Patent Application
20240278176
This invention discloses a calcination process to produce high purity CO2 from solids containing CaCO3. The process operates cyclically and continuously on the solids, arranged in a packed or a moving bed. The process is characterized by each cycle comprising a first step of combustion at atmospheric pressure of a fuel in the bed of solids containing CaCO3 heating them up to 800-900° C., and a second step of application of a vacuum pressure between 0.05 and 0.5 atm to extract pure CO2 from the solids containing CaCO3 while cooling them by 30-200° C. Said combustion can be carried out directly with air, oxygen enriched air or O2/CO2 mixtures when the process is applied to the calcination of a continuous flow of limestone in a moving bed shaft kiln. The process is also applied to calcine CaCO3 formed in reversible calcium looping processes comprising a carbonation reaction step to form CaCO3 from CaO. In these cases, the preferred method of combustion of the fuel used to heating up the bed of solids containing CaCO3 is a chemical looping process, that uses as oxygen carrier a second solid containing a metal oxide (preferably NiO or CuO) mixed in the bed of solids containing CaCO3. In preferred embodiments, the reduced metal (Ni or Cu) resulting at the end of the combustion step is the catalyst used in a sorption enhanced reforming step or in a calcium-enhanced water gas shift reaction step where the bed of solids containing CaCO3 is regenerated.
This is a process to capture CO2 in pure form during the calcination of CaCO3. When the source of CaCO3 is a natural limestone and the product is CaO, the process is applicable to shaft kilns. When the source of CaCO3 is a process involving a step of carbonation of CaO, the process is applicable to calcium looping processes to capture CO2 from flue gases, or to sorption enhanced reforming or sorption enhanced water gas shift processes where carbonaceous fuels are converted into hydrogen, while CaCO3 is formed from CaO.
The capture of CO2 in high purity form (for example with a volume fraction higher than 99% v in CO2) from fuel and flue gases in industrial processes, is of interest in the context of climate change mitigation efforts. Relevant for the purpose of this invention is a family of commercial processes that capture CO2 by separating CO2 from other gases using vacuum pressure swing adsorption (VPSA). VPSA processes operate cyclically and continuously. In each cycle, a gas containing CO2 is arranged to contact a packed or moving bed of solids with affinity and selectivity to adsorb CO2. VPSA CO2 capture systems usually work at relatively mild temperatures (<200° C.), using CO2 sorbent materials that establish weak bonds with the CO2 so as to minimise energy requirements during regeneration. When the bed of solid is saturated in CO2, a system of switching valves seals the bed of solids and a vacuum pump is connected to the vessel containing the bed of solids to force the desorption of the CO2 from the solids under pressures as low as 0.05 atm. A series of vessels containing solid sorbents, operating in parallel, allows the operation in continuous mode. Intermediate gas purge cycles, pressure balancing between beds and well-established switching control technologies allow for VPSA CO2 capture system to be among the most competitive and least energy consuming CO2 capture systems.
On the other hand, there are high temperature calcium looping CO2 capture systems, using CaO as sorbent of CO2 (usually at temperatures between 600-700° C.) to form CaCO3. The CO2 can be captured from combustion flue gases or during sorption enhance reforming processes or calcium-enhanced water gas shift processes, to produce hydrogen from a fuel gas while capturing CO2 at high temperature with CaO.
The operation at these high temperatures facilitates the recovery, during the exothermic carbonation of CaO (−170 kJ/molCaO at 650° C.), of the equivalent energy consumed during the calcination of CaCO3, that takes place at temperatures usually higher than 900° C. (depending on ambient concentration of CO2).
Some of the state of the art of calcium looping processes is making use of vacuum swing principles. US2017/0096335A1 claims higher energy efficiencies by operating an adiabatic calcium looping packed bed reactor at a sufficient high pressure during the exothermic carbonation step to reach sufficiently high temperatures of the carbonated solids as to initiate a reversible vacuum-driven calcination step after each carbonation step. CaO solids heat up by up to 150° C. at the end of the exothermic carbonation period. Such increase in temperature is claimed to be sufficient to sustain a subsequent vacuum calcination step, that causes an equivalent drop in temperature as the CaCO3 is calcined. Note that no heat is added externally to the packed bed of sorbent, as the system is assumed to be adiabatic. This means that the sorbent heats up during carbonation and cools back to its original temperature during calcination indefinitely during successive carbonation/calcination cycles. Such process will however face practical limitations, that may make it not viable in practice. On one hand, real systems are not fully adiabatic. Heat transfer from the outside, or from gas reactants supplied at higher temperature than the gas products, is needed to fulfil the overall heat balance. However, there are constrains imposed to the temperatures of the gas reactants from the additional heat transfer equipment required and because in some cases, preheating such gas reactants is not viable. For example, in sorption enhanced water gas shift processes using CaO as CO2 sorbent, the syngas (containing CO) must be fed at temperatures below 500° C. to avoid carbon deposition. Linked to this, metal dusting phenomena in the heat exchangers used as gas preheaters must be prevented by feeding gases at temperatures as low as 300° C. However, the product gases tend to leave the reactor at higher temperature (i.e., between 650-850° C. depending on the temperature achieved during CaO carbonation). Therefore, such heat imbalance in the gases flowing in and out of the reactor must be compensated by an external heat source. Similarly, sorption enhanced reforming systems using CaO to produce H2 from a fuel gas have even larger thermal energy constrains because reforming reactions are highly endothermic, and therefore, there is no sufficient enthalpy in the overall carbonation-reforming step to sustain the subsequent calcination step. A substantial external heat supply is therefore needed to sustain any of the referred calcium looping systems. The energy released during carbonation is simply not sufficient to drive the reverse calcination reaction in real calcium looping systems.
