VAPOUR COMPRESSION FOR REGENERATION OF A CAPTURE MEDIUM RICH IN A CAPTURED TARGET GAS

Abstract

A process of regeneration of a capture medium rich in a captured target gas comprising the steps of: heating a target gas rich capture medium comprising an absorbent medium, absorbed target gas and water using thermal energy, thereby facilitating the separation of the absorbed target gas from the capture medium into a gas containing vapour phase and a heated lean capture medium, said gas containing vapour phase comprising the target gas and water vapour having a steam fraction of at least 0.8; compressing at least one of: the gas containing vapour phase; or a vapour phase thermally associated with the gas containing vapour phase, to form a compressed vapour; and using the compressed vapour as a source of thermal energy to heat the target gas rich capture medium.

Claims

1. A process of regeneration of a capture medium rich in a captured target gas comprising: heating a target gas rich capture medium comprising an absorbent medium, absorbed target gas and water using thermal energy, thereby facilitating the separation of the absorbed target gas from the capture medium into a gas containing vapour phase and a heated lean capture medium, said gas containing vapour phase comprising the target gas and water vapour having a steam fraction of at least 0.8; compressing at least one of: the gas containing vapour phase; or a vapour phase thermally associated with the gas containing vapour phase, to form a compressed vapour; and using the compressed vapour as a source of thermal energy to heat the target gas rich capture medium.

2. A process according to claim 1, wherein the gas containing vapour phase comprises the target gas and water vapour having a steam fraction of between 0.8 to 0.999, preferably a steam fraction of at least 0.9.

3. A process according to claim 1, wherein the separation of the absorbed target gas from the capture medium into a gas containing vapour phase comprises one of: flash vaporisation, vaporisation, boiling, stripping or a combination thereof, and optionally, the separation of the absorbed target gas from the capture medium into a gas containing vapour phase occurs in one of: a flash vessel or a desorption vessel, and wherein, optionally, the desorption vessel includes a reboiler, and at least the compressed vapour is used to heat the target gas rich capture medium in the reboiler.

4.-5. (canceled)

6. A process according to claim 1, wherein the vapour phase thermally associated with the gas containing vapour phase comprises a heat transfer fluid, and the process further includes: heating a heat transfer fluid using the gas containing vapour phase to produce said a heated heat transfer fluid, wherein, optionally, the heat transfer fluid comprises a water rich phase, and the process further includes heating the water rich phase using the gas containing vapour phase to produce a steam vapour phase, and wherein, optionally, the water rich phase is heated using thermal heat exchange with the gas containing vapour phase.

7.-8. (canceled)

9. A process according to claim 1, wherein the compressed vapour is used to heat the target gas rich capture medium using at least one of: thermal heat exchange with the target gas rich capture medium; or condensing heat exchange with the target gas rich capture medium, thereafter producing a condensed target gas and water phase, and wherein, optionally, the process further comprises using the condensed target gas and water phase as a further source of thermal energy for heating the target gas rich capture medium, and wherein, optionally, the condensed target gas and water phase composition includes a capture medium content, and the process further includes separating the capture medium from the target gas and water phase to form a separated capture medium, and the separated capture medium is optionally mixed with the heated lean capture medium.

10.-12. (canceled)

13. A process according to claim 1, wherein the vapour phase thermally associated with the gas containing vapour phase comprises a steam vapour phase, and the steam vapour phase is fed into the target gas rich capture medium, thereby directly heating the target gas rich capture medium.

14. A process according to claim 1, wherein the heated lean capture medium is used to heat the target gas rich capture medium using: heat exchange with the target gas rich capture medium.

15. A process according to claim 1, wherein the compressing step comprises a single compression stage or a multi-stage compression stage, and optionally the compressing step comprises a multi-stage compression stage with heat recovered from at least one intercooler located between each compression stage, and optionally each compression stage preferably has a compression ratio of between 1.1 and 10.

16.-17. (canceled)

18. A process according to claim 1, wherein: the target gas comprises at least one of CO.sub.2, H.sub.2S, HCl, HF, SO.sub.2, SO.sub.3 or NO.sub.x, preferably CO.sub.2; and optionally the absorbent medium comprises at least one of: a salt solution, preferably an alkaline salt solution or an amino-acid salt solution; or an amine solution that has an amine vapour pressure that is less than 10% of the CO.sub.2 partial pressure in the gas rich capture medium.

19. (canceled)

20. A process according to claim 1, wherein the compressed vapour is used as a first source of thermal energy to provide thermal energy required to heat the target gas rich capture medium, and the heated lean capture medium is used as a second source of thermal energy to provide thermal energy required to heat the target gas rich capture medium.

21. (canceled)

22. An apparatus for the regeneration of a capture medium rich in a captured target gas and the recovery of captured gas therefrom comprising: a process vessel including a process volume; a supply conduit to feed gas rich capture medium into the process volume; at least one heating arrangement configured to heat the gas rich capture medium to a temperature that facilitates the absorbed target gas to separate from the capture medium into a gas containing vapour phase comprising the target gas and water vapour having a steam fraction of at least 0.8, and to produce a heated lean capture medium; a vapour recompression system fluidly connected to the process volume to receive and compress the gas containing vapour phase from at least one of: the process volume; or a vapour phase thermally associated with the gas containing vapour phase, and form a compressed vapour; a compressed vapour supply conduit to supply compressed vapour to the at least one heating arrangement as a source of thermal energy to heat the target gas rich capture medium.

23. An apparatus according to claim 22, further comprising: a heated lean capture medium supply conduit to supply heated lean capture medium to the at least one heating arrangement as a second source of thermal energy to heat the target gas rich capture medium.

