0.03 and 0.1, and restoring the CaCO present in the packed bed of solids at the start of step i).

[0043] In another preferred embodiment, the calcination process of the present invention is applied to an adsorption-enhanced water-gas shift process carried out in a packed bed, characterized by a metal fuel and a Cu water-gas shift catalyst, and an additional carbonation step (iii) of the CaO produced in step (ii) by feeding a gas containing CO and steam to a packed bed of solids containing CaO and Cu to promote the water-gas shift reaction of CO contained in the gas to hydrogen and CO. The CO resulting from the water-gas shift reacts with the CaO, increasing the molar conversion of CaO to CaCO between 0.03 and 0.1, restoring the CaCO present in the bed of solids at the start of step (i). This particular embodiment may be particularly useful for decarbonating blast furnace gas from steel mills, where CO concentrations exceeding 20% ​​vol are not uncommon (see Example 2). Because the water-gas shift reaction is exothermic (-41 kJ / mol), additional heating of the packed bed of solids containing CaO and Cu occurs during the carbonation step (iii), necessitating a specific cooling or heat removal step to extract the energy released in such a reaction. To further improve the thermal efficiency of the process, a preferred embodiment is to carry out such cooling in a variant of the process that further includes an additional step of cooling a packed bed of CaO and Cu-containing solids designed solely to remove the heat released to the solids by the water-gas shift reaction of CO contained in the gas supplied in step iii). Since such additional cooling is typically an order of magnitude smaller than the cooling of the solids in step ii), such additional cooling can be achieved by heat transfer from the bed of solids to a heat exchanger embedded in the bed of solids. However, to minimize capital costs and simplify operation, such additional cooling step can be designed by using an equilibrium-limited endothermic reaction, such as the reverse water-gas shift (41 kJ / mol) or, preferably, a hydrocarbon steam reforming reaction with a high endothermic heat per mole of reformed gas (e.g., methane steam reforming with an enthalpy of 206 kJ / mol).

[0044] As mentioned in the previous paragraph, the CO2 partial pressure swing imposed on the gas atmosphere surrounding the solids during calcination step ii) can be achieved by supplying steam to the bed of calcined solids to achieve the effect of adiabatic cooling during calcination under the new partial pressure of CO2 around the solids. The use of steam has the disadvantage of requiring an additional condensation step to separate the CO2 into water and consuming energy for evaporation. However, it has the advantage of operating with minimal total pressure swing over the vessel, which facilitates the retrofitting of existing kilns to the method of the present invention. The use of steam in step ii) displaces the CO2 generated during calcination. Because CO2 is near equilibrium during step ii), the partial pressure of CO2 at the exit of step ii) changes over time. This means that if the inlet flow of steam is constant, the flow present during step ii) will change (decreasing) over time as the calcined solids undergo the adiabatic cooling provided by step ii). Such flow rate changes can alter the operation of downstream equipment (e.g., heat exchangers, CO2 purifiers, etc.), and they should preferably be avoided or minimized. Thus, a preferred embodiment of all methods and apparatus using steam injection to effect a CO2 partial pressure swing in step ii) is one in which the molar flow rate of steam supplied during step ii) is controlled to be proportional to the total flow rate of gas targeted at the gas outlet during step ii) multiplied by the difference between the total pressure during step ii) and the equilibrium partial pressure of CO2 over CaO, both divided by the total pressure. In this way, at the very beginning of step ii), the product gas from step ii) is pure CO2, and the concentration of CO2 in said product gas decreases with time until it reaches the equilibrium partial pressure of CO2 at the temperature resulting from the adiabatic cooling of 30-200°C that occurs during step ii).

[0045] In the process of the previous embodiment, it is possible to introduce a short additional gas purge / rinse step for safety and gas purity reasons. One particular additional step, which is valid for all disclosed embodiments, is characterized by introducing an additional gas elimination step at the start of step ii) in which a gas volume equal to 1 to 3 times the void volume of the gas in the packed bed of solids is evacuated. This action prevents the small amount of non-CO2 gas present in the interstices of the solid bed at the start of step ii) from mixing with the steam and pure CO2 generated during the calcination of CaCO3. [Brief explanation of the drawings]

[0046] The following diagrams reflect aspects for the actual implementation of the process and facilitate understanding of the key differences of the present invention relative to the state of the art, but they should not be considered limiting of the scope of the present invention. The diagrams use numerals to designate the physical elements, equipment, and materials necessary to carry out the method. Thick, black arrows represent the mass flow of solids and gases into and out of the equipment. The sequences characteristic of the steps of the preferred embodiment are indicated by Roman numerals i), ii), and iii) and last at least one time between a first time t1 and a fourth time t4, with transitions between steps indicated by thick, open arrows. The dashed boundaries represent the complete cycle N of the process of the present invention, including several steps and substeps per cycle. [Figure 1-1] FIG. 1 shows a single cycle (N) of the calcination process of the present invention, which comprises a series of heating steps consisting of the combustion of fuel (8) in air, enriched air, or O2 / CO2 or O2 (9), followed by the change in the partial pressure of CO2 around the calcined particles by supplying steam (100), and the calcination of a specific proportion of CaCO3 (4) contained in a solid containing CaO (5) to produce CO2 (7) diluted in the steam (100). [Figure 1-2]FIG. 1 (continued) shows a single cycle (N) of the calcination process of the present invention, which comprises a series of heating steps consisting of the combustion of fuel (8) in air, enriched air, or O2 / CO2 or O2 (9), followed by the alteration of the partial pressure of CO2 around the calcined particles by supplying steam (100), and the calcination of a specific proportion of CaCO3 (4) contained in a solid containing CaO (5) to produce CO2 (7) diluted in the steam (100). [Figure 2] Figure 2 shows the temperature evolution (Tcore) of a 7 cm limestone rock core preheated to 900 °C in pure CO2 and subjected to a CO2 partial pressure swing from the start of two firing experiments. The dotted line represents the case where the CO2 partial pressure swing is achieved by supplying preheated air at atmospheric pressure, and the solid line represents the case where a vacuum (less than 0.05 atm) is applied to the container with the rock. [Figure 3-1] FIG. 3 is a schematic diagram of the calcination process of the present invention as pure CaCO is calcined to CaO through seven CO partial pressure swings of combustion and steam injection cycles for each stone, feeding a batch of limestone (4) and extracting an equal batch of lime (5) in each cycle. [Figure 3-2] FIG. 3 (continued) is a schematic diagram of the calcination process of the present invention as pure CaCO is calcined to CaO through seven CO partial pressure swings of combustion and steam injection cycles for each stone, feeding a batch of limestone (4) and extracting an equal batch of lime (5) in each cycle. [Figure 4] FIG. 4 is a schematic diagram of a single shaft kiln for carrying out the process of FIG. 3 when the partial pressure of CO2 in the calcination zone is varied during step ii) by injecting steam (100). [Figure 5] FIG. 5 is a schematic diagram of a dual-shaft kiln for carrying out the process of FIG. 3 during the combustion-steam injection cycle, with combustion being the left shaft. [Figure 6] FIG. 6 is a schematic diagram of a dual-shaft kiln for carrying out the process of FIG. 3 during the combustion-steam injection cycle, with combustion being the right shaft. [Figure 7]Figure 7 is a schematic diagram of the calcination process of Figure 1 when the combustion of fuel (8) is carried out by chemical looping combustion by oxidizing the metal to MeO with air and regenerating the metal solid fuel (Me) from its oxidized form by reduction with fuel gas (8). [Figure 8-1] FIG. 8 is a schematic diagram of the process where each step of change in the partial pressure of CO2 is followed by an additional carbonation reaction step to regenerate the CaC03 decomposed during the steps of change in the partial pressure of CO2 by carbonation of CaO using CO2 contained in the gas (13) which reacts with a portion of the CaO (5) in the bed to give CaC03 (5). [Figure 8-2] FIG. 8 (continued) is a schematic diagram of the process where each step of change in the partial pressure of CO2 is followed by an additional carbonation reaction step to regenerate the CaC03 decomposed during the steps of change in the partial pressure of CO2 by carbonation of CaO using CO2 contained in the gas (13), which reacts with a portion of the CaO (5) in the bed to give CaC03 (5). [Figure 9] FIG. 9 is a schematic diagram of a packed-bed temperature and pressure swing system implementing the process of FIG. 8 for enhanced water-gas shift of CO(13)-containing gases using CaO as the CO2 adsorbent. DETAILED DESCRIPTION OF THE INVENTION