The problems associated with the need of external heat supply to sustain the operation of calcium looping systems are aggravated when using packed beds and when taking into account the typically low carbonation molar conversion of real CaO materials when reacting with CO2 during many cycles. Despite intense research effort over the last decades, there is no yet a suitable and widely available CaO sorbent (i.e., with proven mechanical stability and reactivity to react fast and reversibly with CO2 under realistic conditions in the reactors) able to sustain more than 3-10% molar carbonation conversions of CaO to CaCO3 over hundreds or thousands of cycles, as it is the case for CaO derived from calcination from natural limestone. This means that active CaO in calcium looping processes using currently known CaO materials, will typically be accompanied by a very large portion (95-99% w when including other solids in the system such as catalyst and their supports) of materials that are inert towards carbonation. Therefore, despite the exothermic nature of carbonation of CaO, the temperature increment achievable during carbonation will not reach more than 50-150° C. (depending on carbonation conversion and fraction of inerts), which is unsuitable to connect the operating window for effective CO2 capture by carbonation (between 600-700° C.) with the operating window for effective calcination of CaCO3 in pure CO2 (the decomposition temperature of CaCO3 at atmospheric pressure in pure CO2 is about 900° C.). Therefore, additional heat requirements during calcination are known to be needed when operating with known CaO sorbent properties (i.e. a maximum CO2 carrying capacity of 0.05 and 0.1) and known operating windows for carbonation and calcination. Moreover, US2017/0096335A1 is silent about the need of external heat supply to their calcium looping system.
U.S. Pat. No. 9,505,998B2 discloses a solution to solve the problem of heat supply to a packed bed calciner by arranging a combustion chamber for providing additional thermal power to the calciner through a metallic wall, that separates the region containing the bed of solids with CaCO3 from the combustion chamber. However, this solution is limited in applicability, because it requires large heat transfer areas of costly alloys to allow the necessary heat transfer between the combustor chamber and the calciner at temperatures exceeding 900° C., in particular when considering that combustor chamber and calciner may need to operate under substantial differences in pressure.
U.S. Pat. No. 8,506,915B2 discloses another solution to the problem of heat requirements for calcination within a bed of solids containing CaCO3 by introducing a second solid in the bed that contains Cu/CuO. This is a type of oxygen carrier material that has been developed and tested in the state of the art of chemical looping combustion or indirect combustion applications. In chemical looping combustion processes, a metal (such as Ni, Cu, Fe) is oxidised at a high temperature in an “air reactor”, in a process that is highly exothermic. In a subsequent step, the metal oxide or “oxygen carrier” (such as NiO, CuO or Fe2O3) is reduced by a fuel gas to produce a N2-free flue gas and regenerate the metal or “metal fuel”. However, the use of metal fuel oxidation reactions with air to supply energy for CaCO3 calcination has the problem of the unwanted dilution of the CO2 evolved during calcination with the nitrogen contained in the air. To solve this, the key feature of U.S. Pat. No. 8,506,915B2 is the choice of the Cu/CuO chemical loop, as CuO can undertake an exothermic reduction reaction to Cu using a fuel gas (that burns primarily to CO2 and H2O) and supplies the necessary energy to the solids bed to drive the calcination of CaCO3 during the reduction stage of CuO to Cu. This process has been proposed for Sorption Enhanced Reforming of methane and steam to produce a hydrogen-rich stream and high purity CO2, using a bed of solids containing a Nickel reforming catalyst as a third additional solid to the CaO sorbent and Cu oxygen carrier ((Fernandez et al. Conceptual design of a hydrogen production process from natural gas with CO2 capture using a Ca—Cu chemical loop, International Journal of Greenhouse Gas Control 2012, |Volume 6|Page 126-141). Another variant of such process has been recently been proposed to treat steel mill off gases, like blast furnace gas, BFG, in the so called CASOH process, where a bed of solids containing CaO first removes the CO2 contained in the BFG and promotes the enhanced water gas shift of the CO by capturing the resulting CO2 as CaCO3. In a subsequent step, the reduction of CuO with a fuel gas such as CH4 can regenerate CaO while producing CO2 and water vapour (Fernandez et al. Advanced Packed-Bed Ca—Cu Looping Process for the CO2 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 reveals that they can be highly energy efficient if, and only if, an effective heat recovery for power generation is included in the system. This is because these processes require to function with an intermediate additional reaction step involving the oxidation of the Cu with air, before the desired reduction-calcination step. Such oxidation introduces demands for a technically challenging operation at high pressure, needed to minimise the calcination of CaCO3 during the oxidation of Cu. The need of additional power generation equipment is a weakness in practical applications targeting a maximum production yield of H2 from sorption enhanced reforming or calcium-enhanced water gas shift processes, where it is technically more attractive to maximise the H2 yield (because this is the product targeted from such processes) than maximise the thermal energy outputs from the process (even if they are in the form of high temperature gas product streams).
Also related to this invention are patents related to the application of vacuum pressure to a calcination process of limestone, so that the CO2 evolved from the calcination of CaCO3 can be obtained separately or even in pure form.
U.S. Pat. No. 4,748,010 reviews in detail prior state of the art on vacuum calcination of batches of limestone to produce lime. It discloses a semi-continuous CaCO3 calcination process that attempts to exploit the favourable calcination conditions achieved when a vacuum is applied to a preheated stone containing CaO and CaCO3. In a first step, the limestone is heated and partially calcined by air combustion until 50-85% calcination conversion is achieved (i.e., % of mol of CaCO3 converted to CaO) obtaining a hot mass with sufficient accumulated heat to complete the calcination of the remaining unconverted limestone therein in one final cycle. To do this, the hot mass of solids are moved to a vacuum chamber zone, where the complete conversion of the limestone is achieved using the retained heat in the solids, while subjecting the partially calcined hot mass to a sufficiently strong vacuum during the time needed to achieve total conversion of the limestone. It must be noted that U.S. Pat. No. 4,748,010 is not intended to be a CO2 capture process, because most of the CO2 evolved during calcination is emitted in diluted form in flue gases emitted during the first heating up step and partial calcination of the limestone, under standard combustion conditions in the kiln. Furthermore, it is important to note that the fraction of pure CO2 attainable by the method disclosed in U.S. Pat. No. 4,748,010 will not exceed in reality more than 10% molar conversion of the initial mass of limestone, due to known thermodynamic and kinetic constrains, that are briefly explained below.
Hills, AWD; “The mechanism of the thermal decomposition of calcium carbonate” Chemical Engineering Science, 1968, Vol. 23, pp. 297-320, investigated how the decrease in the partial pressure of CO2 around a calcining stone (or a stone containing a core of CaCO3) results in an initially fast calcination rate and a decrease in the temperature of the carbonated core. For example, FIG. 4 of Hills shows a decrease in temperature in the interior of a calcining particle of 1 cm diameter once it is subjected to a step change in the partial pressure of CO2 in the surroundings of the particle held at atmospheric pressure. The decrease in temperature recorded in the interior is a sign of a fast heat transfer from the interior towards the external CaCO3 stone surface, partially driving the calcination. After such initial drop in temperature (about 50° C. in less than 5 minutes) the internal parts of the stone must reach a temperature of the exterior and slow heat transfer from the external CaO part of the partially calcined stones towards the shrinking carbonated core becomes the governing process in the overall calcination rate of the stone. Note that the complete calcination is achieved after 70 minutes in the example of such
To overcome these technical limitations, means for heat transfer is clearly needed towards the calcining solids located in a packed or moving bed of solids. As discussed above for calcium looping systems, there are methods proposed that involve indirect heat exchange to the vacuum calcination chamber through a metallic wall (see for example U.S. Pat. No. 2,080,981). However, such indirect calcination processes are likely to require very large and costly heat transfer areas of high temperature alloys, to withstand the high temperatures characteristic for the calcination of CaCO3 and the differences in pressure between both sides of the heat transfer wall.