24. An apparatus according to claim 22, wherein the gas containing vapour phase comprises the target gas and water vapour having a steam fraction of between 0.8 to 0.999, preferably a steam fraction of at least 0.9.

25. An apparatus according to claim 22, wherein the process vessel comprises one of: flash vaporisation vessel, vaporisation vessel, boiling vessel, desorption vessel or a combination thereof.

26. An apparatus according to claim 22, wherein the at least one heating arrangement includes a heating arrangement configured to heat the feed gas rich capture medium in the process volume.

27. An apparatus according to claim 22, wherein the at least one heating arrangement comprises at least one heat exchanger, condensing heat exchanger, or at least one mixing heat exchanger; and optionally the heat arrangement includes at least one of: a reboiler, preferably a condensing reboiler, configured to heat the target gas rich capture medium in the reboiler; and at least one preheating arrangement configured to heat the feed gas rich capture medium prior to the feed gas rich capture medium being fed into the process volume, and optionally wherein at least the heated lean capture medium is used to heat the target gas rich capture medium in the at least one preheating arrangement.

28.-30. (canceled)

31. An apparatus according to claim 22, wherein the process vessel comprises a flash vaporisation vessel, and the compressed vapour is used to heat the target gas rich capture medium in the at least one preheating arrangement.

32. An apparatus according to claim 22, wherein the vapour phase thermally associated with the gas containing vapour phase comprises a steam vapour phase, and the apparatus further includes an indirect heat exchanger arrangement configured to heat a water rich phase using the gas containing vapour phase to produce said steam vapour phase, and optionally wherein the indirect heat exchanger comprises at least one heat exchanger, at least one condensing heat exchanger, or at least one mixing heat exchanger.

33. (canceled)

34. An apparatus according to claim 22, wherein the at least one heating arrangement includes: a first preheating arrangement comprising a condensing heat exchanger configured to transfer energy from the compressed vapour to heat the feed gas rich capture medium prior to the feed gas rich capture medium being fed into the process volume and a preheating arrangement, and a second preheating arrangement configured to transfer thermal energy from the target gas and water phase after passing through the condensing heat exchanger to heat the feed gas rich capture medium prior to the feed gas rich capture medium being fed into the process volume.

35. An apparatus according to claim 22, wherein the process vessel comprises a desorption vessel, and the at least one heating arrangement comprises at least one mixing heat exchanger configured to heating the target gas rich capture medium by feeding the compressed vapour into the target gas rich capture medium in the process volume of the desorption vessel or a volume fluidly connected to said process volume.

36.-47. (canceled)

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0111] The present invention will now be described with reference to the figures of the accompanying drawings, which illustrate particular preferred embodiments of the present invention, wherein:

[0112] FIG. 1 illustrates (a) an upstream target gas absorption process that can supply a target gas rich capture medium to a process of the present invention; and (b) one process embodiment of the present invention illustrating an absorbent regeneration using vapour compression of flashed gases (CO.sub.2 and steam) with subsequent heat transfer to rich liquid.

[0113] FIG. 2 illustrates one process example of the present invention illustrating absorbent regeneration using vapour compression of desorber off-gases (CO.sub.2 and steam) with subsequent heat transfer in the reboiler.

[0114] FIG. 3 illustrates one process example of the present invention illustrating absorbent regeneration using vapour compression steam raised by heat exchange with desorber off-gases (CO.sub.2 and water vapour) with subsequent heat transfer to the reboiler.

[0115] FIG. 4 illustrates one process example of the present invention illustrating absorbent regeneration using vapour compression steam raised by heat exchange with desorber off-gases (CO.sub.2 and steam) with subsequent injection of steam into desorber.

[0116] FIG. 5A illustrates one process example of the present invention illustrating absorbent (MEA=mono-ethanolamine) regeneration using vapour compression of flashed gases (CO.sub.2 and steam) with subsequent heat transfer to rich liquid.

[0117] FIG. 5B illustrates one process example of the present invention illustrating absorbent (glycinate) regeneration using vapour compression of flashed gases (CO.sub.2 and steam) with subsequent heat transfer to rich liquid.

[0118] FIG. 6A illustrates one process example of the present invention illustrating process flow sheet and conditions for absorbent regeneration using vapour compression of steam raised by heat exchange with desorber off-gases (CO.sub.2 and steam) with subsequent injection of steam into desorber (lean loading=0.02).

[0119] FIG. 6B illustrates one process example of the present invention illustrating process flow sheet and conditions for absorbent regeneration using vapour compression of steam raised by heat exchange with desorber off-gases (CO.sub.2 and steam) with subsequent injection of steam into desorber (lean loading=0.014).

[0120] FIG. 6C illustrates one process example of the present invention illustrating process flow sheet and conditions for absorbent regeneration using vapour compression of steam raised by heat exchange with desorber off-gases (CO.sub.2 and steam) with subsequent injection of steam into desorber (lean loading=0.006).

DETAILED DESCRIPTION

[0121] The present invention provides a process, apparatus and system for regeneration of capture mediums, such as liquid absorbents, used to capture a target as such as CO.sub.2, in which the thermal energy required for the regeneration process is provided by thermal energy within the desorption process, and more particularly the vapour compression process. The present invention typically forms part of a target gas absorption-desorption process which includes an absorber 100 such as illustrated in FIG. 1(a), coupled with a vapour separator vessel such as flash vessel 200 (FIG. 1(b)) or a desorber/regeneration vessel 200, 200A, 200B (FIGS. 2, 3 and 4).