[0047] Referring now to the accompanying drawings, we define preferred embodiments of the present invention and provide evidence supporting the claims for a state-of-the-art model of calcination and calcium looping reactions. Figure 1 depicts a series of process steps for the present invention. All other embodiments disclosed below include the steps depicted in Figure 1, and the goal of the present process is to calcinate a solid containing CaCO3 (4) to produce a pure stream of CO2 (7) from the decomposition of CaCO3 to CaO (5). In this invention, the problems associated with the use of vacuum swing are avoided by conducting the calcination step by injecting steam into a bed of solids containing CaCO3 preheated to approach the equilibrium calcination temperature in pure CO2. The novel method is illustrated in Figure 1, which depicts two steps of the present invention for a single particle containing CaCO3 (optionally located within a central core (4) as a black circle) partially calcined to a particle with a higher CaO (5) content marked in gray. The particle is located in an insulated container (1) with a metal enclosure (2) and a refractory lining (3). The insulated vessel (1) allows for the combustion of fuel (8) with air or other combustion-sustaining mixture (9) within a bed of solids containing CaCO₃, emitting flue gases (10, 11) containing CO₂ (11) and other nitrogen-rich gases (10) when the combustion-sustaining gas (9) for combustion is air. An important aspect of the invention is that for each particle in the bed, calcination occurs gradually, in cycles N-1, N, N+1, etc. This means that multiple cycles are required to achieve complete molar conversion of an individual limestone particle from CaCO₃ to CaO. For simplicity, only four of these cycles (N=1 to N=4 for each particle) are shown in Figure 1. In a continuous process where N proceeds indefinitely through the continuous feeding of limestone particles to the insulated vessel used for calcination and the continuous extraction of an equal amount of limestone particle stream, individual particles enter vessel (1) in cycle N and leave the vessel in cycles greater than N+4, achieving complete calcination as illustrated in Figure 1.

[0048] Alternatively, the CaCO content in the original solid prior to partial calcination in cycle N could be the result of a previous carbonation process step (frame N-1 is not shown in FIG. 1 for simplicity). This is common practice in existing calcium looping processes for CO2 capture, where individual CaO sorbent particles ideally function in an infinite number of reversible carbonation-calcination cycles. For simplicity, only four of these reversible cycles are shown in FIG. 1 (N to N+3), illustrating only two characteristic steps of the present invention for each cycle.

[0049] The process of the present invention therefore involves the partial calcination of a CaCO3-containing solid in repeated cycles N, each cycle characterized by a series of at least these two steps, as shown in the upper part of Figure 1.

[0050] i) a heating step involving fuel combustion, lasting a first time period (t1) preferably between 5 and 20 minutes, in which the fuel (8) is combusted in the bed of solids using air or enriched air, or other combustion-sustaining mixtures such as O2 / CO2 or O2 (9) to heat the bed of solids containing CaCO3 to an average temperature of 800-950° C. The design goal of this combustion step is to only heat the solids with minimal calcination.

[0051] ii) A CO2 partial pressure swing step, lasting a second time (t2), preferably 5-20 minutes, by feeding a stream of steam (100) into the bed of solids to promote the calcination of CaCO3 (4), resulting in pure CO2 (7) diluted with steam (100) and CaO (5) as reaction products, causing a temperature drop of 30-200 °C within the bed of solids. As discussed in the state of the art, any sudden CO2 partial pressure swing (e.g., by replacing the gas atmosphere around the calcined particles with a calcination reactant such as steam by vacuum or inert gas) will cause adiabatic cooling during the adiabatic calcination of CaCO3-containing solids.

[0052] The maximum temperature drop at the end of the adiabatic cooling in step ii) is limited by the adiabatic temperature of the calcination, given the partial pressure of CO in the CO / steam gas mixture surrounding the calcined particles. However, the effective temperature drop also depends on many factors related to the calcination rate of the CaCO-containing solid (e.g., porosity of the solid, particle size, etc.), which affect the actual partial pressure of CO at the calcination reaction front, given the resistance the CO flow experiences leaving the calcination front to reach the bulk gas space between the particles. Because the CO partial pressure swing is achieved by supplying steam to the calcined solid at near atmospheric pressure, the partial pressure of CO in the product gas quickly adjusts to the equilibrium partial pressure of CO on CaO at the temperature resulting from the adiabatic cooling of the solid. For example, if the solid in step i) is heated to 900°C, the equilibrium temperature of pure CO on CaO, CO evolved from the calcination will initially tend to displace any other gas in the free volume of the bed of CaCO-containing particles. Injection of a small steam stream is sufficient to initially draw such CO2 gas from the bed that the partial pressure of CO2 in the gas stream exiting the reactor in step ii) is close to 1. In contrast, towards the end of step ii), when adiabatic firing results in adiabatic cooling of the solid bed from 30 to 200 °C, the partial pressure of CO2 in the product gas (CO2 (7) stream + steam (100) stream) is lower (e.g., 0.09 is the partial pressure in atm of CO2 over CaO at 750 °C, reached if adiabatic cooling was initially 150 °C versus 900 °C).