On the other hand, the state-of-the-art solution to produce concentrated CO2 from the decomposition of CaCO3 is to burn a fuel in the vicinity of the calcining solids, using comburent mixtures of air enriched in O2 or even N2-free CO2/O2 mixtures. Such oxy-combustion CaCO3 calcination systems have been developed for a range of cement manufacturing processes, lime kilns and calcium looping processes (including postcombustion CO2 capture processes and sorption enhanced reforming and water gas shift processes) involving fluidised bed or entrained bed technologies, that facilitate heat transfer processes at particle level toward the calcining particles. Among other disadvantages for their application to packed bed of solids, all oxycombustion methods share the disadvantage of the need of additional equipment to separate O2 from air and the additional electricity requirements to drive the air separation unit (or oxygen enrichment plant). They also need downstream purification devices for CO2, that are also energy consumers, as the flue gases from oxycombustion contain variable fractions of O2, water vapour and inert gases (from air ingress) that need to be removed before transporting the CO2 for use or storage.
Finally, also relevant for this invention is a family of industrial processes where the calcination of limestone minerals containing CaCO3 is carried out to produce calcined materials including CaO, emitting diluted CO2 in the combustion flue gases from the calcination of CaCO3. Modern commercial designs of lime shaft kilns (i.e., countercurrent shaft kilns, parallel flow regenerative kiln, annular shaft kilns etc.), have high energy efficiencies, very high conversions of limestone to lime and a suitable quality of the lime product (“Lime Shaft Kilns”, Hannes Piringer, Energy Procedia 120 (2017) 75-95). About ¾th of the CO2 emitted by the kiln arises from the decomposition of CaCO3 (the rest comes from the combustion of a carbonaceous fuel in air). This means that, unlike other systems relying on combustion, these air-fired kilns cannot be decarbonized today by using renewable fuels such as hydrogen or metal fuels (i.e., Fe, Ni, Cu materials exothermically oxidized by air), because a solution would still be needed to capture the CO2 evolved from calcination. Shaft kilns operate by tuning the combustion conditions in a quasi-adiabatic calcination or burning zone typically located in a central part of the kiln. With such arrangement, hot flue gases evolved from the calcination zone preheat the limestone fed at the top of the kiln and the hot lime produced in the calcination zone is used to preheat the air coming in the kiln to sustain combustion. In parallel flow regenerative shaft kilns (containing two shafts), an inherent part of the process consists on operating the kiln in a cyclic way, by applying a temperature swing to the solids alternating combustion periods (of about 10-20 min long) in each shaft and no-combustion conditions (during which the shaft behaves as a regenerative gas-solid heat exchanger). To achieve continuous operation, these processes require the two parallel shafts operating at opposite modes at any point in time. Very long residence time of the solids are required in these kilns (between 10 and 20 hours, depending on the stone size and on other characteristics affecting the heat transfer to the stone). In contrast, the flue gas resides in the kiln less than 20 seconds, with gas velocities maintained between 1-2 m/s to moderate pressure drop and solids entrainment. The exterior surface of the stone may have temperatures from 900-1800° C. depending on fuel firing conditions. Staging of fuel feeding points and low air-excess are known to allow larger flames and milder local flame temperatures in the vicinity of the stones. Largely unaffected by the temperature in the exterior the stone (Furnas, The Rate of Calcination of Limestone; Industrial and Engineering Chemistry, 1931, 23(5), 534-538), the core of unconverted CaCO3 remains at temperatures quite constant under or around 900° C. (i.e., close to the equilibrium temperature of CO2 on CaO at atmospheric pressure). In contrast, in stones that have not yet reached the calcination zone (i.e., those in the preheating zone of the kiln) the heat transfer from the flue gas to the CaCO3 rich stone allows more uniform temperature profiles between the core and the external surface of the solids and the difference in temperature between the core and the gas may be less than 50° C.
From the previous review of the state of the art, it can be concluded that there is still a need for technically viable, energy and product efficient process to obtain pure CO2 from the decomposition of the CaCO3 contained in the limestone fed to lime kilns or from the CaCO3 formed in calcium looping processes carried out in packed or moving beds. The process of the present invention maintains or exceeds the high energy efficiency and product qualities of non-CO2 capture processes involving CaCO3 calcination, such as limestone calcination processes in lime shaft kilns, sorption enhanced calcium looping reforming processes and calcium-enhanced water gas shift processes for hydrogen production (where CaO acts as a CO2 sorbent).
The main object of the present invention is to provide a process to recover in high purity form part of the CO2 released from the calcination of solids containing CaCO3. The process is applicable to processes for calcination of limestone in shaft kilns, which are thermally insulated vessels that have moving bed of solids containing CaCO3 in the calcination region of the shaft kiln. The method is also applicable to calcium looping processes using thermally insulated vessels containing packed beds of solids containing CaCO3, in this case formed during the capture of CO2 by CaO in contact with a combustion flue gas or any other gas containing CO2, or during the sorption enhanced reforming or sorption enhanced water gas shift process steps, that convert fuel gases to hydrogen when in the presence CaO and other gases, forming CaCO3.
All said processes comprise of a number of known steps to achieve the desired products without capturing CO2 in separate form, or require additional inefficient process steps to capture such CO2 as reviewed in the state of the art. Therefore, they will all benefit from a process to recover in pure form the CO2 released from the calcination of solids containing CaCO3 arranged in a thermally insulated vessel. In the framework of such existing processes, we assume that solids are preheated to a temperature between 650-800° C. and are arranged in a thermally insulated vessel able to withstand vacuum swings. This invention provides a calcination process characterized by the continuous repetition, at least twice, of a sequence of at least the following consecutive steps:
In this way, the process gives as reaction products CO2 and CaO, and causes a temperature drop in the bed of solids of between 30-200° C. Also, there is an increase in molar conversion of CaCO3 to CaO of 0.01 to 0.07.