[0122] The following description relates to the capture of carbon dioxide. However, it should be appreciated that other target gases could equally be used such as H.sub.2S, HCl, HF, SO.sub.2, SO.sub.3 or NO.sub.x which can absorbed by a suitably matched absorption fluid/liquid.

[0123] FIG. 1(a) illustrates a conventional absorption tower system 100 in which where a gas stream including the target gas (CO.sub.2), for example an exhaust gas from combustion of carbonaceous fuel, enters the lower part of absorption vessel/tower 110 through gas inlet line 112. Lean absorbenti.e. absorbent that is stripped for CO.sub.2is introduced into the upper part of the absorption tower 100 through a lean absorbent line 120. The gas flows from the bottom to the top of the absorption tower 110 countercurrent to a lean absorbent, which flows from the top to the bottom of the absorption tower 110. Lean gas, i.e. exhaust gas where a substantial part of the CO.sub.2 is removed, is removed through a gas exit line 130 at the top of the absorption tower 110, whereas rich absorbent, i.e. absorbent having absorbed CO.sub.2, exits from the absorption tower 110 through a rich absorbent line 140.

[0124] The regeneration of aqueous absorption liquids used in absorption systems such as shown in FIG. 1(a) applied for the capture of acid gases such as CO.sub.2 from gas streams is usually carried out at elevated temperature (90 to 150 C.) when CO.sub.2 is released or desorbed from the aqueous solution as a result of the thermal reversibility of the capture process. This is achieved through the supply of heat at the required temperature level from a suitable source, such as the steam cycle of a power plant, a separate boiler, a solar thermal process, a geothermal process, or any other heat source. The supply of heat commonly results in an additional emission of CO.sub.2 that needs to be accounted for in the overall emissions reductions. In some circumstances heat sources are not available and the heat needs to be generated in a separate facility that results in significant, additional CO.sub.2-emissions.

[0125] One type of regeneration system is shown in FIGS. 2 and 3, comprising a desorber vessel/column 310. The rich absorbent 140 from the absorber tower 100 is heated against lean absorbent that is returned to the absorption tower in a heat exchanger 312 (FIG. 1(a)) to a temperature typically in the range between 9 and 110 C., before the rich absorbent is introduced into a desorber column 310 via inlet conduit 340. In the desorber column 310 the rich absorbent flows downwards, countercurrent to steam generated by heating some of the absorbent in a desorber reboiler 315. Lean absorbent leaves the desorber column 310 through a lean absorbent outlet 316. A part of the lean absorbent in the outlet 316 is introduced into the desorber reboiler 315 through line 317 where it is heated to a temperature typically in the range between 12 and 130 C., to produce hot absorbent and steam which is re-introduced into the desorber column 310 through line 318. The heating of the desorber column 310 from the bottom gives a temperature gradient at steady state from the bottom to the top of the column 310, where the temperature at the top is from 10 to 50 C. lower than at the bottom, depending on the actual design of the column 310. The pressure in the desorber column 310 is normally atmospheric pressure or higher to obtain an effective desorption of the absorbent or stripping of CO.sub.2. The pressure in the desorber column 310 is often 1.5 bara or higher. In a practical situation, the pressure is often from about 1.5 to about 2.0 bara but may even exceed this pressure.

[0126] The lean absorbent in line 316 that is not introduced into the desorber reboiler 315, is recycled back to the absorption column 110 (FIG. 1(a)) through the line 320 and cooled in the heat exchanger 312 against rich absorbent in the line 140. In the heat exchanger 315 the relatively cold rich absorbent is heated against the relatively hot lean absorbent leaving the absorption tower 110 at a temperature of about 120 C. Depending on the actual dimensioning and construction of the plant, the temperature of the absorbent leaving the heat exchanger 312 for the absorption tower 110 may be from about 90 to about 110 C.

[0127] CO.sub.2 released from the absorbent, water vapour and minor amounts of absorbent form a gas containing vapour phase, which is withdrawn from the desorber column 310 through a gas withdrawal line 350. It should be appreciated that the water vapour withdrawn through line 350, and the condensed water removed in the desorption column 310, may comprise minor amounts of absorbent. The water and water vapour from this line 350 therefore typically include water and water vapour including minor amounts of absorbent, where appropriate. The processing of that gas containing vapour phase from gas withdrawal line 350 depends on the particular embodiment of the present invention described below.

[0128] In conventional regeneration systems, the lean absorbent in the reboiler 315 is typically heated by means of electricity, or a heating medium, such as steam. Generally, the provision of thermal energy via electricity, for example using resistive or inductive heating, is considered wasteful and electrically driven heat pumps are preferred because of their higher efficiency. In such a heat pump application, gases are compressed, resulting in a temperature increase with heat recovery via condensation of the liquids from the compressed gases.

[0129] The present invention proposes the use of vapour compression systems which more effectively recover thermal energy from that system needed for regeneration of the aqueous absorption liquids and CO.sub.2-desorption. Two systems are proposed: [0130] (1) a direct vapour compression system (FIGS. 1(b) and 2); and [0131] (2) an indirect vapour compression system (FIGS. 3 and 4).

