Patent Application
20110203313
This invention relates to the recovery of carbon dioxide and hydrogen in a concentrated form from a synthesis gas stream comprising hydrogen and carbon dioxide thereby generating a carbon dioxide stream that may be used in a chemical process, or may be sequestered or used for enhanced oil recovery before being ultimately sequestered, and a hydrogen rich stream that may be used as fuel for a power plant thereby generating electricity or as fuel for a low pressure burner of a fired heater, or as fuel for a reformer or boiler or as a refinery feed stream for upgrading of one or more refinery streams or as a hydrogen feed to a chemical process.
U.S. Pat. No. 3,614,872 relates to an autorefrigeration process for separating a shifted synthesis gas feedstream into an enriched carbon dioxide product stream and an enriched hydrogen product stream comprising the steps of:
According to the description of U.S. Pat. No. 3,614,872, the dried feed gas enters the separating portion of the system at substantially initial line pressure, that is a pressure of about 1400 psig (96.5 barg). It is also said that at start-up, back pressure valve 20 is closed and the enriched hydrogen product stream from the top of separator 19 is reduced in pressure from 1400 psig to about 140 psig (9.65 barg) by being passed across an expansion valve. By expansion across the valve, the temperature of this gaseous stream is dropped to −78° F. without solid formation. The cooled enriched hydrogen product gas is then used as internal refrigerant in cooler 13. It is also said that higher refrigeration efficiencies are possible if the compressed hydrogen enriched product gas instead of being expanded at constant enthalpy through a valve, is expanded at constant entropy; that is, the gas is made to operate an expansion engine or move the rotor of a turbo-electric generator. However, after start-up, it may no longer be necessary to supply refrigerant to cooler 13 at a temperature of −78° F. Accordingly, the hydrogen enriched product gas may by-pass the expansion valve and is introduced into the cooler 13 at a temperature of about −55 to −65° F. This scheme is said to avoid the large pressure drop previously experienced across the expansion valve.
It has now been found that by pressurising the synthesis gas stream in a compression system to a pressure in the range of 150 to 400 barg and then passing the compressed synthesis gas feed stream through a heat exchanger system in heat exchange relationship with a plurality of internal refrigerant streams that are produced subsequently in the process, that the high pressure synthesis gas stream that exits the heat exchanger system may be cooled to a temperature in the range of −15 to −55° C. It has also been found that the cooled high pressure (HP) synthesis gas that exits the heat exchanger system may be separated in a gas-liquid separator vessel with negligible pressure drop across the heat exchanger system and separator vessel thereby forming a HP hydrogen rich vapour stream and a HP liquid carbon dioxide stream with the separation achieving 75% to 95% capture of the CO2 from the synthesis gas feed stream. It has also been found that the hydrogen rich vapour stream may be reduced in pressure to any desired pressure by passing the HP hydrogen enriched vapour stream through a turboexpansion system that comprises a plurality of turboexpanders arranged in series. In particular, the hydrogen rich vapour stream may be obtained at the desired fuel gas feed pressure for a combustor of a gas turbine of a power plant (for example, at a pressure of 30 barg). It has also been found that the expanded H2 rich vapour streams that exit each turboexpander of the series may be used as internal refrigerant streams thereby providing a major portion of the refrigeration duty for the heat exchanger system. Preferably, expansion of the H2 rich vapour in the turboexpanders may be used to drive a rotor or shaft of the compressor(s) of the compressor system or to drive the rotor or shaft of a turbo-electric generator thereby achieving a net power consumption for the separation of the synthesis gas stream into a hydrogen rich vapour stream and liquid CO2 stream of less than 30 MW, preferably, less than 25 MW when processing 28,000 kmol/hour of syngas containing 56 mol % hydrogen and 43 mol % CO2.
Thus, according to the present invention there is provided a process for separating a synthesis gas stream into a hydrogen (H2) rich vapour stream and a liquid carbon dioxide (CO2) stream in a CO2 condensation plant that comprises (a) a compression system comprising at least one compressor, (b) a heat exchanger system, (c) a gas-liquid separator vessel, and (d) a turboexpansion system comprising a plurality of turboexpanders arranged in series, the process comprising the steps of:
The machine that is driven by the turboexpanders in step (D) is preferably a compressor of the compression system and/or a pump (for example, for pumping liquid CO2 or supercritical CO2). Where the turboexpanders are used to drive an alternator of an electric generator, the electricity is preferably used to power one or more components of the CO2 condensation plant.
Preferably, the cooled HP synthesis gas stream formed in step (B) is passed to a cryogenic separation system comprising at least one cryogenic separation stage wherein the cryogenic separation stage(s) is comprised of a heat exchanger that employs an external refrigerant and a gas-liquid separator vessel. Accordingly, the gas-liquid separator vessel employed in step (C) of the present invention may be either the gas-liquid separator vessel of a single cryogenic separation stage that employs an external refrigerant or is the final gas-liquid separator vessel of a series of cryogenic separation stages wherein the cryogenic separation stages each employ an external refrigerant and are operated at progressively lower temperatures.
A further advantage of the present invention is that typically, at least 98%, preferably, at least 99%, more preferably, at least 99.5%, in particular, at least 99.8% of the hydrogen is recovered in the H2 rich vapour stream. Yet a further advantage of the process of the present invention is that the hydrogen rich vapour stream that is separated in step (C) is at a pressure substantially above the minimum fuel gas feed pressure (inlet pressure) for a combustor(s) of at least one gas turbine(s) of a power plant. Accordingly, the HP hydrogen rich vapour stream may be reduced in pressure in step (D) to the desired inlet pressure for the combustor(s) of the gas turbine(s) by isentropically expanding the HP hydrogen rich vapour stream in a series of turboexpanders thereby providing cold H2 rich vapours streams (internal refrigerant streams) that may be used to cool the HP synthesis gas stream in step (B). In addition, isentropic expansion of the hydrogen rich vapour streams in each of the turboexpanders of the series generates motive power that may be used to drive the compressor(s) of the compression system and/or to drive at least one alternator of an electric generator thereby generating electricity for use in the process (for example, for operating one or more electric compressors of the compression system) and/or to drive a pump (for example, a pump for a liquid CO2 or supercritical CO2 stream) Thus, a major portion of the compression energy may be recovered using the turboexpanders thereby increasing the overall energy efficiency of the process. However, it is recognised that there will be energy losses and that additional power will be required to run the compressors of the compression system. It is also recognised that the HP hydrogen rich vapour stream may be expanded to pressures below the inlet pressure of the combustor of a gas turbine, if the hydrogen rich vapour stream is to be used for a different purpose, for example, as fuel for a low pressure burner of a fired heater, or as fuel for a reformer or boiler or as a refinery feed stream for upgrading of one or more refinery streams or as a hydrogen feed to a chemical process.