Preferably, the combustion of the fuel is carried out with air, oxygen-enriched air or O2—CO2 comburent mixtures.
The intention with the previous sequence of steps is to exploit the large difference between heating times and calcination times of the typically large particle sizes (10-150 mm) of limestones processed in shaft kiln calciners or in the particles or pellets (2-10 mm) used in calcium looping processes in packed beds. As quantified below in several examples and experiments, the higher heating rate of the particles in step i) respect to their calcination rate in the same step i), allows for the introduction to the bed of particles of a certain amount of thermal energy to heat them up, while limiting the progress of the calcination of CaCO3. It is however unavoidable at this stage to initiate a certain calcination of the CaCO3, and loose some of the CO2 evolved in such calcination as diluted gas in the flue gases generated by the combustion of a fuel in the bed. At the end of step i) the solids have heated up to an average temperature of 800-900° C. from a lower initial temperature (30-200° C. lower) that resulted from the application of vacuum in the previous step ii). The use as comburent of oxygen enriched air, or mixtures of O2 and CO2, instead of air during the combustion of the fuel i), will have a beneficial effect in minimising the decomposition of CaCO3 during the combustion step i). This is because combustion with higher O2 content allows higher concentrations of CO2 in the flue gas, known to reduce the rate of calcination. For example, when the CO2 content of a flue gas enriched in CO2 is 50% v at atmospheric pressure the calcination rate of CaCO3 is extremely slow compared to the heating rate of the stone until the stone reaches a close value to 850° C. (see other examples below). This is because the equilibrium partial pressure of CO2 on CaO at 850° C. is close to 0.5 atm. The use of said comburents with high O2 content during step i), also facilitates the introduction of the necessary energy for calcination in a shorter times of the step i). This is beneficial because it allows the introduction of the vacuum step ii) in the process without altering the overall/average thermal power input to the kiln over long periods of time. In other words, introducing step ii) without increasing the power input during the i) would necessarily translate into a reduction of the average production rate of CaO. This will impact directly the production rate of lime in a shaft kiln or the rate or capacity of the Calcium Looping capture system using such CaO to capture CO2.
Therefore a preferred embodiment of this invention further comprises the combustion of the fuel (8) carried out with oxygen-enriched air or O2—CO2 comburent mixtures (9). As will be illustrated in one example below for a lime kiln, adopting a time for step i) identical that for step ii) requires the power input during step i) to double respect to the state of the art process without capture, in order to maintain the same overall lime production capacity. Therefore, the use of a comburent mixture of about 40% vol in O2, which is part of the state of the art of oxy-combustion systems, will facilitate this doubling of power input while keeping similar flowrates of gases in the kiln, that is important to be able to maintain heat transfer characteristics, pressure drop in the bed, energy efficiencies etc. respect to the those in modern kilns with no CO2 capture.
When the stone or carbonated sorbent has been heated to a temperature between 800-900° C. at the end of step i), the onset of the vacuum step ii) promotes a fast partial calcination of the CaCO3 containing solids, that is associated with an equally fast drop in temperature in said solids. In an industrial scale packed bed, with many stones or carbonated sorbent calcining under vacuum at the same time, the release of CO2 will be controlled at this point by the maximum capacity of the vacuum pump to handle the CO2 flow. However, the rate of calcination will eventually decrease with time, as calcination progresses, the availability of heat in the interior of the stone diminishes, and temperature profiles develop in the stone or sorbent. The temperature in the interior of the stone would decreases to values where the intrinsic calcination rate is too slow at the conditions of temperature and partial pressure of CO2 in the calcination reaction front (which can be assumed to be at equilibrium conditions for temperatures over 700° C.). In view of experimental results shown below, after 5-20 minutes of vacuum a drop in average temperature in the bed of solids (between 30-200° C. is caused. This is associated under adiabatic conditions to an increase in molar conversion of CaCO3 to CaO, or a decrease in CaCO3 molar fraction, of 0.01 to 0.07. At this point, a return path must be initiated to re-heat the bed of solids. This is the purpose of step i), which involves a combustion step taking place within the bed of solids, to facilitate a fast heat transfer of the combustion energy to the solids, that have just been cooled during the vacuum step ii).
When the thermally insulated vessel is a lime shaft kiln, the cyclic sequence noted above is repeated until each stone of limestone initially fed to the kiln is fully calcined in the shaft kiln, following the teachings of the state of the art of lime shaft kilns. Following such teachings, such kiln can be operated in continuous mode in a moving bed configuration, comprising a step of introducing to the top of the kiln a batch of particles of solids containing CaCO3 to the thermally insulated vessel so that the solids are displaced downwards in each combustion step while extracting a batch of particles containing CaO from the bottom of the kiln.
In a preferred embodiment, the thermally insulated vessel of the process of this invention is the calcination section of a shaft kiln that additionally contains a limestone preheating zone at the top of the kiln and a CaO cooling section that also preheats combustion air (or oxygen enriched air or O2/CO2 mixture, if such comburent is preferred to sustain higher thermal power and higher partial pressures of CO2 during the combustion or heating up in step i).
The invention also refers to a shaft kiln suitable to carry out the calcination process noted above. The shaft kiln having means to effectively burn a fuel with preheated air, enriched air or O2/CO2 in its calcination zone, said kiln having walls thermally insulated and capable to operate under vacuum pressure and characterized by having inlet and outlet gas and solid pipes containing switching valves with at least one of the pipes having a vacuum switching valve connected to a vacuum pump to alternate in a cyclic manner between these operating periods:
A particular variant when the comburent mixture is O2/CO2 involves the staging of separate streams of O2 (from an air separation unit) and CO2 (recycled) to prevent carbonation of the CaO exiting the kiln.
If a double shaft kiln is used, this will operate step i) with combustion in just one shaft using part of the air preheated with the bed of limestone preheating region, while arranging an additional flow of air to enter the second shaft, in order to transport heat from the lime cooling region and the hot flue gases to the limestone preheating region and use such heat to preheat part of the air entering the kiln from the top.