[0053] As is known in the state of the art, the effect of CO2 partial pressure swings around calcined particles is more similar when applying a vacuum pressure swing than when changing the CO2 partial pressure around the calcined solid by replacing the CO2 from the calcination with a gas that is inert with respect to the calcination. For example, Figure 2 shows that in the first 5-10 minutes after a rapid swing to reduce the CO2 partial pressure around a limestone rock preheated to 900°C, there is very rapid cooling inside the rock when the CO2 partial pressure around the rock is reduced by switching to airflow (i.e., replacing the CO2 that originally surrounded the 900°C rock) or by applying vacuum pressure to the container in which the rock is housed (T core (decrease with time). The temperature of the rock core, T in Figure 2 core reaches a stable value after about 20 minutes. At such a stable value, the adiabatic cooling caused by the firing of CaCO3 at the firing front of the unfired core must be equal to the heat transferred from the stone periphery to the firing front. This is true when CO2 is replaced by air (T core When vacuum is applied (T core This explains the difference in cooling measured between the two samples (the temperature drops to about 795 °C). Heat transfer towards the calcined stone is probably more effective when air is flowing through the vessel containing the stone than when a vacuum is applied. However, Figure 2 shows that any partial pressure swing that reduces the CO2 partial pressure around a mass of preheated CaCO3 so that it is in equilibrium with CO2 leads to a certain adiabatic cooling of the CaCO3. These results are consistent with the observations by Hills (1968, Figure 3) working with smaller (1 cm diameter) particles.

[0054] Using a heat transfer model for calcining stone, it has been demonstrated that in reasonable heating times (5–20 min), it is possible to transfer the necessary heat to the CaCO₃-containing solid to heat it 30–200°C and return it to an average temperature of 800–950°C. Such a temperature is required to restart a new CO₂ partial pressure swing cycle, in which the temperature again drops 30–200°C in the solid (causing a decrease in CaCO₃ molar content between 0.01 and 0.07). The calcination rate under reduced CO₂ partial pressure in step ii) is fast enough to complete incremental calcination molar conversions of 0.01 and 0.07 in only a 5–20 min CO₂ partial pressure swing step, at the end of which the CO₂ partial pressure in the product gas (containing CO₂(7) and steam (100)) reaches 0.05–0.5 atm CO₂. The reduction in CO2 partial pressure at the end of step ii) has the advantage of a larger decrease in temperature during calcination and a reduction in the number of cycles of steps i) and ii). However, this advantage comes at the expense of higher steam consumption to produce such a low CO2 partial pressure, by using more operating time to extend the calcination period until the point at which the CO2 partial pressure decreases and the steam partial pressure in the product gas increases. Moderate CO2 partial pressure swings require higher heating temperatures of the carbonated solids and / or a greater number of CO2 partial pressure swing cycles, because the temperature decrease of the solids is limited by the equilibrium temperature imposed by the CO2 partial pressure swing (e.g., for a CO2 partial pressure of 0.5 atm, lowering the temperature below 850°C during step ii) is not possible because this is the equilibrium temperature of CO2 on CaO at such a CO2 partial pressure).

[0055] Taking into account the aforementioned constraints, a preferred embodiment of the process of the present invention allows for the continuous calcination of the limestone (4) flow and the extraction of an equivalent volumetric flow of lime (5). A schematic diagram of each single cycle N of the process is shown in FIG. 3. In this case, the vessel (1) of the process contains a large number of particles arranged in a packed bed. In the example of FIG. 3, batches of limestone (4) are periodically charged to the top of the vessel after step ii) and an equal amount of batches, relative to the volume of fully calcined material, is extracted at the bottom of the bed. This allows the preheating of the limestone solids to the temperature required at the start of step i) by the sensitive heat of the combustion flue gases (10 and 11) generated during step i), as well as the preheating of the large flow of air or other combustion-sustaining mixture (9) in contact with the hot CaO leaving the insulated vessel (1). Each cycle N lasts for the sum of the first time (t1) under the combustion step and the second time (t2) of the CO2 partial pressure swing step induced by supplying steam (100), plus the third time (t3) required to displace the limestone batch (4) and extract the CaO product (5). Optionally, the third time (t3) can be zero if such operation of discharging solids is carried out simultaneously with the combustion or heating step i).

[0056] Figure 4 shows a schematic representation of a complete system for carrying out the process of the present invention in a limestone shaft kiln. The kiln includes a refractory-lined outer wall (2) for maintaining high-temperature operation within the stone bed, a lock hopper connected to a solids feed valve (24) and an extraction valve (25), and several gas switching valves (101, 21, 22, 23, 201) installed on all inlet and outlet gas pipes. At least one of the pipes (usually the flue gas (10, 11) pipe) includes a T-pipe bypass with a valve (201) for bypassing product gases (7, 100) during firing. Such kilns typically include a preheating zone (30) for limestone (4) at the top of the kiln and an air preheating section (31) at the bottom of the kiln, where the limestone product (5) is cooled with air or other gases in the combustion-sustaining mixture (9). This process requires a moving bed step iii) after step ii), in which a bed of solids moves downward through the kiln by extracting a batch of CaO product (5) at the bottom of the air preheating section of the kiln while charging a batch of limestone (4) at the top of the limestone preheating section of the kiln. The process operates in a series of operating periods characterized by the positions of the switching valves as shown in Figure 4. During step i), lasting from t1 = 5 to 20 minutes, valve (21) is open to release CO2 (11) produced during fuel combustion mixed with flue gas (10) from the combustion, valve (22) is open to admit fuel (8), and valve (23) is open to admit air, enriched air, or combustion-sustaining gases containing O2 / CO2 or O2 (9). Valve (201), which diverts product gases (7, 100) during calcination, valve (24), and valve (25) are closed. The period of step i) is followed by a second period t2 of 5 to 20 minutes, step ii), during which all valves are closed except for the valve (201) for diverting the product gases (7, 100) during calcination and the steam valve (101) for supplying steam (100) to the bottom of the calcination zone (1). The steam (100) flow must enter the kiln at a sufficiently high temperature (>550 °C) to prevent the risk of hydration of CaO.For simplicity, in this configuration, step iii) is performed for a short period of time when all valves except the solids feed valve (24) and the extraction valve (25) are closed. Minor modifications to this strategy can be implemented to overlap a certain period between steps i) or ii) with step iii). Other minor variations of the process can relate to the feed point of the combustion-sustaining mixture (9) to avoid recarbonation of CaO, for example, if (9) is a combustion-sustaining mixture of O2 / CO2, following the teachings of other oxygen-fired shaft kilns.

[0057] Figures 5-6 are schematic diagrams of a dual-shaft kiln operating with the process steps characteristic of the present invention shown in Figure 3. The number of valves and other minimal mechanical elements has changed relative to the single shaft of Figure 4, but the process remains the same. For simplicity, Figures 5-6 show a system of rotary valves (241 and 242 for charging, 251 and 252 for discharging, operating on either the left or right shaft as shown in Figures 5 and 6) chosen to regulate the charging and discharging of solids during step i) (i.e., step iii) is no longer a separate step as in Figure 4). These valves, along with a lock hopper, allow fuel to be burned in one of the shafts during step i) or to provide steam during step ii) while preventing gas leakage through the solids feed and extraction lines. The process operates in a series of operating periods characterized by the position of the switching valve, as shown in Figure 5 when combustion step i) occurs in the left shaft and as shown in Figure 6 when combustion step i) occurs in the right shaft. Also, for simplicity, it has been assumed that the combustion-sustaining gas (9) used during combustion step i) is air. From here on, it is up to those skilled in the art of oxy-combustion in shaft kilns to adapt the drawing to kilns using other combustion-sustaining gases (9), such as enriched air, O2, or O2 / CO2, in order to maintain the high thermal efficiency characteristic of double shaft kilns.