The synthesis gas feed stream may be generated from a solid fuel such as petroleum coke or coal in a gasifier or from a gaseous hydrocarbon feedstock in a reformer. The synthesis gas from the gasifier or reformer may contain high amounts of carbon monoxide. Accordingly, depending on the desired composition of the hydrogen rich vapour stream, the synthesis gas may be treated in a shift converter unit where substantially all of the carbon monoxide contained in the synthesis gas stream is converted to carbon dioxide over a shift catalyst according to the water gas shift reaction (WGSR)
CO+H2O→CO2+H2
The shift converter unit may be a single shift reactor containing a shift catalyst. However, it is preferred that the shift converter unit comprises a high temperature shift reactor containing a high temperature shift catalyst and a low temperature shift reactor containing a low temperature shift catalyst. The water gas shift reaction is exothermic and results in a significant temperature rise across the shift converter unit. Accordingly, the shift converter unit may be cooled by continuously removing a portion of the shifted synthesis gas stream and cooling this stream by heat exchange with one or more process streams, for example against boiler feed water or against steam (for the generation of superheated steam). The synthesis gas that exits the shift converter unit comprises primarily hydrogen, carbon dioxide and steam and minor amounts of carbon monoxide and methane. Where the synthesis gas is of sufficiently high CO2 content, the shift conversion step may be omitted, in which case the synthesis gas comprises primarily hydrogen, carbon dioxide, carbon monoxide, and steam and minor amounts of methane.
Typically, the synthesis gas stream is cooled to a temperature in the range of 30 to 50° C., for example, about 40° C., upstream of the CO2 condensation plant, by heat exchange with at least one cold process stream, to condense out a condensate (predominantly comprised of water). Typically, the cold process stream is a process stream used during the generation of the synthesis gas. The condensate is then separated from the cooled synthesis gas stream, for example, in a condensate drum.
Where the synthesis gas stream is formed by gasification of petroleum coke or coal in a gasifier, the synthesis gas that exits the gasifier will also comprise minor amounts of hydrogen sulfide (H2S) as an impurity (sour synthesis gas). The H2S impurity is formed by reaction of COS with steam in the shift converter unit. This H2S may be captured upstream of the CO2 condensation plant, for example, by selectively absorbing the H2S from the sour synthesis gas in an absorption tower. Typically, Selexol™ (a mixture of dimethyl ethers of polyethylene glycol) may be employed as the absorbent. Any H2S that is captured may be either converted into elemental sulphur using the Claus Process or into industrial strength sulphuric acid. Alternatively, the sour synthesis gas stream may be fed to the CO2 condensation plant of the present invention where a major portion of the H2S partitions into the liquid CO2 phase and may therefore be sequestered with the CO2. Typically, greater than 95% of the H2S that was contained in the synthesis gas feed stream partitions into the liquid CO2 phase. Any residual H2S in the final H2 rich vapour stream may be removed downstream of the CO2 condensation plant by passing the final H2 rich vapour stream through an adsorbent bed, for example, a zinc oxide bed, or by passing the final H2 rich vapour stream through a scrubber that utilises a suitable liquid absorbent. There is minimal pressure drop, for example, a pressure drop of less than 0.5 bar across the absorbent bed.
After removal of any condensate (see above), the synthesis gas stream is preferably dried prior to being passed to the CO2 condensation plant, as any moisture in the synthesis gas will freeze and potentially cause blockages in the plant. The synthesis gas stream may be dried by being passed through a molecular sieve bed or an absorption tower that employs a solvent, for example, triethylene glycol, to selectively absorb the water. Preferably, the dried synthesis gas stream has a water content of less than 1 ppm (on a molar basis).
Typically, the dried synthesis gas comprises at least 40 mole % H2, preferably, at least 50 mole % H2, in particular, 55 to 60 mole % H2.
The synthesis gas stream is then fed to the compression system of the CO2 condensation plant at a pressure in the range 10 to 120 barg, preferably, 20 to 95 barg, more preferably, 40 to 75 barg, in particular, 40 to 60 barg. The temperature at which the synthesis gas is fed to the compression system of the CO2 condensation plant is not critical. However, it is preferred that the synthesis gas stream is fed to the compression system of the CO2 condensation plant at a temperature in the range of 25 to 50° C., for example, 30 to 40° C. The synthesis gas stream is then compressed, in the compression system, to a pressure in the range of 150 to 400 barg, preferably, 175 to 360 barg, more preferably, 250 to 360 barg, most preferably 300 to 360 barg, in particular, 330 to 360 barg. Preferably, the compression system is a multistage compressor system comprising a plurality of compressors arranged in series. However, where the synthesis gas stream is compressed to lower pressures, for example, pressures in the range of 150 to 200 barg, preferably, 150 to 180 barg, a single stage of compression may be employed. Thus, a multistage compression system is preferred for higher discharge pressures from the compression system and is optional for lower discharge pressures from the compression system. Generally, the compressor(s) of the compression system is mounted on a shaft that may be driven by an electric motor, gas turbine or steam turbine. Alternatively, the compressor(s) of the compression system and the turboexpanders of the turboexpansion system may be mounted on a common shaft so that the isentropic expansion of the hydrogen rich vapour in the turboexpanders may be used to drive the compressor(s).
A typical multistage compression system for use in the process of the present invention comprises at least one low pressure (LP) compressor, preferably two or three LP compressors mounted on a common drive shaft and at least one high pressure (HP) compressor, preferably one or two HP compressors mounted on a further common drive shaft (the drive shafts may be connected via a gear system). The LP and HP compressors are arranged in series. As would be well known to the person skilled in art, increased compression efficiency is achieved by balancing the compression duty across the compressors of the series. Thus, it is preferred that the compression ratios between successive compressors of the series are substantially the same.
Typically, the compressed HP synthesis gas stream is cooled to remove at least part, preferably, substantially all of the heat of compression before being passed through the heat exchanger system thereby reducing the cooling duty for the heat exchanger system. Where there is a single stage of compression, at least part of the heat of compression is removed from the HP synthesis gas by passing the HP synthesis gas stream through at least one heat exchanger of the compression system in heat exchange relationship with an external coolant and/or an external refrigerant. Where there are a plurality of stages of compression, it is preferred that the multistage compression system is provided with at least one interstage heat exchanger where the compressed gas is cooled against an external coolant and/or refrigerant before being passed to the next compressor in the series. Preferably, interstage heat exchangers are provided between each compressor in the series. The multistage compression system is also provided with at least one heat exchanger after the final stage of compression where the HP synthesis gas stream is cooled against an external coolant and/or an external refrigerant before being passed to the heat exchanger system. Typically, the compressed HP synthesis gas stream from the final stage of compression may be passed through a first heat exchanger in heat exchange relationship with an external coolant and then through a second heat exchanger in heat exchange relationship with an external refrigerant, prior to being passed to the heat exchanger system. Suitable external coolants for use in the heat exchanger(s) of the compression system include air, water, or a cold process stream such as the H2 rich vapour stream formed in step (C) or the final H2 rich vapour stream that is exported from the process of the present invention. Suitable external refrigerants for use in the heat exchanger of the compression system include propane, propenes and ammonia.