Following the teachings of the state of the art of gas burners in kilns, air and fuel distribution in the calcination region of the kiln (1) during step i) can be arranged in many forms to accommodate maximum allowed temperatures in the kiln as required by CaO quality requirements. However, it is important to realise that, since kilns are usually firing a fuel continuously, the introduction of the vacuum step ii) in the methods of this invention to operate lime kilns, necessarily represent a drop in lime production capacity proportional to the ratio between the duration of time ii) and the total duration of one cycle of combustion and vacuum steps (t1+t2). Therefore, a preferred embodiment to solve this is to increase the thermal power input during step i) and so that the thermal power of the combustion during step i) is the result of the multiplication of the average power input in the equivalent kiln without CO2 capture and the ratio between the sum of the duration step i) and ii) divided by the duration of step i).
Lime kilns can also be operated with solid fuels in step i), by mixing the right proportion of a solid fuel with the limestone feed. In order to minimise emissions of CO2 and other gas contaminants during the step i) involving combustion of the solid fuel, it is a preferred embodiment to use as solid fuel a metal such as Fe, Ni or Cu, that exothermically oxidizes with air and that is fed contained in a second solid that is mixed with the solids containing CaCO3. The intense premixing of the solid fuel with the limestone will lead to a more controlled temperature profiles and homogeneous combustion in the calcination zone of the kiln. As will be shown in examples below, it is possible to maximise the heating of particles containing CaCO3, while minimising the calcination of the CaCO3 during combustion step using a metal fuel, because the energy for the heating of the solids (by 30-200° C. as mentioned above) is a small fraction of the energy required for their calcination. Therefore the quantity of air needed for metal fuel oxidation is small, and the maximum losses of CO2 in such small air mass (given by the equilibrium of CO2 on CaO at the maximum heating temperature) will also be small.
When the calcination process of this invention is applied to calcium looping processes, the solids containing CaCO3 come from a previous carbonation reaction step that is characteristic of all calcium looping processes, which involves the carbonation of CaO to CaCO3. There is no longer need of a continuous feeding from the outside of CaCO3, because the calcium looping process is reversible: the CaCO3 formed during carbonation step is mainly the CaCO3 that will be calcined during the calcination stage (note that there may be also small fraction of CaCO3 coming from purges or initial batches of CaCO3 or CaO). In these calcium looping systems, due to the limited value o reversible carbonation molar conversions of the known CaO sorbents (0.03 to 0.1), and the dilution of the Ca-containing solids with catalyst and their supports, it is not favourable to arrange a direct combustion of a fuel with air within the bed of solids. This is because local flames and hot spots would damage the integrity and activity of catalyst and/or sorbent. Therefore, it is important in these systems to undertake the heating up step i) by indirect combustion of a fuel, following the teachings of chemical looping combustion processes in packed beds.
Therefore, preferred embodiments of this invention refer to the application of the calcination method of this invention to a calcium looping processes that use packed beds, where solids containing CaCO3 are formed from CaO and CO2 during a carbonation step, in the presence of a second solid containing a metal such as Ni or Cu acting as catalyst and/or oxygen carrier, where the combustion taking place in during the heating up step is an indirect chemical looping combustion of the gaseous fuel with air by dividing the fuel combustion period i) in two sub-steps:
Such second sub-step is therefore regenerating the metal catalyst, that must be in reduced form to maintain catalytic activity during the sorption enhanced reactions of natural gas and steam or water gas shift taking place at the same time than the step of carbonation of CaO.
A preferred embodiment the calcination process of this invention is applied to a postcombustion CO2 capture process by calcium looping characterized by an additional carbonation step iii) of the CaO generated in step ii) by feeding a combustion flue gas or any other gas containing CO2 to the packed bed of solids, to the packed bed of solids to increase the molar conversion of CaO to CaCO3 between 0.03 and 0.1 and restore the CaCO3 that is in the bed of solids at the beginning of step i).
In another preferred embodiment the calcination process of this invention is applied to a sorption enhanced reforming process in packed beds, SER. Such process requires a second solid in the bed that contains Ni, that is known to perform as a metal fuel (in step i)) and as reforming catalyst during SER step. Such SER step includes a carbonation step iii) of the CaO generated in step ii) by feeding a fuel gas and steam to the packed bed of solids containing CaO and Ni. The sorption of CO2 by carbonation of CaO enhances the steam reforming of the fuel gas to hydrogen while increasing the molar conversion of CaO to CaCO3 between 0.03 and 0.1 and restore the CaCO3 that is in the packed bed of solids at the beginning of step i).
A favourable SER process variant of this embodiment allows a higher yield of H2, by introducing an additional step after the sorption enhanced process by carbonation iii) and before the beginning of a new cycle N+1 with step i), consisting in the feeding a fuel gas (13) and steam (14) to the packed bed of carbonated solids to produce a syngas that is used as fuel gas (8) in sub-step jj) of the next step i) in the cycle N+1. As will be shown in the example below, this is possible because of the high quantity of thermal energy contained in the bed of carbonated solids at the end of step iii). Such thermal energy available at high temperature can sustain a partial steam reforming of the natural gas, to generate a syngas with higher energy content while decreasing the temperature of the carbonated bed of solids (see Example 3 below)
In another preferred embodiment, the calcination process of this invention is applied to a sorption enhanced water gas shift process carried out in packed beds, the process being characterized by the metal fuel and water gas shift catalyst being Cu and by an additional carbonation step iii) of the CaO generated in step ii) by feeding a gas containing CO and steam to the packed bed of solids containing CaO and Cu to enhance the water gas shift reaction of the CO contained in to hydrogen an CO2 while increasing the molar conversion of CaO to CaCO3 between 0.03 and 0.1 and restoring the CaCO3 that is in the bed of solids at the beginning of step i).
In previous embodiments of the process, it is possible to introduce short additional gas purging/rinsing steps for security and gas purity reasons. One particular additional step valid for all embodiments disclosed is characterised by the introduction an additional gas rejection step by which a gas volume of 1-3 times of the void volume of gas in the packed bed of solids is discharged at the beginning of the vacuum period ii). This action will avoid the mixing with the pure CO2 evolved during the vacuum calcination of CaCO3 of the small mass of non-CO2 gases present in interstices of the bed of solids at the beginning of step ii).
The figures below reflect aspects for the practical execution of the process and facilitate the understanding of the main differences of the invention with respect to the state of the art, but they should not be considered as to limit the scope of the invention. Figures use numbers to name physical elements, devices and materials required to carry out the methods. Thick black arrows represent solid and gas mass streams entering/exiting the devices. The characterising sequences of steps of preferred embodiments are marked in roman numbers i), ii), iii), lasting a time t1 to t3, with the changes between steps marked with void thick arrows. Dashed line boundaries represent full cycles N of the process of this invention, containing several steps and sub-steps per cycle.