[0058] FIG. 5 shows the period of step i) lasting a first time t1=5-20 minutes, during which valves (231, 233) are opened for the inflow of air (9) and valve (221) for the inflow of fuel (8) to allow combustion of fuel (8) in the left shaft, while the right shaft operates in regeneration mode, valve (232) is opened for the inflow of air (9) and valve (212) is opened for the discharge of CO2 (11) mixed with the remaining flue gases (10) from the combustion. Valve (201) for diverting product gases (7, 100) during calcination, valve (211) in the left shaft for venting flue gases including CO2 (11) and flue gas (10), valve (234) for admitting air to the top of the left shaft, rotary valve (242) for preventing gas leakage through the solids extraction line, and rotary valve (252) for preventing gas leakage through the solids supply line in the right shaft are all closed. Combustion step (i) occurs only in the left shaft, followed by step (ii), which lasts for a second time period (t2) of 5 to 20 minutes with all valves closed except for valve (201) for diverting product gases (7, 100) during calcination and valve (101) for allowing the supply of steam to create the CO2 partial pressure swing.

[0059] Figure 6 shows the cycle of steps i) and ii), characterized by the next combustion period in the right shaft. Following the known operating principle of double-shaft kilns to achieve higher thermal efficiency, the new combustion step i) involves the combustion of fuel (8) in the shaft located on the right side of the kiln, with the following valves opened: rotary valve (242) preventing gas leakage through the solids supply line, rotary valve (252) preventing gas leakage through the solids extraction line, valves (231, 232, 234) for the inflow of air (9), valve (222) for the inflow of fuel (8), and valve (211) for the release of CO2 (11) mixed with other flue gases (10) in the combustion-sustaining mixture. The other valves, valve (202) for diverting the product gas (7, 100) during calcination, valve (241) preventing gas leakage through the solids supply line in the left shaft, rotary valve (251) preventing gas leakage through the solids discharge line, and valves (201, 202) allowing the extraction of CO2 (7) and steam (100), are closed during such step i). The CO2 partial pressure swing of step ii) is applied to the calcination zone (1) in the right shaft by supplying steam (100) while extracting the gas product (composed of CO2 (7) and steam (100)) by opening valve (102) for the steam supply and valve (202) for the CO2 (7) and steam (100) extraction. The remaining valves are closed during such step ii).

[0060] Note that the kilns of Figures 4-6 can recover CO2 (7) after condensation of the steam (100) (heat exchangers or condensers for such downstream processes are omitted for simplicity). However, these methods cannot recover CO2 generated during the combustion of carbon contained in the fuel (8). Such CO2 (11) is released mixed with the nitrogen-enriched flue gas (10) generated during the combustion of the fuel (8) with air or other combustion-sustaining mixture (9) (unless such combustion-sustaining mixture is pure O2 and CO2).

[0061] The application of the process of the present invention to the calcination of limestone in a shaft kiln is not limiting. Several calcium looping CO2 capture processes exist that require the calcination of carbonated solids (formed during the carbonation of CaO with CO2 in flue gas or fuel gas) to produce a pure CO2 stream. These processes can benefit from the calcination sequence disclosed in this invention. In particular, a calcium looping sequence involves cyclic operation in an adiabatic packed bed (1) of stationary solids (4, 5), reversible carbonation of CaO (5) to form solids containing CaCO3 (4), followed by calcination of the formed solids containing CaCO3 (4) to produce concentrated CO2 (7), and CO2 capture by regenerating the CaO adsorbent (5). An example of such a process is the adsorption-enhanced steam reforming of natural gas and other hydrocarbons (including biogas), which uses CaO as an adsorbent to displace the reforming reaction to hydrogen production. Also, adsorption-enhanced water-gas shift using CaO as an adsorbent displaces the water-gas shift equilibrium to H2 production. In order to understand the relevance of the present invention to the technical feasibility of such a calcium looping process, it is important to highlight the following aspects here.

[0062] It is widely known that the ability of naturally sourced CaO as a fully reversible CO2 adsorbent is limited to a maximum calcium conversion to CaCO3 of only 0.03 to 0.1. Furthermore, the necessary mixing of the Ca-based solid with a second solid containing a reforming or water-gas shift catalyst and its support reduces the effective weight content of active CaO in the bed to only 1-5% w of the total bed, with the remaining 95-99% of the bed mass becoming a thermal stabilizer during the carbonation or calcination reaction steps in the calcium looping process.

[0063] The optimal temperature range for rapid and effective carbonation of CaO in CO2 capture systems using calcium looping at atmospheric pressure is typically 600-700°C, which also allows for high CO2 capture efficiencies (>90%) due to the low partial pressure of CO2 in equilibrium with CaO (e.g., 0.012 atm at 650°C). Increasing the operating pressure allows operation at higher temperatures, but at the expense of higher costs and energy requirements. On the other hand, rapid and effective calcination of CaCO3 requires temperatures approaching 950°C in a pure CO2 atmosphere, and despite the swing in CO2 partial pressure during step ii), such a process must begin at temperatures between 800-950°C, as discussed above.

[0064] Therefore, even a small heat loss in the system makes it impossible to reach the calcination temperature at the start of step ii) by relying solely on the energy released during the carbonation of CaO in step iii). This is further exacerbated when considering the heat balance affecting the incoming cold gas reactants and hot gas products, which typically requires additional heat input to the system. Therefore, the method of the present invention is a solution for narrowing the heat balance and temperature gap between the carbonation and calcination temperature requirements in the calcium looping process. By utilizing the role of 95-99% of the remaining bed mass as a thermal stabilizer that can store the sensitive heat released during step i) and provide the energy required for calcination during step ii), the method of the present invention can be demonstrated to enable the use of known CaO adsorbent materials, i.e., CaO adsorbent materials with an effective weight content of active CaO in the bed of only 1-5% w / w of the total bed (see Example 2).

