METHOD AND SYSTEM TO CAPTURE CO2 IN FLUE GASES EMITTED INTERMITTENTLY

Abstract

The disclosure relates to a method to capture CO.sub.2 from the flue gases emitted intermittently from a power plant burning a synthetic fuel in power-to-fuel-to-power systems. The method comprises arranging a reservoir of Ca(OH).sub.2 to feed a flow of such solids to a countercurrent carbonator located in the flue gas path of the power plant, separating the resulting carbonated solids from the CO.sub.2 depleted-gas, storing the carbonated solids in a reservoir of CaCO.sub.3 while the power plant is operating, calcining a steady flow of carbonated solids to produce CaO solids and CO.sub.2 when the power plant is not operating, hydrating the resulting CaO solids with water to replenish the reservoir of Ca(OH).sub.2 and feeding the CO.sub.2 to the power-to-fuel system to manufacture and store the synthetic fuel burned when the power plant is operating.

Claims

1. Method to capture CO.sub.2 from a flue gas (2) emitted intermittently by a fuel turbine (34) power-to-fuel-to-power system (100) when firing with air a carbonaceous synthetic fuel contained in a tank (33), comprising the following steps: a) when the turbine (34) is operating, conducting the flue gas to a carbonator wherein a molar flow of a calcium sorbent feeds from a first reservoir (21) to a carbonator (22) located in the flue gas path of a turbine, forming CaCO.sub.3 containing solids; b) when the turbine (34) is operating, separating the CaCO.sub.3 containing solids (3) obtained in step (a) from the remaining flue gas and storing it in a second reservoir of solids (24); c) when the turbine (34) is not operating, feeding a flow of CaCO.sub.3 containing solids (5) from the second reservoir (24) to an oxy-fired calciner (25), calcining and heating up to a temperature of between 875 C. and 950 C. in presence of a fuel and oxygen; characterized in that the calcium sorbent of step (a) is Ca(OH).sub.2, wherein the flue gas path of a turbine of step (a) is operated with a Ca/C molar ratio between 1 and 2 respect to the carbon flow in the flue gas; wherein the carbonator (22) of the step (a) is a countercurrent carbonator; and wherein the molar flow of the CaCO.sub.3 containing solids (5) of step (c) is of between 1/10 and 1/20 of the molar Ca(OH).sub.2 flow (1), and generating CaO (6) solids and a rich CO.sub.2 stream (14); and wherein it comprises further steps of d) separating CaO solids (6) obtained in step (c) from rich CO.sub.2 stream (14) and return the CO.sub.2 in pure form (7) resulting after purification to the manufacturing plant of synthetic fuel (32) and hydrate the CaO solids (6) with water (8) to produce Ca(OH).sub.2 (9); and e) storing the Ca(OH).sub.2 generated in step (d) in the first reservoir of solids (21) with a capacity between 1 to 2 mol Ca(OH).sub.2 per mol of carbon stored in the fuel tank (33), and re-initiate the sequence in step (a).

2. Method according to claim 1 wherein the flue gas (2) from the fuel turbine at a temperature of between 550 C. and 700 C. is cooled down to a temperature of between 450 C. and 550 C. before entering the countercurrent carbonator (22).

3. Method according to any of claim 1 or 2, wherein the cooling of flue gases from the turbine (16) is carried out by a first section of a heat recovery steam generator of a combined cycle (30) and the flue gas leaving the carbonator (4) at a temperature of between 450 C. and 550 C. is cooled in a second section of the same heat recovery steam generator (31).

4. Method according to any of claims 1 to 3 wherein the molar ratio between the large flow of Ca(OH).sub.2 and the minor flow of CaCO.sub.3 is between 5 and 20.

5. Method according to any of claims 1 to 4 wherein the oxy-fired calciner (25) is supplied by renewable energy such us biomass, renewable electricity or a mixture of H.sub.2 and O.sub.2 from water hydrolysis from renewable electricity.

6. Method according to any of claims 1 to 5, wherein between 1/10 and 1/20 of the Ca(OH).sub.2 is purged from the first reservoir (21) and dispossed to recarbonate in contact with atmosphere.