Where the HP synthesis gas that exits the compression system has not been heat exchanged with an external refrigerant, the HP synthesis gas typically exits the compression system at a temperature in the range of 20 to 50° C., in particular, 25 to 40° C. Where the HP synthesis gas that exits compression system has been heat exchanged with an external refrigerant, the HP synthesis gas typically exits the compression system at a temperature in the range of 0 to −30° C., for example, 0 to −10° C.
The HP synthesis gas stream is then passed through the heat exchanger system of the CO2 condensation plant where the HP synthesis gas stream is cooled against a plurality of internal refrigerant streams i.e. cold process streams that are produced subsequently in the process. The internal refrigerant streams may be selected from cold hydrogen rich vapour stream(s), in particular, cold expanded hydrogen rich vapour stream(s) from the turboexpanders of the turboexpansion system, and liquid CO2 stream(s). Typically, the HP synthesis gas stream is cooled in the heat exchanger system to a temperature in the range −15 to −55° C., preferably, −25 to −50° C., for example, −35 to −40° C. Typically, there is minimal pressure drop across the heat exchanger system, for example, a pressure drop of less than 1.5 bar, preferably, less than 1.0 bar.
Typically, the heat exchanger system comprises at least one multichannel heat exchanger with the HP synthesis gas stream being passed through a channel of the multichannel heat exchanger in heat exchange relationship with a plurality of internal refrigerant streams that are passed through further channels in the multichannel heat exchanger. Representative examples of the multichannel heat exchanger include those described in U.S. Pat. No. 6,622,519, WO 2004/016347, EP 212878 and EP 292245 the disclosures of which are incorporated herein by reference. As an alternative to pre-cooling the HP synthesis gas stream against an external refrigerant in a heat exchanger of the compression system, it is envisaged that one or more external refrigerant streams may be passed through yet further channels in the multichannel heat exchanger thereby providing additional cooling duty for the HP synthesis gas stream. Preferably, the HP synthesis gas stream is passed in a counter-current direction through the multichannel heat exchanger to the internal refrigerant streams and optional external refrigerant stream(s). Preferably, the heat exchanger system comprises a plurality of refrigeration stages arranged in series where each stage in the series comprises either (i) a single multichannel heat exchanger, or (ii) a plurality of multichannel heat exchangers arranged in parallel, for example, 2 or 3 multichannel heat exchangers arranged in parallel. Typically, the heat exchanger system comprises three refrigeration stages arranged in series with the internal refrigerant streams and optional external refrigerant stream(s) being fed to each successive stage of the series at successively lower temperatures. In a preferred heat exchanger system, a first refrigeration stage comprises two single-pass multichannel heat exchangers arranged in parallel, a second refrigeration stage comprises three 3-pass multichannel heat exchangers arranged in parallel, and a third refrigeration stage comprises a single 4-pass multichannel heat exchanger. Thus, the HP synthesis gas stream is divided and recombined as it passes through the stages of the heat exchanger system thereby optimising the heat exchange with the internal refrigerant streams and/or external refrigerant stream(s). However, alternative arrangements for a plurality of multichannel heat exchangers may be adopted.
Alternatively, the heat exchanger system may comprise a plurality of refrigeration stages wherein each refrigeration stage comprises either a single stand-alone heat exchanger or a plurality of stand-alone heat exchangers arranged in parallel. Thus, the HP synthesis gas stream is cooled as it is passed through the refrigeration stages of the heat exchanger system by heat exchange with a plurality of internal refrigerant streams and optional external refrigerant stream(s) that are fed to the stand-alone heat exchanger(s) of each successive refrigeration stage at successively lower temperatures. It is preferred that the HP synthesis gas stream is passed through the stand-alone heat exchangers in a counter-current direction to the internal refrigerant streams and optional external refrigerant stream(s) that are fed to the stand-alone heat exchangers.
It is also envisaged that the heat exchanger system may comprise both multichannel and stand-alone heat exchangers. Thus, the heat exchanger system may comprise a plurality of refrigeration stages arranged in series wherein each refrigeration stage comprises (i) a single multichannel heat exchanger, or (ii) a single stand-alone heat exchanger, or (iii) a plurality of multichannel heat exchangers and/or a plurality of stand-alone heat exchangers arranged in parallel.
The multichannel heat exchanger(s) of the heat exchanger system may be of the type employed in processes for generating liquefied natural gas such as a brazed aluminium plate-fin heat exchanger or a diffusion-bonded heat exchanger (for example, a printed circuit heat exchanger (PCHE) as supplied by Heatric). Alternatively, the multichannel heat exchanger(s) may be a multiple body shell and tube heat exchanger comprising either (a) a tube arranged in the shell of the heat exchanger wherein the shell of the heat exchanger comprises a plurality of compartments and wherein the HP synthesis gas stream is passed through the tube and an internal refrigerant stream or external refrigerant stream is passed through each compartment of the shell in heat exchange relationship with the HP synthesis gas that is flowing through the tube; or (b) a plurality of tubes arranged in the shell of the heat exchanger wherein the shell comprises a single compartment and the HP synthesis gas is passed through the compartment and an internal refrigerant stream or an external refrigerant stream is passed through each of the tubes in heat exchange relationship with the HP synthesis gas that is flowing through the single compartment of the shell. Accordingly, the term “channel” encompasses the channels formed between the plates of a brazed aluminium plate-fin heat exchanger or a diffusion-bonded heat exchanger and also the compartment(s) and tube(s) of a multiple body shell and tube heat exchanger.
The stand-alone heat exchanger(s) of the compression system may be of the shell and tube type (a single body shell and tube heat exchanger) with the HP synthesis gas stream passing through the tube side and an internal refrigerant stream or external refrigerant stream passing through the shell side of the heat exchanger or vice versa. However, a process that employs stand-alone heat exchangers to pre-cool the HP synthesis gas stream may be of reduced efficiency compared with a process that employs a multichannel heat exchanger, in whole or in part, to cool the HP synthesis gas stream in step (B) of the present invention.