[0065] Considering the experimental results and the above discussion, a preferred embodiment of the present invention (see FIG. 7) is characterized in that combustion step i) is carried out according to the chemical loop combustion principle to combust fuel (8) with air (9) using a solid oxygen carrier (MeO / Me) mixed in a bed of solids containing CaCO3. Combustion step i) in a bed of carbonated solids is carried out in two separate substeps:

[0066] j) a first sub-step lasting a fourth time (t4) less than the first time (t1), in which the metal (Me) contained in the second solid (12) is oxidized to MeO by reacting with the combustion-sustaining gas (9) and releasing flue gas (10) (which is enriched with nitrogen when the combustion-sustaining gas is air); and

[0067] jj) The difference between the first time (t1) and the fourth time (t4), i.e., t1-t4, continues with a second sub-step in which the MeO produced in step j) is reduced back to metal (Me), while the fuel (8) is oxidized to combustion products rich in CO2 (11).

[0068] This sequence of sub-steps j) and jj) results in the heating of the bed of CaC03-containing solids in a more uniform and distributed manner than would be the case if direct combustion of fuel (8) with air (9) were carried out in the bed in a single step i). This is important in these applications because, due to the low CaC03 content of the bed (less than 5% w of active CaC03), the appearance of flames and other hot spots in the bed would lead to undesirable localized calcination and localized overheating of the CaC03 during step i), while not sufficiently heating other parts of the bed during step i) and preventing the targeted calcination and CO2 production during step ii).

[0069] As shown in Figure 8, in accordance with the teachings of chemical looping combustion in a packed bed, the oxidation reaction of metal (Me) (open circle) to MeO (marked with a cross on the open circle) proceeds with a neat reaction front that moves upward as process step j) progresses over time. From similar processes, it is known that this oxidation step can be a very short reaction step (the fourth time (t4) is 1-3 minutes) that heats up the entire bed of solids behind the oxidation reaction front, depending on the nature and mass content of the metal (Me) being oxidized to MeO (see Example 2 below). Thus, the process in Figure 8 follows the sequence in Figures 1 and 7 in the sense that a partial calcination of CaCO3 occurs during each cycle of steps i) and ii) (i.e., a reduction in the CaCO3 mole fraction of the bed's Ca adsorbent by 0.01 to 0.1 after the sum of the first time (t1) and the second time (t2). The first time (t1) is the sum of the fourth time (t4) and the difference between the first time (t1) and the fourth time (t4), i.e., t1 - t4). For simplicity, the carbonation step and any other reaction or bed conditioning steps performed in the solid bed are assumed to continue for a third time (t3), reversibly returning the bed to its initial state to restart a new cycle in the calcium looping process. Such a carbonation reaction step is carried out by feeding a carbon-containing gas (13) into the packed bed of solids. During such a process, a carbonation reaction front develops in the bed as the CO2-containing gas encounters a layer of CaO that has not yet been carbonated to CaCO3. Such carbon-containing gases can be flue gas containing CO, or fossil fuel gas or biogas that can be reacted with steam (14) to produce CO during reforming and water-gas shift reactions. In both cases (i.e., adsorption-enhanced reforming or adsorption-enhanced water-gas shift), as noted above, there is a modest increase in the molar conversion of CaO to CaCO between 0.01 and 0.1 when using low-cost CaO material from natural limestone.The carbonation step can be carried out at high pressure by supplying gases (13) and (14) at high pressure to heat the bed of CaC0-containing solids to the initial temperature of step ii), which accelerates the carbonation of CaO, increases CO2 capture efficiency, and allows for a higher temperature of the bed of CaC0-containing solids at the end of the carbonation step, thereby reducing the fuel requirements for step i). The application of a sequence of steps i)-ii), with step i) subdivided into steps j) and jj), makes it possible to bridge the temperature gap between the desired reaction steps in these calcium looping processes, as discussed in Example 2 below.

[0070] Example 2 is presented below by using a preferred embodiment of the invention shown in Figure 9. It is characterized in that the combustion carried out in (1) during the heating step is the indirect combustion of fuel (8) with air (9) by utilizing the dual role of Cu as an oxygen carrier for the indirect combustion of fuel (8) with air (9) in step i) and as a catalyst for the adsorption-enhanced water-gas shift reaction. Variations of this embodiment, also included in the invention, are the adaptation of other adsorption-enhanced processes using Ni or other catalysts (collectively referred to as Me and MeO in Figures 7 and 8). Another variation affects the selection of pure O or an O / CO mixture as the combustion-sustaining gas (9) in step i).

[0071] In all previous embodiments of the process, for safety and gas purity reasons, a short additional gas purge / rinse step can be introduced, for example using additional steam or nitrogen as inert gas. [Example]

[0072] Mass and energy balances are elucidated in the following three examples to illustrate the application of the calcination method disclosed in this invention to the calcination of CaCO3-containing solids to a pure stream of CaO and CO2 in a shaft lime kiln (Example 1), and to a calcium looping process for calcium-enhanced water gas shift (CASOH) catalyzed by Cu to produce H2 from blast furnace gas and pure CO2 (Example 2). For simplicity, the following thermodynamic parameters are assumed to be constant with temperature and represent the enthalpies involved in the temperature interval considered in the examples: enthalpy of calcination and carbonation = ±170 kJ / mol Ca, enthalpy of WGS = -34 kJ / mol CO, enthalpy of steam methane reforming = 190 kJ / mol Enthalpy of CH4 and Cu oxidation = -145kJ / molCu, enthalpy of CuO reduction = -54.5kJ / molCu (when using CH4), -100kJ / molCu (when using H2), -140kJ / molCu (when using CO), enthalpy of blast furnace gas combustion = 2600kJ / kg, enthalpy of methane combustion = 50,000kJ / kg, heat capacity of all solids = 0.8kJ / kgK, heat capacity of blast furnace gas = 1.35kJ / kgK, heat capacity of H2 / N2 product gas from SEWGS = 1.6kJ / kgK, heat capacity of CH4 = 4.1kJ / kgK, heat capacity of H2O = 2.2kJ / kgK, heat capacity of remaining gas = 1kJ / kgK. The displacement of the reaction and heat transfer fronts is calculated assuming infinite reaction rate and infinite solid-gas heat transfer rate. Those skilled in the art can modify the calculations provided below and adapt them to other conditions. Therefore, the examples are illustrative only and are non-limiting.