7. Method according to any of claims 1 to 6, wherein the synthetic fuel containing carbon manufactured in (32) is synthetic natural gas and the fuel turbine (34) firing such fuel is part of a natural gas combined cycle, comprising a further step of feeding of the flue gas leaving the gas turbine (2) into a first heat exchanger section of the heat recovery steam generator (30) of a natural gas combined cycle to produce a cooled flue gas flow containing CO.sub.2 (16); feeding a large flow of Ca(OH).sub.2 solids (1) from the first reservoir (21) with the cooled flue gas flow (16) to a countercurrent carbonator (22) forming CaCO.sub.3 containing solids (3); separating the CaCO.sub.3 containing solids (3) from the lean-CO.sub.2 flue gas (4) to store the solids in the second large reservoir of solids (24) and; feeding the lean CO.sub.2 flue gas (4) into a second heat exchanger section of the heat recovery steam generator (31) and release a low temperature lean CO.sub.2 flue gas (17) into the atmosphere.

8. Method according to claim 1, wherein if the heat recovery steam generators are not available, as it is the case in back-up power plants using open cycle gas turbines, the cooling of the flue gas leaving the gas turbine (2) before the contacting of the Ca(OH).sub.2 flow (1) takes place by mixing with the flue gas (2) with an air flow (18) at ambient temperature to produce the cooled flue gas flow stream (16) entering the carbonator (22).

9. System according to the method described in any of claims 1 to 8, for capturing CO.sub.2 from a flue gas (2) emitted intermittently by a fuel turbine (34) when firing with air a carbonaceous synthetic fuel contained in a tank (33) characterized in that it comprises a first means of conduction configured to conducting the flue gas to a carbonator wherein a molar flow of a calcium sorbent feeds from the first reservoir (21) to a carbonator (22) located in the flue gas path of a turbine, forming CaCO.sub.3 containing solids, when the turbine is operating; a first means of separation configured to separating the CaCO.sub.3 containing solids (3) obtained in step (a) from the remaining flue gas and storing it in a second reservoir of solids (24), when the turbine (34) is operating; a second means of conduction configured to feeding a flow of CaCO.sub.3 containing solids (5) from the second reservoir (24) to an oxy-fired calciner (25); wherein the calcium sorbent of the means for carbonation is Ca(OH).sub.2 and the means of calcination generate CaO (6) solids and a rich CO.sub.2 stream (14); wherein the carbonator (22) is a countercurrent carbonator; a second means of separation configured to separating the generating CaO (6) solids and a rich CO.sub.2 stream (14); a means of purification configured to purifying the rich CO.sub.2 stream (14); a third means of conduction configured to returning the CO.sub.2 in pure form (7) to a fuel manufacturing plant (32) a means of hydration configured to hydrate the CaO solids (6) with water (8); a means of storage configured to storing the Ca(OH).sub.2 generated in the first reservoir of solids (21).

10. System according to claim 9, wherein it further comprises a heat recovery steam generator of a combined cycle (30) configured to cooling flue gases from the turbine (2) a first section of said heat recovery (31) and configured to cooling the flue gases leaving the carbonator (4) in a second section of the same heat recovery steam generator (31).

11. System according to any of claims 9 to 10, wherein it further comprises a means of supply configured to supplying the oxy-fired calciner (25) by renewable energy such us biomass, renewable electricity or a mixture of H.sub.2 and O.sub.2 from water hydrolysis from renewable electricity.

12. System according to any of claims 9 to 11, wherein it further comprises a means of purge configured to purging between 1/10 and 1/20 of the Ca(OH).sub.2 from the first reservoir (21).

13. System according to any of claims 9 to 12, wherein it further comprises a first heat exchanger section of the heat recovery steam generator (30) configured to produce a cooled flue gas flow containing CO.sub.2 (16); feeding a large flow of Ca(OH).sub.2 solids (1) from the first reservoir (21) with the cooled flue gas flow (16) to a countercurrent carbonator (22) forming CaCO.sub.3 containing solids (3); separating the CaCO.sub.3 containing solids (3) from the lean-CO.sub.2 flue gas (4) to store the solids in the second large reservoir of solids (24) and; feeding the lean CO.sub.2 flue gas (4) into a second heat exchanger section of the heat recovery steam generator (31) and release a low temperature lean CO.sub.2 flue gas (17) into the atmosphere, when the synthetic fuel containing carbon manufactured in (32) is synthetic natural gas and the fuel turbine (34) firing such fuel is part of a natural gas combined cycle.

14. System according to any of claims 9 to 13, wherein it further comprises a means of cooling configured to mixing the flue gas (2) with an air flow (18) at ambient temperature to produce the cooled flue gas flow stream (16) entering the carbonator (22) if the heat recovery steam generators are not available, as it is the case in back-up power plants using open cycle gas turbines, where the cooling of the flue gas leaving the gas turbine (2) before contacting of the Ca(OH).sub.2 flow (1) takes place when mixing with the air flow (18) at ambient temperature.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0050] FIG. 1. Schematic of the CO.sub.2 capture processes proposed to close the carbon loop in a power-to-fuel-to-power system (100).

