Patent Application
20120118012
The present invention relates to a process for separating gases from a mixture thereof. In particular it relates to the separation of a first and relatively more condensable gas from a mixture in which it is mixed with one or more relatively less condensable second gases. In its most specific form the invention relates to a two-stage process for separating one or more relatively more condensable gases selected from the group consisting of carbon dioxide, hydrogen sulphide and carbonyl sulphide from mixtures comprising one or more of these components and one or more relatively non-condensable gases selected from the group consisting of carbon monoxide, methane, ethane, hydrogen, nitrogen oxygen and synthesis gas.
For environmental reasons it is becoming increasingly desirable to separate greenhouse and polluting gases from those gases which are environmentally benign, such as hydrogen, nitrogen or oxygen, or to purify gases in order to make them fit for other duties (for example carbon monoxide, methane, ethane or natural gas sources). In particular increasing attention is being given to strategies in which the greenhouse gas, carbon dioxide, the principle carbon containing product of the combustion of hydrocarbon fuels, is separated and stored underground in rock formations. However there exists other situations where it is desirable to separate gases on a large scale for example the need to sweeten sour natural gas which contains significant levels of hydrogen sulphide. In such sweetening processes the strategy it to isolate a relatively pure stream of hydrogen sulphide which can then be converted into elemental sulphur, for example by the Claus process.
In our co-pending application PCT/GB/2009/001810 there is described a process for separating relatively condensable carbon dioxide from gas mixtures wherein the other major component is principally non-condensable hydrogen. The process described comprises in general terms first compressing and cooling the mixture to a pressure and temperature at which carbon dioxide is liquid and thereafter separating the liquid carbon dioxide condensed out from the other non-condensable gases. Thereafter the separated components are returned to a temperature and pressure suitable for further use by a series of heat exchangers and turbo expanders integrated amongst themselves and with those used to cool the feedstock so that the total energy across the whole process is managed for optimum efficiency. Our application describes process configurations for achieving this outcome and in particular the use of compact, diffusion-bonded heat exchangers to simplify the demands on the hardware needed.
Whilst the process described in this application allows good separation of carbon dioxide from hydrogen-rich gases, it has now been found that its efficiency and product purity can be further improved if after separation the hydrogen-rich gas is subjected to scrubbing with a physical solvent (e.g. methanol) capable of absorbing residual carbon dioxide before being changed in pressure and temperature, for example at least part of the carbon dioxide being returned to a desired temperature and pressure. In some cases, some or all of the carbon dioxide is subjected to further process steps before being returned to the desired temperature and pressure.
The improvement of the purity of the carbon dioxide can in turn leads to improved environmental benefits when the hydrogen rich gas is used subsequently for example in a hydrogen fuelled decarbonised power station. Furthermore in developing this improvement it has been recognised that such a two step process has a more general utility for example in the separation of hydrogen sulphide from natural gas and the removal of carbon dioxide and/or carbonyl sulphide and/or hydrogen sulphide from nitrogen, oxygen, air or synthesis gas.
According to an aspect of the present invention there is therefore provided a process for separating a mixture of gases into a relatively condensable first component and a relatively non-condensable second component and wherein (1) the first component comprises one or more gases selected from the group consisting of carbon dioxide, carbonyl sulphide and hydrogen sulphide and (2) the second component comprises one or more gases selected from the group consisting of hydrogen, methane, ethane, carbon monoxide, nitrogen, oxygen and synthesis gas wherein the process comprises the steps of:
The present invention has the particular advantage that by scrubbing the second component fraction whilst it is at a low temperature and high pressure the amount of the residual first component(s) retained by the physical solvent is improved. This means that relative to conventional alcohol scrubbing processes, for example the Rectisol™ and Selexol™ processes' which operate at relatively low pressure the rate of alcohol flow though the scrubber can be significantly reduced making the unit much more compact than is conventional in the art. The present invention therefore exploits the thermodynamic synergy between the fractionation unit and the scrubber to achieve an improved and more energy efficient process design. As a consequence, it also allows the capital cost of the scrubber and the footprint of any plant using the process to be minimized. The latter is an especially significant benefit where the process is to be used on a site where space is at a premium e.g. an offshore gas platform.
Alternatively, or in addition, where the size of the fractionation column used for the scrubbing is dominated by the volumetric flow of the gas, the use of high pressure gas can allow for a smaller column to be used because high pressure gas has lower specific volume than lower pressure gas.
According to a further aspect of the invention there is provided a method of separating a relatively condensable first component from a mixture of gases including the first component and a relatively non-condensable second component, wherein the method includes the steps of:
The physical solvent is methanol in the examples described herein, but other solvents could be used. For example the solvent may include one or more of dimethyl ethers of polyethylene glycol and N-methyl-2-pyrrolidone (NMP). Preferably the physical solvent is one in which the adsorption of the CO2 in the solvent relies on the solubility of the CO2 in the solvent rather than, for example a chemical reaction with the solvent. The solvent may for example be regenerated by changing pressure and/or temperature.