[0073] Example 1 (lime vertical shaft kiln) Using Figure 4 as a reference, this example is calculated to demonstrate the application of the method to a lime single-shaft kiln designed to generate an average target flow rate of 1 kg CO₂ / s (7) during step ii), using a first period (t1) of 10 minutes duration for the combustion step and a second period (t2) of 15 minutes duration for the CO₂ partial pressure swing step. This example may also be adapted to a double-shaft kiln, such as those shown in Figures 5 and 6. In fact, the selected duration of step i) coincides with the typical combustion period of a double-shaft kiln to achieve similar goals regarding heat transfer efficiency between gas and solids. Initially, the efficiency of CO₂ recovery during step ii) is assumed to be 80% (the remaining CO₂ generated from the calcination is considered to be lost with the flue gas during step i). This means that there is a flow of 1.70 kg / s of limestone (assumed to be pure CaCO₂) entering the kiln. The combustion performed during heating step i) (represented by two flames in the kiln's calcination zone (1)) is assumed to be within the calcination zone. In this example, it is assumed that the supply of limestone (4) and the removal of lime (5) occur before the combustion step i) (see Figure 4). Complete calcination of limestone to lime (0.95 kg CaO / s) requires a minimum of 3.0 MJ / kg of CaO, corresponding to a minimum average energy requirement for calcination of 2.85 MW during the entire cycle. In this example kiln, if the lime flow rate is maintained, heat input must be supplied to the shaft kiln only during the combustion step i) for a shorter period of time. This means that a calcination heat input of 7.12 MW is required during step i). To resolve the energy balance of the limestone and air preheating zones (30, 31), it is assumed that the limestone (4) and oxygen-enriched air are supplied at a temperature of 20°C, while the lime (5) and flue gases (10, 11) exit the kiln at 50°C. Furthermore, a temperature of 950°C is assumed for the gas and solids streams exiting the calcination zone (1). Therefore, a methane flow rate of 0.17 kg / s is calculated to be sufficient to compensate for the firing energy requirements, the 10% characteristic value accounting for heat losses, and the sensitive heat leaving at the outlet of CaO and flue gas mass flow rate.A low excess of 5% and an oxygen concentration of 45% v in oxygen-enriched air (9) were used in the above calculations.

[0074] The assumption of enriched air with an oxygen concentration of 45%v promotes mass efficient closure and energy balance in a single shaft kiln, since in this particular example the gas flow rate at the outlet of the calcination zone contains enough energy to preheat the incoming limestone stream to said calcination zone. Assuming a typical superficial gas velocity of the combustion flue gases in the calcination zone of 1.5m / s, this corresponds to a flow rate of 4.5m / s. 2 This translates to an average superficial velocity of solids moving down the kiln during each cycle of 0.0004 m / s, a void fraction unoccupied by stone of 0.4, and an average density of solids of 2000 kg / m 3 Assuming a total height of 8 m for the calcination zone (1, similar to the other kilns), we can estimate a total mass of solids of about 29,000 kg in this calcination zone of the kiln. This translates to a total residence time of the solids in the calcination zone (1) of 6.2 hours. Such a long residence time in the calcination zone is within the range of residence times for stones in existing kilns and can be adjusted to fit the stone particle size and stone calcination rate parameters. Considering known models of limestone calcination rates, it is shown below that this is sufficient to achieve a degree of calcination close to 100%.

[0075] During each stage of the kiln's combustion step (i), a preliminary heat and mass balance indicates that approximately 115 kJ / kg of solids in the calcination zone are absorbed on average during each combustion step (i). After the completion of this combustion period, to initiate step (ii), the valve (22) for the inlet of fuel (8), the valve (23) for the inlet of air (9), and the valve (21) for the release of CO2 (11) mixed with nitrogen-enriched flue gas (10) are closed, and the steam valve (101) for supplying steam (100) to the bottom of the calcination zone (1) and the valve (201) for diverting the product gases (7, 100) during calcination are opened. The 15-minute duration of step (ii) remains consistent with known calcination models, allowing for the conversion of 6.9% wt. of CaCO3 to CaO during step (ii). Therefore, a minimum of 12 combustion / steam partial pressure swing cycles are required for the stones in the calcination zone to achieve their full calcination. Since each cycle lasts for the first time (t1) plus the second time (t2), i.e., t1 + t2 = 25 minutes, the cumulative time required for the particles to be fully calcined is 4.9 hours. This calcination time is less than the residence time of the solids in the calcination zone, which is estimated to be over 6.2 hours. Therefore, under these assumptions, ensuring full calcination of the limestone fed to the kiln is consistent with the state of the art for existing lime kilns.

[0076] A temperature drop of 115°C of the solids is estimated during the vapor partial pressure swing step. This results in a decrease in the equilibrium partial pressure of CO2 from 1.0 atm at the beginning of the step to 0.17 atm at the end. Furthermore, the cooling of the solids during this step in the calcination zone is assumed to be constant with time. A total of 900 kg of CO2 is calcined in each CO2 partial pressure swing step. This results in a 1.1 m2 CO2 at the exit of the calcination zone (1). 3 0 to 0.89 m to maintain a total flow rate of N / s 3 A variable steam flow (100) of up to N / s is required, which corresponds to an average gas velocity of 1.0 m / s in the calcination zone.

[0077] For an equivalent single lime shaft kiln without CO2 recovery, with the same gas and solids stream inlet temperatures and heat losses, and operating under the same conditions in the calcination zone, when air is used as the combustion feed, a maximum heat input of 3.8 MW can be calculated, resulting in a maximum lime flow of approximately 0.84 kg / s. This shows that the method disclosed in this invention for obtaining a pure CO2 flow of 1 kg / s when oxygen-enriched air is used as the combustion feed can be implemented in an existing shaft kiln without reducing lime production capacity, and the kiln is retrofitted in a similar manner as described above with the additional equipment required for steam injection (100), steam condensation (not shown in Figure 4 for simplicity), and switching valves as shown in Figure 4.

[0078] Example 2 (CO using water vapor) 2 CASOH process by pressure swing firing) The method of the present invention applies to a calcium-enhanced water-gas shift process (CASOH) for treating blast furnace gas (BFG) from a steel plant, producing H-enriched gas while carbonating CaO to CaCO in a packed bed. This process is carried out in a system consisting of several adiabatic packed-bed reactors operating in parallel with different reaction steps to enable continuous processing of the gas stream. Figure 9 is used as a reference. As in the previous example, an average flow rate of 1 kg / s of pure CO (7) and steam (100) is recovered during step ii) and used as a basis for calculations. It is also assumed that the BFG consists of 23% v CO, 24% CO, 5% H, and 48% N. The CaO-based material has 4.5% w active CaO. In this particular example, the Cu-based material also present in the bed has 5% w Cu on an inert support. It is assumed that Cu can be fully oxidized to CuO in substep j) and reduced to Cu in step jj), as shown in Figure 9. These material compositions are available in the state-of-the-art chemical looping combustion and water-gas shift catalytic processes. For the calculation, the adiabatic packed-bed reactor (1) has a length of 10 m, an internal diameter of 2.2 m, and a bed density of 1500 kg / m. 3 It was assumed that this was the case.

[0079] The reactor contains a mixture of solids with a molar ratio of active Cu / CaO components of 0.10 when the CaO adsorbent is combined with a Cu catalyst containing 5% w Cu, which means 4.1% w active CaO, 89.6% w inactive CaO, 0.3% w Cu, and 5.9% w Cu catalyst support.