[0132] The starting point for both the direct and indirect vapour compression systems is to analyse the thermal energy requirement per unit mass of CO.sub.2 released, the so-called specific reboiler duty (MJ/kg CO.sub.2), in more detail. This thermal energy requirement has the following three contributions: [0133] 1) The energy required to break the bond between CO.sub.2 and the active components in the absorption liquid as determined by its chemical formulation. [0134] 2) The heat required for bringing the solvent up to the regeneration temperature, which is a function of the CO.sub.2 loading of the liquid absorbent and temperature difference across the CO.sub.2-desorber. [0135] 3) The evaporation enthalpy for the steam that leaves the CO.sub.2-desorber together with the CO.sub.2, which is determined by the temperature of the steam-saturated CO.sub.2-product leaving the CO.sub.2-desorber at the top.

[0136] The minimisation of the thermal energy requirement is an important objective in CO.sub.2-capture technology development work involving extensive evaluation of different absorption liquids, considering their intrinsic characteristics, such as vapour-liquid equilibria and reaction enthalpies, and process conditions such as regeneration temperature and pressure in the CO.sub.2-desorber. In particular, as the amount of steam leaving the CO.sub.2-desorber represents a loss of energy, process and design optimisation focuses on the minimisation of the steam/CO.sub.2 ratio in the stream exiting the CO.sub.2-desorber. The Applicant's U.S. Pat. No. 10,040,023, Process and apparatus for heat integrated liquid absorbent regeneration through gas desorption is an example of how the thermal energy requirement can be minimised by new process and equipment designs.

[0137] As discussed in the background, a number of prior art processes have been developed that use compression of the CO.sub.2-stream from the desorber with subsequent heat recovery to lower the thermal energy requirement for CO.sub.2 desorption. However, none of these processes include a CO.sub.2-desorption process in which all the thermal energy required is provided by energy transfer from the regeneration stage/desorption stage within that process. This is because energy efficient liquid regeneration process concepts are based on the minimisation of the water content of the wet CO.sub.2-product leaving the CO.sub.2-desorber. Such processes fail to recognise the full benefits of the vapour compression process.

[0138] In the present invention, the steam content in the CO.sub.2-product in the vapour compression process is high (within the steam fractions discussed below) so to achieve high recovery of energy via the condensation of water to supply the overall heat requirement of the capture process. This is in contrast to the common approach of minimisation of steam content in the exit stream from the CO.sub.2-desorber in prior art regeneration processes.

[0139] As noted above, latent heat contained in the exit gas stream from the CO.sub.2-desorber can be recovered and reused in the regeneration of the absorption liquid and CO.sub.2-desorption through two different pathways means: [0140] 1. Directly using the condensate from the steam/CO.sub.2 stream (for example the systems illustrated in FIGS. 1(b) and 2); or [0141] 2. Indirectly using steam generation via heat recovery of the condensate (for example the systems illustrated in FIGS. 3 and 4).

Direct Latent Heat Recovery

[0142] In addition to heat exchange between the lean liquid stream and rich liquid stream (for example as shown using heat exchanger 312 in FIG. 1(a)), the direct system entails compression of the exit gas stream from the CO.sub.2-desorber vessel, which consists of a mixture of steam and CO.sub.2, with subsequent heat recovery to such an extent that it suffices to provide the total additional thermal energy requirement for the CO.sub.2-desorption process. As such the CO.sub.2-capture process can then be operated using electrical energy only, that can be sourced from renewable energy sources such as solar cells or wind turbines. Coarsely speaking two pathways for heat recovery have been identified by the inventors:

Concept 1

[0143] As illustrated in FIG. 1(b), the system 200 has incoming rich liquid absorbent 240 fed into a flash vessel 210 to generate CO.sub.2 and steam (line 250) and separate lean liquid (line 220) through flash vaporisation. Flashing is followed by compression of the CO.sub.2/steam mixture in compressor 260 and recovery of heat from the compressed vapour stream in condenser 215 via condensation of water. Heat is transferred in condenser 215 to the incoming rich liquid absorbent 240, to produce a further heated rich liquid absorbent 240A. The pressure in the flash vessel 210 is normally atmospheric pressure or higher to obtain an effective desorption of the absorbent, or stripping of CO.sub.2, and is more typically 1.5 bar or higher. Practically, the pressure is often from about 1.5 to about 2.0 bar but may even exceed this pressure.

[0144] This first concept is akin to a Mechanical Vapour Recompression (MVR) process in which low pressurisation and high rates of evaporation are used. MVR is used in industrial applications such as the production of freshwater from seawater and in liquid and solids dehydration.

Concept 2

[0145] This concept (system 300A) includes the previously described desorber column 310 and desorber reboiler 315 and processes the gas containing vapour phase withdrawn from the desorber column 310 through a gas withdrawal line 350. As illustrated in FIG. 2, the CO.sub.2/steam mixture from gas withdrawal line 350 of desorber column 310 is compressed in compressor 360 and then the resulting compressed vapour (in line 370) is fed into the reboiler 315 in which heat is recovered from the compressed vapour stream via condensation of water and transferred to/heats up the recirculating lean absorption liquid at the bottom of the desorber column 310 thus generating steam required for the desorption of CO.sub.2 in the desorber column 310. This is akin to the steam stripping process in standard desorber columns where steam is generated at the bottom of the desorber column 310.

[0146] Compression of the steam/CO.sub.2 stream is advantageous for most subsequent storage and utilisation options which often require CO.sub.2 at elevated pressure, for example in excess of 100 bar, if CO.sub.2 is used for geological storage. Such high pressures are not needed for other cases where CO.sub.2 is used as reactant to produce carbonates or fuels or where CO.sub.2 is used to promote plant growth.