The cooled HP synthesis gas stream that exits the heat exchanger system is a two phase stream comprised of a liquid phase and vapour phase. There is a limit on the temperature to which the HP synthesis gas stream may be cooled in the heat exchanger system as the temperature must be maintained above a value where solid CO2 will form. This typically occurs at a temperature of −56° C. (the triple point for pure CO2 is at 5.18 bar and at a temperature of −56.4° C.) although the presence of H2 may depress this freezing point. The amount of cooling that is achieved in the heat exchanger system owing to heat exchange with the plurality of internal refrigerant streams will be dependent upon the amount of cooling of the isentropically expanded hydrogen rich vapour streams that is achieved in the turboexpansion system which, in turn, is dependent on the pressure of the HP hydrogen rich vapour stream that is formed in step (C) and the pressure of the H2 rich vapour stream that exits the final turboexpander of the turboexpansion system in step (D). The amount of electricity generated by the turboexpanders of the turboexpansion system will also be dependent on the extent to which the hydrogen rich vapour is subjected to isentropic expansion in the turboexpansion system which is also dependent on the pressure of the H2 rich vapour stream formed in step (C) and the pressure of the H2 rich vapour stream that exits the final turboexpander of the turboexpansion system in step (D).
The cooled synthesis gas stream from the heat exchanger system may be passed directly to a gas-liquid separator vessel that is preferably operated at substantially the same pressure as the heat exchanger system. Thus, the pressure drop across the separator vessel is typically in the range of 0.1 to 5 bar, preferably, 0.1 to 1 bar, in particular, 0.1 to 0.5 bar. Accordingly, a HP hydrogen rich vapour phase is withdrawn from at or near the top of the gas-liquid separator vessel and is passed to the turboexpander system while a HP liquid CO2 stream is withdrawn from at or near the bottom of the gas-liquid separator vessel. If necessary, the HP liquid CO2 stream is reduced in pressure to the desired export pressure for the liquid CO2 product stream (as described below).
Where relatively high carbon capture levels are desired, for example, greater than 90%, preferably, greater than 95% CO2 capture from the synthesis gas feed stream (based on the molar composition of the synthesis gas feed stream), and insufficient cooling of the HP synthesis gas has been achieved in the compression system and heat exchanger system, the cooled HP synthesis gas that exits the heat exchanger system may be passed to a cryogenic separation system of the CO2 condensation plant wherein the cryogenic separation system comprises at least one cryogenic separation stage. The cryogenic separation stage(s) comprises a heat exchanger that employs an external refrigerant and a gas-liquid separator vessel. The cryogenic separation system is operated without any substantial reduction in pressure across the cryogenic separation stage(s). Where the cryogenic separation system comprises a single cryogenic separation stage, a HP hydrogen rich vapour stream and a HP liquid CO2 stream are formed by: (i) passing the cooled HP synthesis gas that exits the heat exchanger system through the heat exchanger of the single cryogenic separation stage where it is cooled against an external refrigerant, (ii) passing the resulting cooled HP stream to the separator vessel of the single cryogenic separation stage where a H2 rich vapour phase separates from a liquid CO2 phase, and (iii) withdrawing a HP hydrogen rich vapour stream and a HP liquid CO2 stream from the separator vessel of the single cryogenic separation stage. Where the desired carbon dioxide capture level cannot be achieved using a single cryogenic separation stage, a HP hydrogen rich vapour stream and a plurality of HP liquid CO2 streams are formed by: (i) passing the cooled HP synthesis gas that exits the heat exchanger system through the first heat exchanger of a multistage cryogenic separation system where it is cooled against an external refrigerant, (ii) passing the resulting cooled HP stream to the separator vessel of the first cryogenic separation stage where a H2 rich vapour phase separates from a liquid CO2 phase, (iii) withdrawing a HP hydrogen rich vapour stream and a HP liquid CO2 stream from the separator vessel of the first cryogenic separation stage, (iv) passing the HP hydrogen rich vapour stream through the heat exchanger of a further cryogenic separation stage where it is cooled against a further external refrigerant to below its dew point thereby forming a cooled two phase stream, (v) passing the cooled two phase stream to the separator vessel of the further cryogenic separation stage where a hydrogen rich vapour phase separates from a liquid CO2 phase, and (vi) withdrawing a HP hydrogen rich vapour stream and a HP liquid CO2 stream from the separator vessel of the further separation stage and (vii) if necessary, repeating steps (iv) to (vi) by passing the HP hydrogen rich vapour stream through one or more further cryogenic separation stages until the desired CO2 capture has been achieved. Accordingly, the HP hydrogen rich vapour stream is removed from the gas-liquid separation vessel of the final cryogenic separation stage of the series.
Where the HP synthesis gas stream is cooled in the heat exchanger system to a temperature in the range of −30 to −40° C., for example, about −37° C., the cryogenic separation system may comprise a single cryogenic separation stage having an operating temperature in the range of −40 to −55° C., preferably, −45 to −50° C. Where the HP synthesis gas stream exits the heat exchanger system is at a higher temperature, for example, a temperature in the range of −15 to greater than −30° C., the cryogenic separation system may comprise a plurality of cryogenic separation stages that are arranged in series with the separator vessels of the cryogenic separation stages operated at successively lower temperatures. The operating temperature of each cryogenic separation stage will depend on the temperature to which the HP synthesis gas stream has been cooled in the heat exchanger system, the number of cryogenic separation stages and the desired carbon dioxide capture level. There is a limit on the lowest temperature in the final cryogenic separation stage, as the temperature must be maintained above a value where solid CO2 will form (see above). Generally, the final cryogenic separation stage is operated at a temperature in the range of −40 to −55° C., preferably, −45 to −50° C.
Suitable external refrigerants that may be used as refrigerant in the heat exchanger(s) of the cryogenic separation stages(s) include propanes, ethane, ethylene, ammonia, hydrochlorofluorocarbons (HCFC's) and mixed refrigerants. Typical mixed refrigerants comprise at least two refrigerants selected from the group consisting of butanes, propanes, ethane, and ethylene. These refrigerants may be cooled to the desired refrigeration temperature in external refrigerant circuits using any method known to the person skilled in the art including methods known in the production of liquefied natural gas. These refrigerants may also be cooled to the desired refrigeration temperature by heat exchange with one or more cold isentropically expanded H2 rich vapour streams from the turboexpanders of the turboexpansion system. The external refrigerant for the cryogenic separation stage is selected so as to achieve the desired operating temperature. For example, propane may be used as refrigerant when the feed temperature of the HP synthesis gas stream is in the range of −15 to greater than −30° C. and the desired operating temperature of the cryogenic separation stage is in the range of −20 to greater than −30° C. while ethane and/or ethylene may be used as external refrigerant when the feed temperature of the HP synthesis gas stream is in the range of −30 to −40° C. and the desired operating temperature for the cryogenic separation stage is in the range of −40 to −55° C., preferably, −45 to −50° C.