Reference is made from this point to the accompanying drawings to define preferred embodiments of the invention and provide experimental evidence to support the claims.
Alternatively, the content of the CaCO3 in the original solid undertaking a partial calcination in cycle N may be the result of a previous carbonation process step (not shown in
Therefore, the process of this invention comprises the partial calcination of CaCO3 containing solids in repetitive cycles, N, characterized by each cycle having a sequence of at least these two steps as represented in the top of
In order to achieve a high recovery efficiency of pure CO2 respect to the total CO2 contained in the solids as CaCO2, it is important to maximise the quantity of CO2 released by calcination during ii) and minimise such quantity during i). To gain insight into dynamic heat transfer and calcination phenomena during step i) and ii) we built an experimental device to illustrate the process of this invention for a single limestone rock and for a small bed of CaCO3 containing solids. In addition, existing models for dynamic heat transfer (solved in COMSOL) and state of the art thermodynamics and kinetics of calcination of CaCO3 have been used to define preferred but not restrictive values for t1 and t2 as a function of the characteristics of the material containing CaCO3 and operating conditions.
The laboratory scale experimental device to test calcination by successive heating-vacuum cycles is a steel vessel of 130 mm id and 200 mm height surrounded by an electric oven to control the temperature in the gas environment of the device. In typical experiments as plotted in
We must highlight at this point that the advantage of a much faster calcination rate when the core of the stone starts at a temperature higher than the minimum given by calcination controlled by external heat transfer, has not yet been exploited in practice. As reviewed in the state of the art, in industrial scale vessels with packed beds of solids containing CaCO3, the limiting factor of calcination is not the transfer of heat to an individual stone from the inside of the stone, but the transfer of heat to the whole packed bed of material in the vessel.
The method disclosed in this invention, involving a continuous repetition of cycles with a heating up step intended to restore the heat in the interior of the CaCO3 before a new vacuum step begins, is intended to exploit such heat transfer advantage clearly observed in
The experiment of
Modelling of dynamic heat transfer phenomena in spherical stones of CaCO3 or CaO also help to support the chosen conditions and times in the two steps that characterise this invention. This is particularly important to elucidate conditions to avoid excessive calcination of CaCO3 during the heating up step i).
The experimental proof of the previous statement is implicit in
The previous experimental results, and the modelling carried out to interpret them, demonstrates that in reasonable heating up times (5-20 minutes) it is possible to transfer the necessary energy (50-150 kJ/kg) to solids containing CaCO3 to bring them back to an average temperature of 800-900° C., necessary to re-initiate a new vacuum cycle step, where temperatures will drop between 30-200° C. (causing a decrease of molar CaCO3 content between 0.01 and 0.07) in the solids. Calcination rates under vacuum step ii) are sufficiently fast to complete an incremental calcination molar conversion of 0.01 and 0.07 in just 5-20 minutes of vacuum stage, at vacuum pressures between 0.05-0.5 atm. The skilled in the art will note that reducing vacuum pressure allows larger drops in temperature during calcination, and a reduced number of cycles of steps i) and ii), at the expense of a higher energy consumption to generate such low vacuum pressure. Moderate vacuum pressure demand for higher temperatures of heating of the carbonated solids and/or higher number of vacuum cycles, as the drop in temperature of the solids will be limited by the equilibrium temperature imposed by the vacuum pressure (for example, for 0.5 atm of vacuum, it will not be possible to lower the temperature below 850° C. during vacuum as this is the equilibrium temperature of CO2 on CaO at such partial pressure of CO2).
Taking the previous constrains into account, a preferred embodiment of the process of this invention allows the calcination in continuous mode of a flow of limestone (4) and the extraction of an equivalent volumetric flow of lime (5). The schematic of each single cycle N of the process is presented in
It must be noted that kilns of
The application of the process of this invention to the calcination of limestone in a shaft kilns is non-restrictive. As known from the state of the art, there are a number of CO2 capture processes by calcium looping requiring calcination of carbonated solids (formed during the carbonation of CaO with CO2 in a flue gas or fuel gas), to generate a pure stream of CO2. These processes can benefit from the sequence of heating-vacuum steps for calcination disclosed in this invention. In particular, there is a family of calcium looping processes that include cyclic operations in adiabatic packed beds (1) of solids (4-5) that do not move in the bed, and that rely on the capture of CO2 by the reversible carbonation of CaO (5) to form solids containing CaCO3 (4), followed by the calcination of the formed solids containing CaCO3 (4) to produce concentrated CO2 (7) and regenerate the CaO sorbent (5). Example of such processes are sorption enhanced steam reforming process of natural gas and other hydrocarbons (including biogas) using CaO as a sorbent to displace reforming reactions towards hydrogen production. Also, sorption enhanced water gas shift using CaO as sorbent to displace the water gas shift equilibrium towards the production of H2.
In order to understand the relevance of this invention for the technical viability of such calcium looping processes, it is important to highlight here these known elements of the state of art:
It should be evident to the skilled in the art that for these realistic materials and conditions it is not possible to established the desired temperature swing between carbonation and calcination using only the energy involved in carbonation and calcination reactions. The lack of sufficient thermal energy for calcination worsens when considering heat losses and the heat balance affecting incoming low temperature gas reactants and high temperature gas products. Therefore, the method of this invention is a solution to close the heat balances and the temperature gaps between carbonation and calcination temperature conditions in calcium looping processes. This will make them technically viable without the need of high performance, high activity Ca-materials, that have not yet become commercially available after decades of worldwide research work.
In view of all experimental results and discussions above, a preferred embodiment of this invention (see
As indicated in
Examples 2 and 3 are provided below by using a preferred embodiment of this invention as represented in
A process variant in previous reforming processes, that translates into higher yields of hydrogen, involves the partial reforming of the natural gas with steam using the sensitive heat contained in the high temperature bed of solids resulting after the carbonation step, as explained in an example below using
For conciseness, detail examples of the application of this invention to other calcium-enhanced reactions for biogas or syngas derived from biomass have also been omitted. Such processes can be considered variants where the origin of the solids containing CaCO3 are the calcium enhanced reactions forming CaCO3. All these processes will benefit from the solution provided in this invention to calcine said CaCO3 while producing pure CO2 and regenerating the CaO for a new cycle of enhanced calcium enhanced reactions.