Indirect Latent Heat Recovery

[0147] The indirect process is based on the transfer the latent heat in the exit gas stream from the CO.sub.2-desorber via condensation to an incoming water stream that will be converted into steam. The latent heat from the high steam fraction in the desorber exit can thus be converted into pure steam that is subsequently compressed with heat recovery via condensation by the following pathways:

Concept 3

[0148] This concept (system 300B) also includes the previously described desorber column 310 and desorber reboiler 315 and processes the gas containing vapour phase withdrawn from the desorber column 310 through a gas withdrawal line 350. As illustrated in FIG. 3, the CO.sub.2/steam mixture from gas withdrawal line 350 of desorber column 310 passes into heat exchanger 440 which transfers the latent heat of the CO.sub.2/steam mixture from gas withdrawal line 350 to water in line 442 to produce a water vapour/steam phase which exits the heat exchanger 440 in line 450. The CO.sub.2/steam mixture exits the heat exchanger in line 473 as a wet CO.sub.2 phase and water/condensate. The steam phase in line 450 is then compressed in compressor 460 and fed into the reboiler 315 via line 462. Heat transfer to the absorption liquid in the desorber column 310 is via condensation of steam and latent heat transfer in the reboiler 315 and/or via a preheating arrangement (not illustrated) to the incoming rich solution in line 340. Condensate exits the system in line 470 which can be recycled back to heat exchanger 440 via line 442 to generate further steam. This will facilitate a complete closed vapour compression-based steam cycle for heat recovery from the desorber column 310 and heat supply to the desorber column 310. A closed system is can also be based on other working fluidsthermal transfer fluidsother than water.

Concept 4

[0149] As illustrated in FIG. 3, the system 300C has the CO.sub.2/steam mixture from gas withdrawal line 350 of desorber column 310 passing into heat exchanger 440 which transfers the latent heat of the CO.sub.2/steam mixture from gas withdrawal line 350 to water in line 442 to produce a water vapour/steam phase which exits the heat exchanger 440 in line 450. The CO.sub.2/steam mixture exits the heat exchanger in line 473 as a wet CO.sub.2 phase and a water/condensate in line 474 (see for example FIG. 6A). That steam phase in line 450 is then compressed and fed directly into the desorber column 310. Injection of the steam into the desorber column 310 is used to replace the reboiler as heat supply and for desorption of CO.sub.2. This embodiment does not require any additional heat exchange equipment. The input water stream can be made up by the condensate from the desorber column 310 with the resulting steam acting as a stripping gas in the desorber column 310.

[0150] The process operation in both direct and indirect process will be most beneficial at high steam content of the CO.sub.2 product in the desorber as this will lead to the highest recovery of latent heat from the stream fraction. In this regard, each of the processes illustrated in FIGS. 1(b) to 4 operate with the following operational parameters: [0151] Steam/CO.sub.2 ratio in exit stream of CO.sub.2 desorber needs to be high to enable significant recovery of latent heat in additional to recovery of the sensible heat. This will require a steam fraction of at least 0.8, preferably above 0.9.

[0152] A variety of liquid absorbents can be used. The preferred absorption liquids will be: [0153] salt solutions, i.e. solutions without vapour pressure, such as alkaline salt solutions or amino-acid salt solutions. Exemplary examples include: Potassium, lithium and sodium salts of carbonate, phosphate, glycine, taurine, alanine, sarcosine, proline, lysine, methyltaurine, methionine, aminohexanoic acid, phenylalanine, glutamic acid, arginine aspartic acid, leucine, serine, threonine, glucosamine, dimethyl-glycine, methyl-amino-propionic acid, amino-butyric acid, pipecolic acid preferably having concentrations between 0.005 and 5.0 M, more preferably between 0.5 and 2.5 M. [0154] It should be appreciated that a variety of other absorbents could be used, for example as taught in Australian Patent No. AU552657B2 entitled amino-acids for gas absorption, United States Patent No. U.S. Pat. No. 2,176,441 entitled Removal of gaseous weak acid from gases containing the same; and Canadian Patent No. CA 619,193 entitled Process for separating carbon dioxide from gas mixtures, the contents of which should be understood to be incorporated into this specification by this reference; [0155] Amine solutions that have an amine vapour pressure that is less than 10% of the CO.sub.2 partial process both determined at the CO.sub.2-desorber exit would be acceptable to use.

[0156] Further optimisation for different absorption liquids and under varying process conditions will entail a trade-off between power consumption for compression and lean loading requirement. A low lean loading will be beneficial for the CO.sub.2-absorption process, resulting in higher mass transfer and therefore smaller equipment. However, this entails a deeper degree of absorbent regeneration that results in a higher specific power consumption.

[0157] The presence of amine vapours will result in co-absorption of CO.sub.2 and amines into the condensed water, which results in part removal of the CO.sub.2 due to the reaction with amine. If the amine vapour pressure is small compared to the CO.sub.2 pressure, for example 10%, this loss is acceptable.

[0158] The compression arrangements in the systems shown in FIGS. 1(a) to 4 can be single stage or multi-stage with heat recovered from the intercoolers between each stage. Typical compression ratios will be between 1.1 and 10 for each stage. As explained previously, the compression ratio is dependent on the steam/CO.sub.2 ratio in the exit stream and the requirements stemming from the uses of the CO.sub.2-product. In the range of steam fractions 0.8 to 0.999 in the exit stream a higher compression ratio would generate more sensible heat that is most beneficial for lower zone range, whereas a low compression ratio would be most effective at the higher zone range. These considerations might be overruled by the requirement to deliver CO.sub.2 at elevated pressure in which case a high compression ratio is desirable regardless of the steam/CO.sub.2 ratio. This will be the case if the CO.sub.2 is sent to be transported as supercritical liquid to a geological storage reservoir.