As discussed above, the cryogenic separation system is operated with minimal pressure drop across the cryogenic separation stage or stages of the system. Typically, the pressure drop across the cryogenic separation system is in the range of 1 to 5 bar, preferably, 1 to 3 bar, in particular, 1 to 2 bar. Thus, where there are a plurality of cryogenic separation stages, these may be operated at substantially the same pressure. However, higher pressure drops across the cryogenic separation system may be tolerated, for example, pressure drops in the range of 5 to 20 bar, preferably 5 to 10 bar.
Preferably, prior to being fed to the turboexpansion system, the HP hydrogen rich vapour stream formed in step (C) is used to cool the HP synthesis gas stream in step (B) by passing the cold HP hydrogen rich vapour stream through a further channel in a multichannel heat exchanger (or through a stand-alone heat exchanger) in heat exchange relationship with the HP synthesis gas stream. Alternatively the HP hydrogen rich vapour stream may be used as coolant for the heat exchanger(s) of the compression system.
In a another preferred embodiment the HP H2 rich vapour stream is first fed to a solvent extraction system in which the vapour stream is contacted with a solvent which absorbs any residual CO2 contained therein. Solvent extraction processes for effecting this separation include the Rectisol™ and Selexol™ processes which respectively uses refrigerated methanol and a refrigerated mixture of dimethyl ethers of polyethylene glycol as the absorbent. Alternatively the absorbent can be amine based for example monoethanolamine, diethanolamine, methyldiethanolamine diisopropylamine or the like.
The HP H2 rich vapour stream that is fed to the turboexpansion system is at elevated pressure. Accordingly, the H2 rich vapour stream is reduced in pressure to the desired exit pressure by being passed through the turboexpansion system. The expansion energy recovered from the H2 rich vapour streams in the turboexpanders can be used to drive an electric turbine or can be used to directly drive the compressors of the compressor system. As discussed above, isentropic expansion of the H2 rich vapour stream results in significant cooling. Where one or more, preferably all, of the cold isentropically expanded hydrogen rich vapour streams that are withdrawn from the turboexpanders of the turboexpansion system are used to cool the HP synthesis gas stream in step (B), the operating pressures of the turboexpanders are set to optimise the expander efficiency and to ensure the discharge temperatures for the expanded H2 rich vapour streams do not fall below −56° C. (the freezing point of CO2). Typically, the discharge temperatures of the expanded streams are in the range of −15° C. to −50° C., preferably, −20° C. to −50° C., in particular, −30 to −50° C. However, it is also envisaged that one or more of the cold isentropically expanded hydrogen rich vapour streams may be used as an internal refrigerant for a different purpose, for example, to cool an external refrigerant (such as propane or ethane) that is employed in the process or to cool a non-isentropically expanded H2 rich vapour stream so that it can be used as internal refrigerant in the process of the present invention. As there is no risk of forming solidified CO2 when an internal refrigerant is used for these purposes, the discharge temperature of the expanded stream may be below −55° C.
Where the HP synthesis gas stream that exits the heat exchanger system is either passed directly to a gas-liquid separator vessel or is passed to a cryogenic separation system that comprises a single cryogenic separation stage, a single liquid CO2 stream is obtained. Where the HP synthesis gas stream that exits the heat exchanger system is passed to a cryogenic separation system that employs a plurality of cryogenic separation stages, a separate liquid CO2 stream will be obtained from each cryogenic separation stage. Preferably, the liquid CO2 streams from the plurality of cryogenic separation stages are combined to form a combined liquid CO2 stream. Where the liquid CO2 stream or combined liquid CO2 stream is obtained at the desired CO2 export pressure (typically, 100 to 175 barg), the liquid CO2 stream or combined liquid CO2 stream is preferably used as internal refrigerant in the heat exchanger system thereby providing further cooling for the HP synthesis gas feed stream in step (B). The CO2 product stream that exits the heat exchanger system is in a dense phase state. By “dense phase state” is meant that the CO2 product stream is either in a supercritical fluid state or a liquid state depending upon whether the temperature of the CO2 product stream is above or below its critical temperature. The dense phase CO2 product stream is then exported from the process and sequestered and/or used in a chemical process.
Where the liquid CO2 stream or the combined liquid CO2 stream is produced at a pressure above the desired liquid CO2 export pressure, the liquid CO2 stream or the combined liquid CO2 stream is reduced in pressure to the desired liquid CO2 export pressure, typically, a pressure in the range of 100 to 175 barg. It is envisaged that the energy associated with reducing the pressure of the liquid CO2 stream or the combined liquid CO2 stream to export pressure may be recovered using a hydro turbine system. Alternatively, the liquid CO2 stream or the combined liquid CO2 stream may be flashed to lower pressure across a valve before being passed to a flash separation vessel (the valve may be located at the inlet to the flash separation vessel). Any H2 rich vapour that separates from the liquid CO2 stream owing to the reduction in pressure may be removed as a H2 rich vapour stream from at or near the top of the flash separator vessel. The H2 rich vapour stream that is withdrawn from the flash separator vessel may be combined with a similar pressure expanded H2 rich vapour stream from one of the turboexpanders of the turboexpander system and the combined H2 rich vapour stream is then used as an internal refrigerant stream in the heat exchanger system thereby providing cooling of the HP synthesis gas stream in step (B). For example, the combined H2 rich vapour stream may be routed through a channel of a multichannel heat exchanger in heat exchange relationship with the HP synthesis gas stream. Alternatively, when further recovery of CO2 from the H2 rich vapour stream is desired, the H2 rich vapour stream that is withdrawn from the flash separator vessel may be combined with a similar pressure compressor stream that exits one of the compressors of the compressor system. Preferably, the liquid CO2 stream that is withdrawn from the flash separator vessel is used as an internal refrigerant in the heat exchanger system thereby providing further cooling for the HP synthesis gas stream in step (B). As described above, a dense phase CO2 product stream (either supercritical CO2 or liquid CO2) exits the heat exchanger system. This dense phase CO2 product stream is then exported from the process and sequestered and/or used in a chemical process. Thus, an advantage of the present invention is that the dense phase CO2 product stream is produced at the desired CO2 export pressure so that there is no requirement to provide additional CO2 pumps.
The dense phase CO2 product stream that is exported from the process of the present invention preferably comprises at least 95 mole % CO2, in particular, at least 98 mole % CO2, the remainder being mostly hydrogen with some inerts, for example, nitrogen and/or CO. However, where the synthesis gas feed stream is a sour synthesis gas, the dense phase CO2 product stream will also comprise minor amounts of H2S. Where the dense phase CO2 product stream comprises H2S, the product is either sequestered or the H2S is removed prior to the CO2 being used for a different purpose, such as a feed to a chemical process. Where the dense phase CO2 product stream is sequestered, it is typically delivered to a pipeline that transfers the dense phase CO2 product stream to a reception facility of an oil field where the dense phase CO2 product stream may be used as an injection fluid for an oil reservoir. If necessary, the dense phase CO2 product stream is pumped to above the pressure of the oil reservoir before being injected down an injection well and into the oil reservoir. The injected CO2 displaces the hydrocarbons contained in the reservoir rock towards a production well for enhanced recovery of hydrocarbons therefrom. If any carbon dioxide is produced from the production well together with the hydrocarbons, the carbon dioxide may be separated from the hydrocarbons for re-injection into the oil reservoir. It is also envisaged that the dense phase CO2 product stream may be sequestered by being injected into an aquifer or a depleted oil or gas reservoir for storage therein.