In all previous embodiments of the process, the skilled in the art will introduce short additional gas purging/rinsing steps for security and gas purity reasons, for example using steam or nitrogen as inert gases.
One particular additional step valid for all embodiments disclosed here is characterised by the introduction of an additional gas rejection step by which a gas volume of about 1-3 times of the void volume of gas in the solids packed bed (1) is discharged at the beginning of the vacuum period ii). This action will avoid the mixing of the small mass of non-CO2 gases present in the interstices of the bed of solids at the beginning of step ii) with the pure CO2 evolved during the vacuum calcination of CaCO3.
Mass and energy balances are solved in the three examples below to illustrate the application of the calcination method disclosed in this invention to calcine solids containing CaCO3 into CaO and a pure stream of CO2 in a shaft lime kiln (Example 1), a calcium looping process for calcium-enhanced water gas shift (CASOH) catalysed by Cu to produce H2 from Blast Furnace Gases and pure CO2 (Example 2), and a calcium looping process for sorption enhanced reforming process catalysed by Ni to produce H2 from natural gas and pure CO2 (Example 3). For simplicity and transparency, the following thermodynamic parameters are assumed to be constant: enthalpy of carbonation and calcination=170 kJ/mol Ca; enthalpy of WGS=−34 kJ/mol CO, enthalpy of steam methane reforming=190 kJ/mol CH4, enthalpy of Cu oxidation=−145 kJ/molCu, enthalpy of CuO reduction=−54.5 kJ/molCu with CH4, −100 kJ/molCu with H2, −140 kJ/molCu with CO, enthalpy of blast furnace gas combustion=2600 kJ/kg, enthalpy of methane combustion=50000 kJ/kg, heat capacity of all solids=0.8 kJ/kgK; heat capacity of blast furnace gas=1.35 kJ/kgK, heat capacity of H2/N2 product gas from SEWGS=1.6 kJ/kgK, heat capacity of CH4=4.1 kJ/kgK, heat capacity of H2O=2.2 kJ/kgK and heat capacity of the rest of gases=1 kJ/kgK. The displacement of reaction and heat transfer fronts are calculated assuming infinite reaction rates and infinite solid-gas heat transfer rates. The skilled in the art will be able to refine the calculations provided below and adapt them to other conditions. The examples are therefore only illustrative and non-restrictive.
The assumption of an enriched air with oxygen concentration of 45% v facilitates in this particular example an efficient closure of mass and energy balances in a single shaft kiln, because the flow rate of gases at the exit of the calcination zone contains sufficient energy to preheat the incoming limestone flow to said calcination zone. Assuming, a typical superficial gas velocity of the combustion flue gases of 1.5 m/s in the calcination zone, this yields a cross-section for the kiln of 4.5 m. This translates into an average superficial solid velocity moving downwards in the kiln of 0.0004 m/s during each cycle. Assuming a void fraction not occupied by stones of 0.4, an average density of 2000 kg/m3 for the solids and a total height of 8 m for the calcination zone (1, as in other kilns), a total amount of solids of around 29000 kg can be estimated in this zone of the kiln. This translates into a total residence time of the solids in the calcination zone (1) of 6.2 h. Such long residence time in the calcination zone is within the range of stone residence times in existing kilns and can be tuned by the skilled in the art to adapt it to particle size of the stones and calcination rate parameters of the stone. It is shown below to be sufficient to achieve a calcination degree close to 100% in view of experimental results and modelling simulations as in
During each stage under the combustion step i) in the kiln, the previous heat and mass balances indicate that approximately 115 kJ/kg of solids in the calcination zone are absorbed during on average in each combustion step i). After such combustion period is completed, the 15 min of the vacuum step ii) allows for a conversion of CaCO3 to CaO of 6.9% wt during the vacuum step ii), consistent with the experimental and modelling results presented in
For an equivalent single lime shaft kiln with no CO2 capture, operating with the same conditions in the calcination zone, with the same gas and solid flows inlet temperature and heat losses, a maximum thermal input of 3.8 MW can be calculated when using air as comburent, leading to a maximum flow of lime of around 0.84 kg/s. This shows that the method disclosed in this invention to obtain a pure stream of 1 kg/s of CO2, could be implemented in existing shaft kilns without reducing the lime production capacity, when oxygen enriched air is used a comburent and the kiln is retrofitted with additional equipment required for vacuum generation, reinforced walls, lock-hoppers for solid handling minimizing air ingress during vacuum, switching valves etc.
The method of this invention is applied to a calcium-enhanced water gas shift process (CASOH) to treat Blast Furnace Gas (BFG) from a steel mill plant and generate a H2-rich gas while carbonating CaO to CaCO3 in a packed bed. The process is carried out in a system composed of several adiabatic packed-bed reactors operating in parallel in different reaction stages, to allow the continuous treatment of gas streams. The calcination method of this invention recovers pure CO2 when calcining the CaCO3 formed during carbonation of CaO taking place in the CASOH stage.
The first half of the reactor contains a mixture of solids with a molar ratio of active Cu/CaO components of 0.85 when combining the CaO sorbent with the catalyst with 25% w of Cu. This means 2.7% w of active CaO, 87.1% w of inert CaO, 2.6% w of Cu and 7.6% w of Cu-catalyst support. The second half of the reactor contains a mixture of solids with an active Cu/CaO molar ratio of 0.16 combining the CaO sorbent with the catalyst with 5% w of Cu. This means 2.7% w of active CaO, 87.1% w of inert CaO, 0.5% w of Cu and 9.7% w of Cu-catalyst support. This distribution of solids for this particular example will facilitate the operation with a uniform temperature during the heating up by the chemical looping combustion step i), as explained below. The skilled in the art of chemical looping combustion in packed beds could identify other methods in the state of the art to handle the generation of axial temperature distributions in the bed, by introducing additional solids to the bed acting as thermal ballast, by using gas recycles or by introducing additional regenerative heat exchangers or exploiting the excess of thermal energy in high temperature packed beds to drive other endothermic reactions (see Example 3 below).
The CASOH process works cyclically, and the description can therefore be initiated from any stage. Since the application of the method of this invention (steps i) and ii) in
As indicated in
Following the method of this invention, the next step involves vacuum stage ii) in which vacuum pressure (0.05 atm) is applied to extract pure CO2 from the heated solids with CaCO3. The packed bed of solids contains at the beginning of this step sufficient sensitive heat to accomplish the calcination of CaCO3, leaving the bed at 725° C., which is the assumed temperature to start a new CASOH stage in a subsequent cycle. It is assumed that the time given for such a vacuum step (t1=t2=t3=20 min) in this example allows the complete calcination of CaCO3, leading to a production of a steady flow of 1 kg/s of pure CO2 during this period.