[0159] It should be appreciated that further process steps or stages may also be included in the concepts described above. For example, the target gas and water/water vapour phase from line 250 (FIG. 1(b) and line 350 (FIGS. 2, 3 and 4) may include minor amounts of absorbent. Once condensed (i.e. passed through heat exchanger/condensers 215, 315, 440), the process may further include a separator (for example separator 280 in FIG. 5A or separator 380 in FIG. 5B) configured to separate the absorbent from the target gas and water/water vapour). That separated absorbent can be mixed back into the lean liquid being recycled back to the absorber column 110 (FIG. 1(a)). It should be noted that this additional stage may entail a deeper degree of absorbent regeneration that results in a higher specific power consumption.

EXAMPLES

[0160] The following examples were developed using a process model developed in ProTreat simulatora process simulator tool for gas treating available from Optimized Gas Treating, Inc. (OGT), Buda, Texas, United States of America.

Example 1Concept 1

[0161] An example of absorbent regeneration system based on concept 1 (FIG. 1(b)) is provided in FIG. 5A for a process that uses an aqueous 2M Monoethanolamine (MEA) solution. The same reference numerals have been used in FIG. 5A for like features as used in FIGS. 1(a) and 1(b).

[0162] The process model was developed in ProTreat discussed above based in the process flow diagram provided in FIG. 5A. In the model, the rich liquid 140 coming from the absorber 110 was determined to have a CO.sub.2-loading of 0.3 mol/mol when capturing CO.sub.2 from the air. The rich solution 140 was heated up in two heat exchangers 312 and 215 and subsequently sent to a flash vessel 210 in which vapour was released at 1.5 bar in vapour stream 250. The vapour stream 250, having a steam fraction of 0.99, was subsequently compressed in compressor 260 to 3 bar with the temperature of the vapour mixture increased to 208 C. The vapour is partially condensed in condenser/heat exchanger 215, and heat is transferred to the incoming rich absorption liquid. That condensed vapour stream passes into separator 280 which separates CO.sub.2/water (line 275A) and any absorbent liquid (line 275B). The absorbent liquid is mixed into the hot-lean liquid 220 through mixer 282 to form a mixed hot-lean liquid 220A. After the flash the lean liquid leaves the flash vessel 210 having a lean loading of 0.056 mol/mol and passes through heat exchanger 312 transferring heat to the incoming rich solution 140 and is then recycled back to the absorber to be used to capture CO.sub.2 from the air.

[0163] The specific power consumption for the vapour compression (efficiency=0.8) is equal to 1.94 MWh/tonne CO.sub.2.

Example 2Concept 1

[0164] A second example of the process of concept 1 (FIG. 1(b)) is shown in FIG. 5B. Again, the same reference numerals have been used in FIG. 5A for like features as used in FIGS. 1(a) and 1(b). As with the first example, the process was modelled in ProTreat for an amino-acid salt solution (2M sodium glycinate) based on the process flow diagram shown in FIG. 5B.

[0165] In this example under similar process conditions as described above for Example 1, the H.sub.2O vapour fraction of the vapour stream from the flash operation was lower (0.835) together with a lower compression ratio (1.5 bar to 2.2 bar) and a lower degree of regeneration (lean liquid loading=0.088 mol/mol) was achieved. In addition, an additional heat exchanger 313 is included transferring latent heat from the condensed vapour stream 275 to the rich solution 140 from the absorber 110 (FIG. 1(a)). The vapour stream 275A is also shown as going through separator 284 to separate the CO.sub.2 content and water content thereof.

[0166] The overall process resulted in a lower compression energy requirement of 1.12 MWh/tonne CO.sub.2.

[0167] It is considered that further optimisation for different absorption liquids and under varying process conditions will entail a trade-off between power consumption for compression and lean loading requirement. A low lean loading will be beneficial for the CO.sub.2-absorption process, resulting in higher mass transfer and therefore smaller equipment. However, this entails a deeper degree of absorbent regeneration that results in a higher specific power consumption.

Example 3

[0168] A process modelling example is provided where the exhaust from the desorber is compressed to 10 bar via a two-stage adiabatic compression process following the process illustrated in FIG. 2. That pressure level would be suitable to supply CO.sub.2 to a methanation process in which CO.sub.2 is reacted with hydrogen to produce methane.

[0169] The compression process of the process in FIG. 2 was modelled estimating the heat that was produced from the compression process by cooling down the compressed vapour to 135 C. and compared this to estimates for the reboiler duty. The reboiler duty was estimated for three cases of the heat associated with the CO.sub.2 absorption process representing the relevant range reaction enthalpies. The heat for production of steam that is present with CO.sub.2 was then added to this. In this modelling, two desorber exhaust temperatures, 100 C. and 120 C. at two different pressure levels are considered with the temperature of the compressed vapour decreased to 135 C. at the end of each stage with heat transferred in the reboiler as a result of the water condensation. Various steam fractions were considered, and the process simulation results (specific compressor power consumption and specific heat recovery from compressors) are presented in Table 1.

[0170] Table 1 indicates where the heat from the compression process exceeded the reboiler heat requirementthus not providing a result for a single case but rather serving to provide the minimum steam ratio range. It was assumed that there was negligible heat requirement for heating up the absorption liquid in the desorber. Referring to Table 1, it can be seen that at higher steam fraction the specific compressor power consumption increases but also the amount of recoverable heat. At the higher inlet temperature of 120 C. and higher exit pressure of 2 bar the specific power consumption is lower than for 100 C./1 bar. The amount of heat recovery is slightly lower.