[0051] FIG. 2. The graph shows the evolution of the conversion of Ca(OH).sub.2 to CaCO.sub.3 with the reaction time (dots represents experimental values and trends are represented by solid lines, following an Avrami's model curve).

[0052] FIG. 3. The graph shows the evolution of the molar carbonate conversion achieved by Ca(OH).sub.2 to CaCO.sub.3 with temperature after 30 seconds of carbonation reaction time (i.e. the value of X.sub.max noted in FIG. 2).

EXAMPLES

Example 1

[0053] The particular example of design of the method disclosed in this patent is presented below, with reference to the notation of FIG. 1 to refer to the mass and energy flows most relevant for the method. FIG. 1 discloses a power-to-fuel-to-power (100) producing methane as synthetic fuel containing carbon and using a gas turbine combined cycle with a thermal input of 150 MW.sub.th. The fuel turbine (34) operates with a capacity factor of 0.1 during discharge periods and the fuel manufacturing plant (32) with a capacity factor of 0.9 during charge periods. During discharge periods, a methane flow of 3.0 kg/s is burned in the fuel turbine (34) producing a flue gas flow (2) of 136 kg/s at 630 C. with a CO.sub.2 concentration of 4 vol % CO.sub.2 and containing a CO.sub.2 molar flow of 0.188 kmol/s. The flue gas flow is fed into a first section of the heat recovery steam generator (30) and cooled down up to 530 C. by extracting a thermal flow of 16.3 MW.sub.th before feeding it into the countercurrent carbonator (22).

[0054] The countercurrent carbonator (22) is composed of three cyclone reactors. The CO.sub.2 capture efficiency in the countercurrent carbonator (22) is 0.95 and the CO.sub.2 volume fraction of the flue gas is reduced up to 0.002 before being emitted to the atmosphere. The sorbent conversion and the composition of the gas and solid streams leaving each cyclone reactor has been calculated by assuming an effective reaction time in each reactor cyclone of 4 seconds. The average concentration of CO.sub.2 between the inlet and the outlet of the individual cyclone reactors is found by iteration until the carbon mass balance is fulfilled (carbon disappeared from the gas phase must equal the carbonate formed in the solid phase). This can be done by calculating solid conversion using the Avrami equation, that has been fitted to the experimental data of FIG. 2 with X=(1exp(k t.sup.0.69) X.sub.max (with k=16.46 exp(25040/RT)). Two experimental devices are needed to access experimental information in the time interval of interest. Experimental results marked with black symbols have been obtained in a drop tube reactor with an internal diameter of 100 mm and a total length of 6 m. The device is electrically heated and experiments were carried out at isothermal conditions. Experiments have been carried out with a 5%.sub.vol CO.sub.2 and 20%.sub.vol H.sub.2O in the reacting gases. Commercial Ca(OH).sub.2 with a purity >95% and an average particle size (dp.sub.50) of 5 microns has been used as sorbent. Experimental results marked with white symbols are obtained from a thermogravimetric analyser operated under similar reaction temperature and gas composition, with sample weight of 3 mg of the same material.

[0055] According to the data presented in FIG. 3, a linear dependence of X.sub.max (value noted in FIG. 2) with the carbonation temperature is assumed in the range of temperatures of interest. Following this procedure, it can be estimated that a Ca(OH).sub.2 feeding rate of 0.420 kmol/s is needed in the carbonator reactor to achieve the target volume fraction of 0.002. Therefore, the evolution of said molar carbonate conversion achieved by Ca(OH).sub.2 to CaCO.sub.3 with temperature after 30 seconds of carbonation is the technical limit for the design of the Ca(OH).sub.2 carbonator, as carbonate conversion cannot go over X.sub.max for reasonable gas-solid contact times in the countercurrent carbonator (22).

[0056] Consequently, cooling the flue gas below 450 C., where the equilibrium partial pressure of CO.sub.2 is already as low 0.0027 vol %, allows theoretical capture efficiencies higher than 99%. Carbonation rates measured under these low temperatures remain very high (FIG. 2) but there are indications that maximum carbonation conversion of the solids will be lower (FIG. 3).

[0057] According to scheme depicted in FIG. 1 and the method above, the flue gas stream entering the countercurrent carbonator (22) meets a stream of partially carbonated solids coming from the second cyclone reactor (37) at 552 C. with a carbonate conversion of 0.410 and enters the first cyclone reactor (36) where further carbonation of the sorbent particles takes place. Then, the CaCO.sub.3 containing solids (3) are discharged from the countercurrent carbonator reactor (22) at a temperature of 536 C. with a final carbonate conversion of 0.426 and sent to the second reservoir (24) to be stored.