Typically, the amount of CO2 contained in the H2 rich vapour stream that is obtained from the CO2 condensation plant is less than 10 mole %, preferably less than 5 mole %, more preferably, less than 2 mole %, in particular, less than 1 mole %.
Where the synthesis gas feed stream has not been subjected to a shift conversion, the H2 rich vapour stream may comprise up to 20 mole % carbon monoxide (CO). Where the synthesis gas feed stream has been subjected to a shift conversion, the H2 rich vapour stream typically comprises trace amounts of CO and methane, for example, less than 500 ppm on a molar basis. The remainder of the hydrogen rich vapour stream that is obtained from the CO2 condensation plant is predominantly hydrogen.
As discussed above, the H2 rich vapour stream obtained from the CO2 condensation plant may be used as fuel for a low pressure burner of a fired heater, or as fuel for a reformer or boiler or as a refinery feed stream for upgrading of one or more refinery streams or as a feed to a chemical process. However, it is preferred to use the H2 rich vapour stream as a fuel gas stream for the combustor of at least one gas turbine of a power plant thereby producing electricity. Where the H2 rich vapour stream is to be used as a fuel, it is preferred that the H2 rich vapour stream contains only trace amounts of CO thereby minimizing the amount of CO2 released to atmosphere (where the CO2 arises from combustion of CO).
As discussed above, an advantage of the present invention is that the fuel gas stream may be obtained at above the minimum inlet pressure for the combustor(s) of the gas turbine(s). Typically, the feed pressure for the fuel gas stream (inlet pressure for the combustor of the gas turbine(s)) is in the range of 25 to 45 barg, preferably, 28 to 40 barg, in plute. Accordingly, there is no requirement for a compressor to compress the fuel gas stream to the inlet pressure for the combustor(s) of the gas turbine(s). Preferably, the H2 rich vapour stream is diluted with medium pressure nitrogen and/or medium pressure steam prior to being fed as fuel gas to the combustor(s) of the gas turbine(s). Accordingly, the fuel gas stream that is fed to the combustor(s) of the gas turbine(s) preferably contains 35 to 65 mole % hydrogen, more preferably, 45 to 60 mole % hydrogen, for example, 48 to 52 mole % of hydrogen.
It has been found that the expansion energy recovered from the H2 rich vapour streams in the turboexpanders of the turboexpansion system (when the fuel gas stream is obtained at a pressure of about 30 barg) can reduce the net power consumption of the separation process to less than 30 MW, preferably, less than 25 MW, in particular less than 23 MW when processing 28,000 kmol/hour of synthesis gas containing 56 mol % hydrogen and 43 mol % CO2. The net power consumption is defined as: [power consumption in the compression system+power consumption in the external refrigerant compression system (if used)]−[power generated in the turboexpansion system].
The exhaust gas from the gas turbine(s) of the power plant is passed to a heat recovery and steam generator unit (HRSG) where the exhaust gas may be heat exchanged with various process streams. Optionally, the temperature of the exhaust gas of the gas turbine is increased by providing the HRSG with a post-firing system, for example, a post-firing burner. Suitably, the post-firing burner is fed with a portion of the fuel gas stream wherein the fuel gas stream is combusted in the burner using residual oxygen contained in the exhaust gas. Suitably, the exhaust gas is raised in temperature in the post-firing system to a temperature in the range of 500 to 800° C.
Typically, the HRSG generates and superheats steam for use in at least one steam turbine and/or in a process for generating the synthesis gas and/or for diluting the fuel gas stream. Typically, the HRSG is capable of generating high pressure (HP) steam, medium pressure (MP) steam and low pressure (LP) steam and of superheating these steam streams. The HRSG may also be capable of reheating MP steam that is produced as an exhaust stream from the high pressure stage of a multistage steam turbine. In addition, the HRSG may be used to heat boiler feed water (for example, boiler feed water that is fed to the waste heat boiler of a shift converter unit).
The cooled exhaust gas is discharged from the HRSG to the atmosphere through a stack. Preferably, the stack is provided with a continuous emission monitoring system for monitoring, for example, the NOx content of the cooled exhaust gas.
In a further embodiment of the present invention there is provided a carbon dioxide condensation plant for separating carbon dioxide and a hydrogen from a synthesis gas stream, the plant comprising:
As discussed above, the CO2 condensation plant may optionally comprise a valve for letting down the pressure of the separated liquid CO2 stream to the CO2 export pressure. Accordingly, the CO2 condensation plant optionally comprises a flash separation vessel for separating any hydrogen rich vapour from the reduced pressure liquid CO2 stream.
Also, as discussed above, the CO2 condensation plant may optionally comprise a cryogenic separation system comprising at least one cryogenic separation stage that is comprised of a heat exchanger that employs an external refrigerant and a gas-liquid separator vessel with the cryogenic separation system being operated with minimal pressure drop across the cryogenic separation stage or stages of the system. Accordingly, the gas-liquid separator vessel for separating the partially condensed compressed synthesis gas stream into a hydrogen rich vapour stream and liquid CO2 stream is either the gas-liquid separator vessel of a single cryogenic separation stage or is the final gas-liquid separator vessel of a plurality of cryogenic separator vessels that are arranged in series. Where the cryogenic separation system comprises a plurality of separation stages, a liquid CO2 stream will be withdrawn from, the final cryogenic separation stages and additional liquid CO2 streams will be withdrawn from each of the gas-liquid separator vessels of the preceding cryogenic separation stages of the series.