[0080] Since the CASOH process operates cyclically, the description can begin with any step. Since application of the method of the present invention (steps i) and ii) in Figure 9) requires a packed bed of solids containing CaCO3, we will first describe the step involving the production of CaCO3 in the bed (CASOH step, represented as iii in Figure 9). When such a calcium-enhanced water-gas shift reaction begins, the bed is loaded with approximately 57,000 kg of solids and is at 770°C as a result of the previous calcination step (see below). To run the CASOH step, valves (22) and (23) are opened to allow the supply of steam (14) and BFG (13), respectively. A BFG (13) flow of 2.3 kg / s (i.e., 6 MW LHVThe heat input (13) is mixed with 0.63 kg / s of steam (14) to obtain a steam-to-CO molar ratio of 2 in the combined feed to the reactor. Valve (21) is also open to allow extraction of the CASOH H2-enriched product gas (15). For simplicity, we assume that the gas mixture (13, 14) is fed to the reactor at the same temperature as the solids (770 °C) and 5 bar. These conditions allow for a very fast water-gas shift reaction between CO and HO, catalyzed by the Cu-based solids, to form CO2 and H2. As CO2 is produced, it reacts with activated CaO to form CaCO3, which shifts the water-gas shift equilibrium toward more vigorous production of H2. The enhanced water-gas shift of CO in the presence of CaO is exothermic. For simplicity, we calculate the temperature rise of the reaction front and the rate at which both the reaction and heat exchange fronts progress through the packed bed, assuming a well-defined reaction front and negligible resistance to mass and heat transfer during operation. When these gases reach the region containing activated CaO solids, exothermic carbonation occurs with a sharply defined carbonation front, which reaches a maximum temperature of 902 °C. With a low content of activated CaO in the bed, the carbonation front advances quickly, leaving the carbonated solids behind at 902 °C, while the product gas (15) obtained during this step iii) (product gas flow rate: 1.52 kg / s, consisting of 24.8% v H2, 2.5% v CO, 2.6% v CO2, 46.9% v N2, and 23.2% v H2O) is discharged at 770 °C (i.e., the initial temperature of the packed bed). After approximately 25 minutes of operation under these conditions, the activated bed of solids reaches full carbonation, and the packed bed is brought to 902 °C. Since the gases (13 and 14) are assumed to enter at the same temperature, there is no temperature profile at the bottom of the reactor.

[0081] As shown in Figure 9, this example method begins with step i) of heating a bed of solids containing CaCO3 by combustion of fuel (8) (in this example, BFG is also used for this purpose, but other fuel gases are suitable for the heating step). The combustion of fuel gas (8) in step i) is divided into two substeps, j) and jj). In substep j), oxidation of the Cu-based solids to CuO is carried out using pure oxygen (9) (the oxidation gas can be air, enriched air, or O2 / CO2). Due to the low Cu content in the reactor as described above, the oxidation front progresses very quickly, leaving behind the generated heat. For simplicity, we assume that the reactor is fully adiabatic, that the heating of the bed of solids is completed using the energy released by the oxidation of Cu to CuO, and that no axial or radial temperature profile is generated. The highly exothermic nature of the Cu oxidation reaction and the moderate Cu content selected result in a temperature rise in the packed bed of approximately 8°C. The high temperatures achieved at the oxidation front (i.e., 910 °C) and in the solids remaining after said oxidation reaction front ensure the complete conversion of pure oxygen, and the outlet flow (10) during this stage is zero in this example (if the oxidation gas (9) is air or enriched air, it will be composed mainly of nitrogen. The use of pure O (or an O / CO mixture) avoids the loss of CO due to the calcination of CaCO during this sub-step j). Once Cu has been completely oxidized with oxygen (9), the next step j) completes the chemical looping combustion of the BFG fuel gas stream (8). This consists of reducing the CuO solids with CO and H contained in the BFG. In this case, the operation is carried out with a flow rate of 0.52 kg / s of BFG (i.e., 1.35 MW). LHV(heat input). During this step, a 0.59 kg / s flow consisting of 47% v CO2, 48% v N2, and 5% v H2O exits at 910 °C. During the reduction of CuO, the reaction front moves forward rapidly, and the heat released during CuO reduction is used to raise the temperature of the solids remaining behind the reduction reaction front. Under these operating conditions, the temperature rise in the reactor is calculated to be 8 °C. When CuO is completely reduced to Cu, the packed bed is at a uniform temperature of 918 °C. For simplicity, the inlet BFG (8) is assumed to be at the same temperature as the solids at the end of step jj). Feeding a lower temperature gas reactant into the packed bed reactor creates a heat transfer front that grows from the gas reactant entry point and creates a cooler region at the temperature of the gas reactant. Those skilled in the art of chemical looping combustion in packed beds can provide solutions to manage the appearance of such cooler regions, for example, by operating the packed bed with countercurrent flow of the inlet gas, by using a regenerative heat exchanger to preheat the feed gas, or by loading the packed bed with a higher concentration of metal Me at the gas inlet to compensate for the cooling effect of the cooler feed gas with additional heat (released in the combustion of the fuel gas in step i).

[0082] According to the method of the present invention, the next step involves step ii) of extracting pure CO (7) and steam (100) from the previously heated CaCO-containing solids using a CO partial pressure swing by feeding steam (100) to the bed of solids by opening valve (101). This is possible because the packed bed of solids preheated in step i) contains sufficient heat sensitivity at the start of this step ii) to achieve the calcination of 7.5% w of CaCO (corresponding to 4.2% w of active CaO fully carbonated to CaCO), and the bed is at 770°C, the assumed temperature for starting a new CASOH step in the next cycle. The time given for such a CO partial pressure swing step is assumed to allow the calcination of CaCO resulting in the production of a steady flow of pure CO of 1 kg / s during this step iii). The consumption of steam (100) in step ii) is estimated to be an average flow rate of 0.72 kg / s for the duration of step ii).

[0083] It should be noted that due to the exothermic nature of the WGS reaction of CO contained in the BFG in step iii), the heat stored in the bed of solids at the end of steps iii) and i) is greater than the heat required for the calcination of CaCO in step ii). In this particular example, this can be seen by the slight difference between the carbonate (4.1% w) formed at the end of CASOH stage iii) and the 4.2% w of CaCO calcined at the end of step ii). While this difference is small, it accumulates with the number of cycles. Therefore, for the previous temperature to be cyclically repeatable and steady-state, it must be assumed that there is a sufficient reservoir of CaCO available in the bed during calcination, for example, by starting the cycle with a bed that has excess limestone. If this is not the case and the CaCO in the bed is only formed by CaO carbonation, an additional step involving moderate cooling of the bed of solids must be performed before the CASOH step to extract the heat released by the water-gas shift reaction of CO from the packed bed of solids. There are several means for achieving such moderate cooling of the bed. In addition to standard heat removal techniques (i.e., by heat exchangers embedded in a bed of solids), endothermic reactions such as steam reforming of small streams of hydrocarbons such as CH4 can be used, with one of the gas reactants being segregated along the length of the reactor to ensure that cooling is evenly distributed along the length of the reactor.