It is important to highlight that, from the mass and energy balanced noted above, it is evident that the application of the method of this invention to calcine CaCO3 generated during the CASOH stage almost doubles the efficiency in the generation of H2-rich gas from BFG respect to the previous state of the art design of the CASOH process. In this example, about 60% of the BFG is converted into a H2-rich gas without CO2 while only the remaining 40% of the BFG is used to supply energy for the calcination of CaCO3 during the step i). This is a substantial improvement with respect to the configuration proposed by the state of the art given by Fernandez et al. (2020), in which only 30% of the BFG could be used to produce H2-rich gas from the CASOH stage, while the energy contained in the remaining 70% of the BFG had to be extracted as high temperature heat. Furthermore, it will be evident to the skilled in the art that the CO2 purity of the 1 kg/s gas stream during the vacuum stage must be close to 100% CO2, while the CO2 rich stream in the CASOH state of the art is only about 55-60% v.
A further example of application of this invention is a variant of this Example 2, when feeding a fuel gas containing CO2 to the packed bed of solids containing CaO. The fuel gas can be Blast Furnace Gases (BFG or Oxy-BFG) or Basic Oxygen Furnace gas (BOF) from a steel mill plant, which contain CO. Feeding such gas to packed bed of solids containing CaO will form CaCO3 in the packed bed and generate a gas stream rich in CO. The CO rich gas resulting from this step can be used as a reducing gas in the steel making process, as in other processes for top gas recycle known in the state-of-the-art. In other words, this is a carbonation reaction step similar to the CASOH stage described in the above paragraphs, but without feeding any additional steam with the BFG to the bed of CaO. The remaining steps of the method are similar to the ones described above in this Example 2.
Also note that the application of the method to capture CO2 from a combustion flue gas by carbonation of CaO in a packed bed, would not differ in any essential feature from the case described in the previous paragraph, except in that the resulting decarbonized flue gas would be emitted to the atmosphere. The remaining steps of the method are similar to the ones described above in this Example 2.
The method of this invention is applied to calcine solids containing CaCO3 formed during a sorption enhanced reforming (SER) process to convert methane and steam into a H2-rich product gas, in the presence of CaO as CO2 sorbent to form CaCO3 during SER. As in Example 2, the overall process is carried out in a reactor system composed of several adiabatic packed-bed reactors (see
As the carbonation front moves forward, the Ca-based sorbent is left behind carbonated to its maximum possible conversion (i.e., 3% of the total mol of Ca are carbonated or 2.1% w CaO becomes a 3.7% w of CaCO3). The carbonated part of de packed bed acts as a conventional steam methane reformer. Therefore, the high endothermicity of this reaction reduces the temperature of the bottom carbonated part towards an adiabatic equilibrium temperature of 625° C. (given by the steam-methane reforming equilibrium in the absence of CaO). The CO2 formed is mostly captured downstream in the carbonation reaction front (marked in
As in previous example, once that the previous reaction stage iii), involving carbonation of CaO to CaCO3 is completed, the thermally insulated vessel (1) contains CaCO3 at a certain temperature (625° C. in this example) that is insufficient for calcination. This is when the method of this invention provides a solution of the calcination of such CaCO3 while generating pure CO2. The following stage is therefore a fuel combustion step i) to heat up the solids towards calcination temperatures. This is done in this case using chemical looping combustion principles. In a first substep j) of step i), the oxidation of the Ni-based solids to NiO takes place. A flow of 12.5 kg/s of air (9) at 10 bar and 625° C. allows the complete oxidation of Ni to NiO in only 5 minutes (t4). The heat released due to the oxidation of Ni leads to an increase in temperature in the oxidation front of about 175° C. The high temperature achieved in the oxidation front (i.e., 800° C.) ensures the complete conversion of the oxygen, giving as a result an outlet flow (10) of 9.96 kg/s basically composed of nitrogen, which is discharged at 625° C. The operation is carried out at 10 bar to moderate the leakage of CO2 caused by partial calcination of CaCO3 (PCO2eq at 800° C. is 0.217). In these conditions, the product gas (10) only contains 2.17% v of CO2 and only 8% of the carbonate is calcined in this stage. Once the Ni has been totally oxidized, the bed is divided into two zones at different temperatures as the result of the advance at different velocities of the oxidation and heat exchange fronts. About 6% of the bed (located close to the reactor inlet) is left at 625° C. and the remaining 94% at 800° C. In the following sub-step jj) of i), lasting 15 minutes (t1-t4), the combustion of the fuel gas (8) is completed. The fuel gas (8), used to reduce the NiO to Ni, is a mixture of the syngas (16) generated in the steam reforming stage kk), with the additional flow of methane noted above ((8) is made of 0.60 kg/s of syngas at 685° C. (i.e., the average temperature of the syngas coming from the steam reforming step kk)) and 0.06 kg/s of methane at 25° C.). This flow rate (at an average temperature of 650° C.) allows the complete reduction of NiO in the packed bed in the same duration of the SER stage (i.e., t1=t4+(t1-t4)=20 min). The resulting composition of the reducing gas (49.9% v H2, 8.8% v CO and 8.6% v CH4) makes the reduction of NiO almost thermally neutral, giving as a result a final bed divided into two fractions: a small part of about 7% located at the bottom at 650° C. and the rest at 800° C. In variants of this design, such temperature in the carbonated bed of solids could be increased by increasing the Ni content in the bed at the expense of higher temperatures during the SER stage that will reduce the hydrogen yield.
Once the NiO reduction has finished, and the combustion and heating up step i) is completed, the vacuum calcination of CaCO3 in the bed of solids ii) is carried out to extract pure CO2 while solids cool down as a consequence of the application of vacuum (see
The overall mass and energy balances described above indicate that the application of the method of this invention to the SER process represents a reduction of more than 30% in the energy requirements for calcination (about 20% of the total input of methane to the process is burned during step i) in this process; while 30% is needed in existing Ca—Cu processes derived from the Fernandez et al. (2012).
| Number | Date | Country | Kind |
|---|---|---|---|
| 21382495.6 | Jun 2021 | EP | regional |
| Filing Document | Filing Date | Country | Kind |
|---|---|---|---|
| PCT/EP2022/064747 | 5/31/2022 | WO |