The process of the present invention will now be illustrated by reference to
The low pressure synthesis gas stream 1D that exits the first LP compressor C1 is at a pressure of 94 bar and a temperature of 84.62° C., the increase in temperature arising from heat of compression. Stream 1D is then cooled in heat exchanger E1 against a cold stream (for example, water or air) and is passed to the second LP compressor C2 thereby generating a second LP synthesis gas stream 2D having a pressure of 147.0 bar and a temperature of 84.21° C. Stream 2D is then cooled in heat exchanger E2 against a cold stream (for example, water or air) and is passed to the first high pressure (HP) compressor C3 thereby generating a first high pressure stream 3D having a pressure of 230.0 barg and a temperature of 83.56° C. Stream 3D is then cooled in heat exchanger E3 against a cold stream (for example, water or air) and is passed to the second HP compressor C4 thereby forming a second HP synthesis gas stream 4D having a pressure of 360.0 barg and a temperature of 82.25° C. Stream 4D is then cooled in heat exchanger E4 against a cold stream (for example, water or air) thereby generating HP synthesis gas stream S1 having a temperature of about 40° C. The person skilled in the art would understand that the number of compression stages has been optimised to minimise power consumption and that compression from 60 to 360 barg could also have been achieved using two or three compression stages. The person skilled in the art would also understand that the pressure of the synthesis gas exiting each compression stage can be varied. However, the discharge pressure from the final stage of compression determines the amount of carbon dioxide that is captured from the synthesis gas stream. The person skilled in the art will also understand that the HP synthesis gas may be cooled in heat exchanger E4 against an external refrigerant, for example, propane, propylene or ammonia, to a temperature in the range of 0 to −30° C. thereby reducing the cooling duty for the multichannel heat exchanger LNG-100.
HP synthesis gas stream S1 is then passed through multichannel heat exchanger LNG-100 where it is cooled against a plurality of cold process stream (see below) thereby generating a cooled HP synthesis gas stream S2 having a pressure of 358.5 bar and a temperature of −21.33° C. (where stream S1 has a temperature of about 40° C.).
Stream S2 is passed directly to gas-liquid separator vessel F360 where a hydrogen rich vapour phase separates from a liquid CO2 phase. Accordingly, a hydrogen rich vapour stream S2V is removed overhead from separator vessel F360 and is passed through multichannel heat exchanger LNG-100 thereby cooling the HP synthesis gas stream S1. H2 rich vapour stream 1N exits the multichannel heat exchanger LNG-100 and is then routed to expander EX1 where it is expanded to lower pressure. The person skilled in the art would understand that isentropic expansion of the vapour stream results in cooling. Accordingly, stream 1T exits the expander EX1 at a pressure of 205.0 barg and a temperature of −26.25° C. and is routed through multichannel heat exchanger LNG-100 where it is heat exchanged with HP synthesis gas stream S1 thereby cooling the HP synthesis gas stream S1. H2 rich vapour stream 2N exits the multichannel heat exchanger LNG-100 and is passed to expander EX2 where the vapour stream is expanded to lower pressure. H2 rich vapour stream 2T exits the expander at a pressure of 112.0 barg and a temperature of −26.01° C. and is passed to manifold M1. A liquid CO2 stream S2L is withdrawn from the bottom of separator vessel F360 and is flashed across valve VLV-109 thereby generating a further two phase stream 18 that is passed to flash vessel F150. H2 rich vapour stream S2LV that is withdrawn from the top of flash vessel F150 is routed to manifold M1 where is it combined with H2 rich vapour stream 2T to form combined vapour stream 2TM. Vapour stream 2TM is then passed through multichannel heat exchanger LNG-100 thereby cooling HP synthesis gas stream S1. H2 rich vapour stream 3N that exits the multichannel heat exchanger LNG-100 is then passed to expander EX3 where it is expanded to lower pressure. Vapour stream 3T exits expander EX3 at a pressure of 65.0 barg and a temperature of −24.56° C. and is passed through multichannel heat exchanger LNG-100 thereby cooling HP synthesis gas stream S1. H2 rich vapour stream 4N is then passed to expander EX4 where it is expanded to lower pressure. H2 rich vapour stream 4T exits expander EX4 at a pressure of 31.0 barg and a temperature of −24.59° C. and is passed through multichannel heat exchanger LNG-100 thereby cooling HP synthesis gas stream S1. The final H2 rich vapour stream that exits the multichannel heat exchanger LNG-100 has a pressure of 30.5 barg and a temperature of 37.00° C. This vapour stream comprises 88.42 mole % H2 and 9.30 mole % CO2 and, after dilution with medium pressure N2 and/or medium pressure steam, may be passed as hydrogen fuel gas to the combustor(s) of the gas turbine(s) of a Power Island (not shown). The person skilled in the art will understand that the number of expanders can be increased or decreased (minimum of two expanders). The person skilled in the art will also understand that the operating pressure and temperature of the expanders can also be varied provided that the vapour stream that exits the multichannel heat exchanger has a pressure of at least 30 barg. The expanders EX1, EX2, EX3 and EX4 may be connected to electric motors to recover energy and the electricity may be either used in the process or is exported from the process. Alternatively, the expanders may be directly coupled to the compressors, for example, by mounting the expanders and compressors on a common shaft so that the isentropic expansion of the hydrogen rich vapour in the expanders is used to turn the common shaft and to drive the compressors. Accordingly, the net power consumption for the flow scheme of
A liquid CO2 stream S2LL is withdrawn from the bottom of flash vessel F_150 and is passed through multichannel heat exchanger LNG-100 where it is used to cool HP synthesis gas stream S1. A dense phase CO2 stream exits the multichannel heat exchanger LNG-100 at a temperature of 24° C. and a pressure of 149.5 barg. Accordingly, the dense phase CO2 stream is above the minimum pipeline export pressure (typically, 130 barg). The dense phase CO2 stream that exits the multichannel heat exchanger comprises 96.52 mole % CO2 and 2.69 mole % hydrogen. The hydrogen recovery is 98.15% based on the moles of H2 in the dry synthesis gas stream. The amount of carbon dioxide captured from the dry synthesis gas stream is 84.58% based on the moles of CO2 in the dry synthesis gas stream