TABLE-US-00001 TABLE 1 Specific compressor power consumption and specific heat recovery for desorber exit vapour compression at high water content (2-stage compression; 80% efficiency compressor) and comparison with reboiler duty Minimum Minimum Minimum Specific reboiler reboiler reboiler Outlet compressor Specific heat duty; heat of duty; heat of duty; heat of Steam Inlet Inlet Outlet temperature power recovery from reaction = 50 reaction = 75 reaction = 100 fraction pressure temperature pressure after cooling consumption compressors kJ/mol CO.sub.2 kJ/mol CO.sub.2 kJ/mol CO.sub.2 [] [bar] [ C.] [bar] [ C.] [MWh/tonne CO.sub.2] [GJ/ton CO.sub.2] [GJ/ton CO.sub.2] [GJ/ton CO.sub.2] [GJ/ton CO.sub.2] 0.50 1 100 10 135 0.13 0.89 () 1.99 2.56 3.13 0.60 1 100 10 135 0.17 1.45 () 2.41 2.98 3.55 0.70 1 100 10 135 0.23 2.38 () 3.13 3.69 4.26 0.80 1 100 10 135 0.33 4.22 () 4.55 5.11 5.68 0.90 1 100 10 135 0.73 9.98 (+) 8.81 9.38 9.94 0.95 1 100 10 135 1.47 21.22 (+) 17.33 17.90 18.47 0.95 2 120 10 135 0.97 19.90 (+) 17.33 17.90 18.47

TABLE-US-00002 TABLE 2 Heat recovery (H in GJ/ton CO.sub.2) for desorber exit vapour compression at high water content and compressor power consumption (P in MWh/tonne CO.sub.2); single stage compression from 1 bar; 100% efficiency compressor; full latent heat recovery from water condensation Minimum Minimum Minimum reboiler reboiler reboiler duty; heat of duty; heat of duty; heat of Steam Exit Exit Exit Exit reaction = 50 reaction = 75 reaction = 100 fraction pressure = pressure = pressure = pressure = kJ/mol CO.sub.2 kJ/mol CO.sub.2 kJ/mol CO.sub.2 [] 1.25 bar 2.5 bar 5 bar 10 bar [GJ/ton CO.sub.2] [GJ/ton CO.sub.2] [GJ/ton CO.sub.2] 0.50 H = 0.90 H = 1.05 H = 1.24 H = 1.47 1.99 2.56 3.13 P = 0.010 P = 0.043 P = 0.084 P = 0.134 0.60 H = 1.33 H = 1.53 H = 1.76 H = 2.05 2.41 2.98 3.55 P = 0.011 P = 0.054 P = 0.105 P = 0.168 0.70 H = 2.06 H = 2.32 H = 2.63 H = 3.01 3.13 3.69 4.26 P = 0.016 P = 0.072 P = 0.140 P = 0.224 0.80 H = 3.52 H = 3.90 H = 4.37 H = 4.95 4.55 5.11 5.68 P = 0.024 P = 0.108 P = 0.211 P = 0.336 0.90 H = 7.89 H = 8.66 H = 9.60 H = 10.75 8.81 9.38 9.94 P = 0.048 P = 0.216 P = 0.421 P = 0.672 0.95 H = 16.63 H = 18.17 H = 20.05 H = 22.34 17.33 17.90 18.47 P = 0.095 P = 0.432 P = 0.843 P = 1.34 0.99 H = 86.55 H = 94.26 H = 103.7 H = 115.1 85.5 86.1 86.7 P = 0.48 P = 2.16 P = 4.21 P = 6.72 0.995 H = 174.0 H = 189.4 H = 208.2 H = 231.1 170.7 171.2 171.9 P = 0.95 P = 4.32 P = 8.43 P = 13.4 0.999 H = 873.2 H = 950.3 H = 1044.4 H = 1159 852.6 853.1 853.7 P = 4.75 P = 21.6 P = 42.1 P = 67.2

[0171] Table 1 also shows the specific energy requirement for regeneration of the absorption liquids, the so-called reboiler duty. In the calculation, the regeneration energy is calculated as the sum of the reaction enthalpy requirement for CO.sub.2-desorption and the latent heat of evaporation for water. In this relatively simple analysis, the heat requirement for heating up the absorption liquid after exchanging heat with the regenerated solution is ignored, assuming an ideal temperature approach for the heat exchange. The reboiler duty thus calculated is a minimum and has been based on three values for the reaction enthalpy for the chemical absorption, 50, 75, 100 KJ/mol CO.sub.2, which cover the range of values for chemical absorption liquids. For water evaporation an average value of 37.5 KJ/mol H.sub.2O was used for the temperature range 100 to 120 C. The reboiler duty increases with an increase in the reaction enthalpy for the chemical absorbent chosen.

[0172] In Table 1 it can be seen that for steam fractions equal to 0.8 and lower the compressor is not able to provide sufficient energy for the regeneration of the absorption liquid. At a steam fraction of 0.9 and higher there is sufficient heat available for regeneration of absorption liquid. At higher steam fraction the overall compression energy requirement will increase, and the optimal conditions are those at which the water vapour content is high enough to provide the heat to the reboiler via the vapour compression process but not higher than that, as this would unnecessarily increase the energy requirement. The water vapour content is determined by the characteristics of the absorption liquid and will be dependent on CO.sub.2-loading of the absorption liquid and temperature. For the example of compression to 10 bar it is shown that the break-even steam fraction would be between 0.8 and 0.9. This break-even point is also dependent on the compression ratio of the vapour compression process. In another example the pressure ratio for the vapour compression is varied between 1.25 and 10 assuming a single-stage compression process and the vapour exiting the desorber at 100 C./1 bar. To understand the trends and the potential for heat recovery, in this calculation a 100% compressor efficiency was used and it was assumed that all latent heat could be recovered.