In examples of the invention described herein, at least a part of the expanding and/or warming of the gaseous second component fraction occurs downstream of the CO2 absorber. Thus the invention may further include the step of warming and/or expanding the scrubbed second component fraction.
Preferably the warming uses at least one heat exchanger to exchange heat with a process stream, for example the mixture of step (a).
Preferably the expansion includes using at least one turbo-expander capable of recovering mechanical work.
Alternatively, or in addition, expansion and/or warming of the second component fraction may occur before the separated gaseous second component is passed to the physical solvent.
The method may further include the step of warming and/expanding the second component fraction before the scrubbing step.
A described above, the warming may include heat exchange with another process stream. The expansion preferably recovers mechanical work. Preferably the expansion is carried out using a turbine device.
In some examples, the second component fraction will be expanded and warmed both before and after the CO2 absorption step.
It will be understood therefore that the pressure at which the CO2 absorption is carried out may be substantially the same as the pressure at which the two-phase separation step occurs, or may be lower, for example if expansion of the second component fraction has occurred between the separator and the solvent system.
Preferably, the solvent separation is nevertheless carried out at an elevated pressure. For example, the pressure of the solvent separation step may be 50 bar or more, for example 60 bar or more, for example 80 bar or more, for example 100 bar or more, for example 120 bar or more. In some applications, the pressure of the solvent stage may be between 80 bar and 400 bar.
Preferably at least 50% of the first component is separated from the two-phase mixture in step b. For example, preferably at least 50%, for example at least 70%, of the carbon dioxide in the gas mixture is separated in the two-phase separation in this or other aspects.
In the case where carbon dioxide is being separated from gas mixtures comprising carbon dioxide, hydrogen and optionally nitrogen, the hydrogen/nitrogen mixture exiting the fractionation unit, depending upon that unit's exact operating temperature and pressure, can contain as much as 10% of the original carbon dioxide level. By using the fractionation unit and scrubber in the synergistic way described herein this residuum can be reduced to 5% or even less.
In stage (a) of the process of the present invention, the gas mixture is cooled and compressed until the first components contained therein liquefy. This is done by at least one, typically a series of, compressors optionally linked through inter-stage heat exchangers which remove the heat generated by the compression. After the final compression, the compressed mixture is suitably cooled once more in a final heat exchanger against either or both of the cold separated first and second components thereby allowing the latter streams to be warmed back towards their final desired temperature. The inter-stage heat exchangers can also use the first and second components as coolant if desired. In this case the flows of these components are configured so that as they warm up they progressively encounter warmer and warmer streams to be cooled. However it is preferred that the flows of the cold separated first and second components are configured so as to flow though one or more multi-channel diffusion bonded and/or micro channel heat exchangers designed to cool the already cooled gas mixture to the operating temperature of the fractionation unit. Examples of such hear exchangers are described for example in EP 0212878 and WO 2004/017008 the contents of which are incorporated by reference herein.
In the final heat exchanger the mixture is further cooled to create a two phase gas/liquid mixture. The exact temperature and pressure required to achieve this will depend on exactly how selective this fractionation unit is required to be but in all cases it is critical that the mixture is prevented from becoming supercritical on the one hand or that the first component(s) freeze out on the other. In practical terms, this means ensuring that at one extreme the temperature and pressure of the mixture should not exceed both the critical temperature and critical pressure. For each of the first components of the present invention these key parameters for pure components are as follows: carbon dioxide Tc=31.4° C. and Pc=73.9 bar; carbonyl sulphide Tc=105.8° C. and Pc=63.5 bar and hydrogen sulphide Tc=100.5° C. and Pc=90 bar. It will be appreciated that the critical temperature and critical pressure of gas mixtures may be higher depending on their compositions. At the other extreme it is necessary to ensure that the temperature of the mixture should not fall below each of the first components' triple point temperatures which are as follows: carbon dioxide Tt=−56° C.; carbonyl sulphide Tt=−138.9° C. and hydrogen sulphide Tf=−85.5° C. Within these boundary conditions it is desirable that the operating temperature of the fractionation unit is at least 20° C. below the boiling point of the first component at the operating pressure in order to obtain efficient separation
It will be appreciated by one of ordinary skill that the thermodynamic constraints referred to above relate to ideal systems and that for mixtures which exhibit significant deviation from ideality it is possible that in certain cases the presence of one or more of the second components my depress or elevate a first component's triple point temperature further. To allow for this possibility it is generally preferred that the temperature of the fractionation unit should be at least 3° C. preferably at least 5° C. above the theoretical triple point temperature of the relevant first component. In practical terms and for the mixtures described herein this generally means operating the fractionation unit at a temperature in the range −25 to −53° C. and preferably in the range −40 to −50° C. At the same time the pressure should suitably be in the range 80 to 400 bar, preferably 150 to 250 bar.