Stream S2 is then cooled to a temperature of −50° C. against ethane refrigerant that is fed to the shell side of kettle E-101 of a refrigeration circuit (discussed below) thereby generating a two phase stream 20 that is passed to separator vessel F360 where a hydrogen rich vapour phase separates from a liquid CO2 phase. A H2 rich vapour stream S2V is removed overhead from the separator vessel F360 and is passed through multichannel heat exchanger LNG-100 where it is heated against HP synthesis gas stream S1 thereby forming vapour stream 1N (and cooling the HP synthesis gas stream S1). H2 rich vapour stream 1N is then routed to expander EX1 where it is expanded to lower pressure. H2 rich vapour stream 1T exits the expander EX1 at a pressure of 205.0 barg and a temperature of −31.63° C. and is routed through multichannel heat exchanger LNG-100 where it is heat exchanged with HP synthesis gas stream S1 thereby forming vapour stream 2N (and cooling the HP synthesis gas stream S1). H2 rich vapour stream 2N is passed to expander EX2 where the vapour stream is expanded to lower pressure. H2 rich vapour stream 2T exits the expander at a pressure of 112.0 barg and a temperature of −35.04° C. and is passed to manifold M1. A liquid CO2 stream S2L is withdrawn from the bottom of separator vessel F_360 and is flashed across valve VLV-109 thereby generating a further two phase stream 18 that is passed to flash vessel F150. H2 rich vapour stream S2LV that is withdrawn from the top of flash vessel F150 is routed to manifold M1 where is it combined with H2 rich vapour stream 2T to form combined vapour stream 2TM. The combined H2 rich vapour stream 2TM is then passed through multichannel heat exchanger LNG-100 where it is heated against HP synthesis gas stream S1 thereby forming vapour stream 3N (and cooling the HP synthesis gas stream S1). H2 rich vapour stream 3N is passed to expander EX3 where it is expanded to lower pressure. H2 rich vapour stream 3T exits expander EX3 at a pressure of 65.0 barg and a temperature of −42.57° C. and is passed through multichannel heat exchanger LNG-100 where it is heated against HP synthesis gas stream Si thereby forming vapour stream 4N (and cooling the HP synthesis gas stream S1). H2 rich vapour stream 4N is then passed to expander EX4 where it is expanded to lower pressure. H2 rich vapour stream 4T exits expander EX4 at a pressure of 31.0 barg and a temperature of −53.81° C. and is heated against HP synthesis gas stream S1 in multichannel heat exchanger LNG-100 (thereby cooling HP synthesis gas stream S1). The final H2 rich vapour stream that exits the multichannel heat exchanger has a pressure of 30.5 barg and a temperature of 37° C. and, after being diluted with MP N2 and/or steam, is passed as hydrogen fuel gas to the combustor(s) of the gas turbine(s) of a Power Island (not shown). An advantage of the flow scheme of
The person skilled in the art will understand that the number of expanders can be increased or decreased (minimum of two expanders). The person skilled in the art will also understand that the operating pressure and temperature of the expanders can also be varied provided that the vapour stream that exits the multichannel heat exchanger has a pressure of at least 30 barg. The expanders EX1, EX2, EX3 and EX4 are each connected to electric motors to recover energy and the electricity may be either used in the process or is exported from the process. Alternatively, the expanders may be directly coupled to the compressors, for example, by mounting the expanders and compressors on a common shaft so that the isentropic expansion of the hydrogen rich vapour in the expanders is used to turn the common shaft and to drive the compressors. Accordingly, the net power consumption for the flow scheme of
A liquid CO2 stream S2LL is withdrawn from the bottom of separator vessel F_150 and is passed through multichannel heat exchanger LNG-100 where it is used to cool HP synthesis gas stream S1. The dense phase CO2 stream exits the multichannel heat exchanger LNG-100 at a temperature of 24° C. and a pressure of 149.5 barg. Accordingly, the dense phase CO2 stream is above the minimum pipeline export pressure (typically, 130 barg).
The dense phase CO2 stream that exits the multichannel heat exchanger comprises 97.87 mole % CO2 and 1.36 mole % hydrogen. The final H2 rich vapour stream that exits the multichannel heat exchanger LNG-100 comprises 94.38 mole % H2 and 3.24 mole % CO2. The hydrogen recovery is 98.8% based on the moles of H2 in the dry synthesis gas stream. The amount of carbon dioxide captured from the dry synthesis gas stream is 92.5% based on the moles of CO2 in the dry synthesis gas stream.
The ethane refrigerant that is fed to the shell side of kettle E-101 of the refrigeration circuit may be obtained using any cryogenic method known to the person skilled in the art, including cryogenic methods for producing refrigerants for liquefying natural gas. The ethane refrigerant for heat exchanger E-101 may be replaced with ethylene.
Typically, the Ethane refrigerant that is employed in Kettle E-101 of refrigeration circuit is compressed in Ethane Compressor K-100. Ethane vapour stream 4 from the discharge of Ethane Compressor K-100 is fully condensed against propane refrigerant in Ethane Condenser E-102 against propane refrigerant. The liquefied ethane stream 6 exiting E-102 is then routed to kettle E-101 with ethane flow to kettle E-101 being controlled by means of an inlet level control valve VLV-101. The ethane vapour stream 2 exiting the E-101 kettle is at a temperature of −51.59° C. and a pressure of 5 bar and is routed to Ethane Compressor K-100 via valve VLV-100.
The propane refrigerant in the CO2 Condensation Circuit is compressed in Propane Compressor K-100-2. Propane vapour stream 4-2 from the compressor K-100-2 discharge is fully condensed in air cooled Condenser E-102-2. The liquefied propane 5-2 exiting Condenser E-102-2 routed to kettle E-101-2 that is operated at a temperature of −32.62° C. and a pressure of 1.5 barg. The propane flow to kettle E-101-2 is controlled by means of an inlet level control valve VLV-101-2. Vapour stream 2-2 exiting the kettle E-101-2 is routed to the propane compressor K-100-2 via valve VLV 100-2.
It is also envisaged that the CO2 condensation plant of
Stream S2 is cooled to a temperature of −50° C. against ethane refrigerant that is fed to the shell side of kettle E-101 of a refrigeration circuit (the refrigeration circuit being as described for
The person skilled in the art will understand that the number of expanders can be increased or decreased (minimum of two expanders). The person skilled in the art will also understand that the operating pressure and temperature of the expanders can also be varied provided that the vapour stream that exits the multichannel heat exchanger has a pressure of at least 30 barg. The expanders EX1, EX2, EX3 and EX4 are each connected to electric motors to recover energy and the electricity may be either used in the process or is exported from the process. Alternatively, the expanders may be directly coupled to the compressors, for example, by mounting the expanders and compressors on a common shaft so that the isentropic expansion of the hydrogen rich vapour in the expanders is used to turn the common shaft and to drive the compressors. Accordingly, the net power consumption for the flow scheme of
A liquid CO2 stream S2LL is withdrawn from the bottom of separator vessel F150 and is passed through multichannel heat exchanger LNG-100 where it is used to cool HP synthesis gas stream S1. The dense phase CO2 stream exits the multichannel heat exchanger LNG-100 at a temperature of 24° C. and a pressure of 149.5 barg. Accordingly, the dense phase CO2 stream is above the minimum pipeline export pressure (typically, 130 barg). If desired, the liquid CO2 stream may be boosted in pressure, using a CO2 pump (not shown) to a pressure of up to 200 barg.
The dense phase CO2 stream that exits the multichannel heat exchanger comprises 98.28 mole % CO2 and 1.19 mole % hydrogen. The final H2 rich vapour stream that exits the multichannel heat exchanger LNG-100 comprises 92.06 mole % H2 and 5.44 mole % CO2. The hydrogen recovery is 99.14% based on the moles of H2 in the dry synthesis gas feed stream. The amount of carbon dioxide captured from the dry synthesis gas stream is 90.08% based on the moles of CO2 in the dry synthesis gas feed stream.
| Number | Date | Country | Kind |
|---|---|---|---|
| 08252610.4 | Jul 2008 | EP | regional |
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/GB2009/001810 | 7/24/2009 | WO | 00 | 5/16/2011 |