[0173] Table 2 presents results the specific heat recovery (H) and the specific compressor power consumption (P) for four compression ratios (1.25, 2.5, 5, 10) and a range of steam fractions of the exit vapour from the CO.sub.2 desorber (0.5 to 0.999). Also presented are the estimated minimum specific reboiler duties for the CO.sub.2-desorption process for three values of the reaction enthalpy. These reboiler duties can be compared with the heat recovery from the compression process (H). The values where the heat from the compression process exceeds the reboiler duty for all values of the reboiler duty are shown in bold for clarity. At low compression ratio the steam fraction needs to be high for to achieve high enough heat recovery from the compression process, e.g. at 1.25 bar pressure after compression, the steam fraction needs to be at least 0.99 to 0.995 to achieve. At 10 bar pressure after compression the steam fraction can be somewhat lower, i.e. more than 0.8 to 0.9 to achieve break-even. Higher steam fractions lead to steep increases in the specific power consumption as shown in Table 2. This indicates that careful optimisation is needed for the specific absorption liquid, CO.sub.2-loading and regeneration process design to determine the optimum compression conditions that will minimise compression power consumption and achieve sufficient heat recovery from the compression process.

[0174] The simplified analysis here indicates that for the relevant pressure conditions of 1.25 to 10 bar the steam fraction will need to be in the range 0.8 to 0.995 to achieve the situation where the heat of the compression process covers is sufficient to supply the reboiler duty. At lower vapour fraction the heat generated by the compression will be insufficient to fully cover the reboiler duty. At higher vapour fractions the equipment might be become quite large and expensive as large amounts of water vapour need to be treated relative to the amount of CO.sub.2. Also, the compressor power consumption increases quite rapidly which is not desired.

Example 4Concept 4

[0175] The process configuration shown in FIG. 4 was modelled in Protreat for a range of conditions with differing degrees of regeneration using a 2M sodium glycinate as a representative amino-acid salt solution.

[0176] The absorption liquid entering the regeneration system had a liquid loading equivalent to 0.3 mol/mol that can be achieved through CO.sub.2-capture from the air. The desorber was operated at a pressure of 150 kPa (1.5 bar) under isothermal conditions (112 C.) for all process simulations. The steam compressor was assumed to have an 80% efficiency. The rich solution 140 was preheat in heat exchangers 312 using the hot lean liquid exiting the bottom of the desorber column 310. The heated rich liquid is subsequently fed into desorber column 310.

[0177] The resulting process flow sheets are given in FIGS. 6A, 6B, 6C for lean loadings equal to 0.02, 0.014 and 0.006. The flow sheets indicated the modelled conditions. The results are also summarised in Table 3. It is noted that the same reference numerals have been used in FIG. 6A, 6B, 6C for like features as used in FIGS. 1(a) and 4.

TABLE-US-00003 TABLE 3 Results summary for absorbent regeneration using vapour compression steam raised by heat exchange with desorber off-gases with subsequent injection of steam into desorber Steam fraction Specific compression Lean loading in CO.sub.2-desorber Compression power consumption [mol/mol] exit ratio [MWh/tonne CO.sub.2] 0.006 0.994 1.34 1.17 0.016 0.988 1.55 0.92 0.02 0.977 2.2 0.85

Example 5Concept 3

[0178] The process configuration shown in FIG. 3 was modelled in Protreat for a range of conditions with differing degrees of regeneration using a 2M sodium glycinate as a representative amino-acid salt solution.

[0179] The rich absorption liquid 340 entering the regeneration system at a temperature of 104.5 C. (=0.3, 14420 kg/hr) had a liquid loading equivalent to 0.3 mol/mol that can be achieved through CO.sub.2-capture from the air. The desorber 310 was operated at a pressure of 150 kPa (1.5 bar) and a bottom temperature of 115 C. with the regenerated, lean absorption liquid 320 leaving the regeneration at a temperature of 112.2 C., (=0.036, 14120 kg/hr) having a liquid loading equivalent to 0.036 mol/mol.

[0180] The vapour mixture leaving the desorber 350 exchanged heat with the incoming water stream 442 (@40 C., 8150 kg/h) that resulted in the production of steam 450 at a temperature of 110 C. and a pressure of 150 kPa (8150 kg/hr). CO.sub.2 and condensate exit in stream 475 at 45 C., a pressure of 150 kPa (CO.sub.2: 290 kg/h). This steam is compressed using the compressor 460 to 290 kPa (8150 kg/hr) with the adiabatic compression resulting in an outlet temperature of 195 C. in line 462 and condenses in the reboiler 315 at a temperature of 132.5 C. in the condensate 470 transferring heat to the absorption liquid which exits the reboiler in stream 318. The steam compressor was assumed to have an 80% efficiency and the electricity consumption of the process equals 1.2 MWh/tonne CO.sub.2.

[0181] Those skilled in the art will appreciate that the invention described herein is susceptible to variations and modifications other than those specifically described. It is understood that the invention includes all such variations and modifications which fall within the spirit and scope of the present invention.

[0182] Where the terms comprise, comprises, comprised or comprising are used in this specification (including the claims) they are to be interpreted as specifying the presence of the stated features, integers, steps or components, but not precluding the presence of one or more other feature, integer, step, component or group thereof.