[0084] From the above mass and energy balance, it is important to emphasize that application of the method of the present invention to calcining CaCO produced during the CASOH process nearly doubles the efficiency of generating H-enriched gas from BFG compared to prior art designs for the CASOH process. In this example, approximately 82% of the BFG used in the process (as gas (13) in step iii) and fuel (8) in step i) is converted to H-enriched gas (15), while only the remaining 18% of the BFG is used to provide energy for the calcination of CaCO during step i). This is a substantial improvement over the configuration proposed by the state-of-the-art presented by Fernandez et al. (2020), in which only 30% of the BFG could be used to generate H-enriched gas from the CASOH process, while the energy contained in the remaining 70% of the BFG had to be extracted as high-temperature heat. Furthermore, the CO2 purity of the 1 kg / s gas stream during step ii) is close to 100% CO2 after condensation of the vapor (100), whereas the CO2-enriched stream at the CASOH state of the art is only about 55-60% v, still requiring separation from the N2 contained in the BFG.

[0085] It should be noted that the application of the method for capturing CO2 to other fuel gases containing CO2 or to combustion flue gases by carbonation of CaO in a packed bed does not differ in essential features from the case described in the previous paragraph.

Claims

1. Steam (100) to CO ２ After separating CO ２ (7) A process for recovering CO ２ is CaCO ３ The solid is preheated to a temperature of 650-800°C and placed in an insulated container (1), the process comprising at least the following series of successive steps: i) burning fuel (8) in said bed of solids to produce CaCO ３ a heating step of heating the bed of solids containing ii) CO by feeding steam (100) to the bed of said solids. ２ Partial pressure swing process, A process characterized by at least two successive repetitions of

2. The combustion of the fuel (8) is carried out in air, oxygen-enriched air or O ２ -CO ２ 2. The process according to claim 1, carried out using a combustion-sustaining mixture (9).

3. 3. The process according to claim 1 or 2, wherein the fuel (8) is a hydrocarbon gas or hydrogen.

4. CaCO3 was added so that the solid was replaced in each combustion step while extracting a batch of particles containing CaO(5). ３ 4. The process according to any one of claims 1 to 3, further comprising the step of introducing a batch of solid particles containing (4) into the insulated vessel (1).

5. The combustion carried out in the heating step i) divides the step i) into two sub-steps: j) the oxidation by a combustion-sustaining gas (9) of a second solid (12) containing a metal such as Ni or Cu, said second solid being said CaCO ３ a first sub-step in which the solids are mixed into a bed of solids containing jj) concentrated CO ２ a second sub-step comprising reducing the oxidized form of the metal fuel formed in sub-step j) by supplying a fuel gas (8) to regenerate the metal fuel in the second solid (12) while releasing (11); 5. The process according to claim 1, wherein the gaseous fuel (8) and the combustion-sustaining gas (9) are chemically looped by dividing the gaseous fuel (8) into

6. CaCO between 0.03 and 0.1 CaO ３ and the CaCO3 present in the solid bed (1) at the start of step i). ３ (4) To regenerate the combustion flue gas or CO ２ 6. The process of claim 1, further comprising an additional carbonation step iii) of the CaO(5) produced in step ii) by feeding any other gas containing

7. At the start of step i) CaCO between 0.03 and 0.1 ３ and the CaCO in the packed bed of solids (1) ３ 6. The process of any one of claims 1 to 5, further comprising an additional carbonation step iii) of the CaO (5) produced in step ii) by feeding the fuel gas (13) and steam (14) to a packed bed of solids containing CaO and Ni to catalyze the steam reforming of fuel gas (13) to hydrogen (15) while reviving (4).

8. A gas containing CO (13) and water vapor (14) is fed to the packed bed of solids containing CaO and Cu to obtain a CaCO of between 0.03 and 0.1 at the start of step i). ３ and increasing the molar conversion of the CaCO3 in the solid bed (1) to ３ (4) is revived, while the hydrogen (15) of the CO contained in (13) and CO ２ 6. The process of claim 1, further comprising an additional carbonation step iii) of the CaO(5) produced in step ii) by catalyzing the water gas shift reaction to

9. 9. The process of claim 8, further comprising the additional step of cooling the packed bed of solids containing CaO and Cu to remove the heat released to the solids by the water-gas shift reaction of the CO contained in the feed of carbon-containing gas (13) in step iii).

10. The thermal energy of the combustion during step i) is ２ 10. The process of claim 1, wherein the average energy input to the equivalent kiln without recovery is the result of multiplying the average energy input to the equivalent kiln without recovery by the ratio of the sum of the durations of steps i) and ii) divided by the duration of step i).

11. The heating step lasts for 5 to 20 minutes, and / or CO ２ 11. The process of any one of claims 1 to 10, wherein the step of partial pressure swing lasts from 5 to 20 minutes.

12. the molar flow rate of steam supplied during step ii) is the total flow rate of gas targeted at the gas outlet during step ii), both divided by the total pressure during step ii) and the CO over CaO ２ 12. The process of claim 1, wherein the partial pressure of the sulphur dioxide is controlled to be proportional to the difference between the equilibrium partial pressure of the sulphur dioxide and the equilibrium partial pressure of the sulphur dioxide.

13. 10. The process of claim 1 is carried out to produce CaCO in the form of limestone. ３ (4) is calcined and CO ２ A shaft kiln for producing a mixture of (7) and steam (100) and CaO (5), the shaft kiln comprising: - a first vertical shaft; - comprising means adapted to effect said combustion of fuel (8) with preheated air or other combustion-sustaining mixture (9) in its calcination zone, The kiln is at least one inlet pipe with a first switching valve (101) configured to supply steam (100) to the lower part of the calcination zone; ２ (7) and at least one outlet pipe with a second switching valve (201) configured to allow extraction of said product gas including steam (100), both the inlet and outlet pipes configured to allow continuous repetition of at least a series of first and second operating periods; means configured to feed batches of limestone through a third solid valve and extract lime (5) at a fourth solid valve during said first operating period, lasting between 5 and 20 minutes, wherein fuel (8) is combusted in said calcination zone in a running mode, said air or other combustion-sustaining mixture (9) being preheated by said CaO (5), and said limestone (4) being preheated by said combustion flue gases flowing from said calcination zone; and a gas inlet configured to supply steam (100) to the bottom of the calcination zone and pure CO ２ means configured to seal said gas inlet / outlet and solids valve during said second operating period, lasting from 5 to 20 minutes, except for (7) and a gas outlet configured to extract product gas containing steam (100); The shaft kiln further comprises:

14. 14. The shaft kiln of claim 13, further comprising a lime cooling zone and a second vertical shaft connected to the first vertical shaft, the second vertical shaft including an additional air flow entering the second vertical shaft and configured to transfer heat from the lime cooling zone and the hot flue gases to the limestone preheating zone.

15. 15. The shaft kiln according to claim 13 or 14, wherein the calcination section of either the first vertical shaft or the second vertical shaft comprises an insulated vessel (1), the insulated vessel (1) further comprising a limestone preheating zone at the top of the kiln (30) and a CaO cooling section (31) configured to preheat the combustion air or other combustion-sustaining mixture (9).

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