[0058] The flue gas discharged from the first cyclone reactor (36) at a temperature of 536 C. with a CO.sub.2 volume fraction of 0.0386 is mixed with a stream of solids coming from the third cyclone reactor (38) at a temperature of 577 C. and a carbonate content of 0.357 to react in the second cyclone reactor (37). After reacting, the partially carbonated solids are discharged from the second cyclone reactor (37) at a temperature of 552 C. with a carbonate conversion of 0.410.

[0059] The flue gas discharged from the second cyclone reactor (37) at a temperature of 552 C. with a CO.sub.2 volume fraction of 0.034 is mixed with a flow of Ca(OH).sub.s solids (1) at 350 C. to react in the third cyclone reactor (38). After that, the flue gas leaving the third cyclone reactor (38) is cooled down in the second heat exchanger section of the heat recovery steam generator (31) to a temperature of 105 C., by extracting 73.6 MW.sub.th from the gases before being emitted to the atmosphere (17). This results in a leakage of just 0.009 kmol/s of CO.sub.2 into the atmosphere.

[0060] In order to regenerate the Ca(OH).sub.2, a flow of CaCO.sub.3 containing solids (5) of 2.77 kg/s is fed from the second solid reservoir (24) into an oxy-fired calciner (25). Also, a flow of 0.10 kg/s of fresh limestone (12) at 20 C. is fed into the oxy-fired calciner (25) to compensate the CO.sub.2 leakage in the lean CO.sub.2 flue gas emitted to the atmosphere (17) during discharging periods and to close the carbon loop. In order to ensure a full calcination of the CaCO.sub.3, the oxy-fired calciner (25) operates at temperatures between 875-950 C. (the adiabatic heat balance in the calciner is closed at 910 C. for the conditions of this particular example). This results into a heat demand in the calciner of 8.1 MW.sub.th for calcining the CaCO.sub.3 and heating up the stream of solids entering the calciner (5 and 12). In this specific example, the heat demand is fulfilled by feeding a flow of fuel (10) of 0.06 kg H.sub.2/s and a flow of oxygen (11) of 0.46 kg O.sub.2/s. The gases and solids leaving the oxy-fired calciner (25) are separated in a second solids separation device (26) to produce a rich CO.sub.2 stream (14) and CaO solids (6). The rich CO.sub.2 stream (14) is sent to a purification unit (29) to produce a pure CO.sub.2 stream (7) of 0.021 kmol/s that is fed into the manufacturing plant of synthetic fuel (32). On the other hand, the CaO solids (6) are sent to a hydrator reactor (27) where they react with a stream of 0.43 kg/s of H.sub.2O (8) to produce 0.037 kmol/s of Ca(OH).sub.2 at a temperature of 350 C. Then, a stream of Ca(OH).sub.2 solids (9) of 0.047 kmol/s is diverted to the first reservoir (21) while a stream of Ca(OH).sub.2 solids (13) of 0.001 kmol/s is sent for disposal.

[0061] In another particular configuration of this example, biomass can be used as fuel (10) to fulfil the heat demand in the oxy-fired calciner (25). This results into an input of carbon into the loop that avoids the need of feeding a make-up flow (12) of natural limestone into the system. For the case described above, there is a biomass with a calorific value of 20 MJ/kg, a carbon content of 50 wt % and an oxygen requirements of 1.5 kg O.sub.2/kg biomass, a flow of fuel (10) of 0.41 kg/s and a flow oxygen of 0.61 kg O.sub.2/s of oxygen (11) are fed into the oxy-fired calciner (25) to fulfil the heat demand. After separation in the second solids separation device (26), the rich CO.sub.2 stream (14) containing 0.036 kmolCO.sub.2/s is sent to a to a purification unit (29) to produce a pure CO.sub.2 stream (7) of 0.021 kmol/s that is fed into the manufacturing plant of synthetic fuel (32) and a second stream of pure CO.sub.2 (15) of 0.016 kmolCO.sub.2/s. There are other opportunities to reduce the heat demand in the oxy-fired calciner by preheating the stream of solids entering the reactor (5 and 12) using the rich-CO.sub.2 flue gas (14) and to increase the carbonator temperature by preheating of the Ca(OH).sub.2 solids (1) using the flue gas (4) leaving the countercurrent carbonator reactor (22) are not discussed. It will be obvious for a skilled in the art to take benefit from these method alternatives.