The fractionation unit used in the process of the present invention is suitably a conventional gas-liquid separator adapted to work at the high pressures and low temperatures set out above. In such vessels the gaseous second fraction is typically taken off overhead and the liquid first component is removed at or near the bottom. The pressure drop across the fractionator is typically no more than between 0.1 and 0.5 bar.
After separation, the gaseous second component is fed to a scrubber where in step (c) of the process it is contacted with preferably a continuously fed and continuously removed stream of cold solvent, for example alcohol, in order to extract residual first component therefrom. This is typically effected by continuously contacting a stream of the second component with the cold solvent, for example alcohol, stream under conditions which cause intimate and turbulent mixing of the two for example by counter-current mixing or by sparging the second component through the solvent. Under these conditions the residual first component dissolves in the solvent, for example alcohol, and is removed from the system by way of the effluent from the scrubber. As mentioned above by effecting this contacting at the high pressure and low temperature characteristic of the fractionation unit a significant part of the residual first component is caused to be absorbed by and to dissolve in the solvent, for example alcohol, in accordance with Henry's law. The thermodynamic driving force behind this absorption process, which is enhanced at high pressures, works synergistically with the increased capacity of the solvent at low temperature to hold proportionately more first component making a highly efficient system. In particular it is more efficient than the alternative i.e. conventional use of a Rectisol or Selexol treatment carried out at much lower pressures after the second component had been fully returned to or near to its final desired state.
When conducting step (c) it is preferred that where an alcohol solvent is used it is selected from methanol, ethanol, the isomers of propanol and glycols and glycol ethers formed by oligomerisation of ethylene or propylene glycol. For obvious reasons the alcohol or other solvent chosen should be one which will not freeze under the operating conditions of the scrubber. Since it is preferred that the step (c) is conducted immediately after step (b) with no intermediate treatment of the second component the operating temperature and pressure of the scrubber should be the same as or substantially the same as those of the fractionation unit. However the temperature and pressure ranges disclosed above for the fractionation unit are applicable mutatis mutandis to the scrubber irrespective of whether any treatment of the second component has occurred between the fractionation unit and the scrubber. It will be appreciated however that step (c) works most efficiently when the cold solvent, for example alcohol, solvent is fed to the scrubber at or close to the latter's operating temperature.
Furthermore, as discussed further herein, the second component may be expanded after the phase separation step and prior to the scrubber. In such an example, the operating pressure of the solvent system is lower than that of the phase separation apparatus.
The effluent solvent from the scrubber may be passed to a treater where fresh solvent may be regenerated by distillation and overhead removal of the first component in gaseous form. Thereafter the regenerated solvent can be cooled and recycled to the scrubber. The gaseous first component so obtained can thereafter be either disposed of or liquefied and combined with the main first component stream before doing so.
After leaving the scrubber the second component is in step (d) warmed and decompressed to restore it to any temperature and pressure required for its further utilisation. In order to ensure that the energy utilisation of the process is as efficient as possible this step is effected by passing the second component through one or more turbo expanders and in the case of a series of turbo expanders a series of inter-stage heat exchangers. In each turbo expander the second component is progressively expanded isentropically reducing its pressure and releasing expansion energy which in turn drives a turbine capable of recovering this energy as mechanical work. This mechanical work can if desired be used elsewhere in the process, for example to drive the compressors used in stage (a), thereby minimising overall energy usage. At the same time the expansion of the second component causes it to cool and the cooling capacity generated can be used in the interstage coolers to cool warmer streams especially those generated in stage (a). In a preferred embodiment these interstage coolers are integrated into a single or array of multichannel heat exchangers in order to manage the cooling capacity of the second component even more efficiently. In performing these series of expansions and coolings it is important not to let the temperature of the expanded second component after each turbo expansion fall below the triple point temperature of the first component in order to prevent progressive blockage of the transfer line between each turbo expander and interstage cooler by build-up over time of frozen, solid first component derived from any small amounts still remaining in the treated second component. Once the second component has been reduced to its desired temperature and pressure it can be used for its chosen duty.
Subsequent treatment of the first component recovered in stage (b) will depend to a certain extent on its exact nature. If it is liquid carbon dioxide it may for example be piped or tankered offsite for underground storage. In this case it is desirable to liquefy any further gaseous first component recovered in step (c) and combine it with the material recovered in step (b). If on the other hand the liquid first component is carbonyl sulphide or hydrogen sulphide it may be more beneficial to restore it to the gaseous state so that the sulphur it contains can be recovered as elemental sulphur for example by feeding the first component to a Claus plant. In this case any gaseous first component recovered in step (c) can either be fed directly to such a plant or combined with that recovered in step (b) after it has been warmed and decompressed in a manner analogous to step (d).
It will be apparent that the novel process of the present application can manifest itself as a separation plant employing the process described above. Accordingly there is provided, in an embodiment of the present invention, a gas separation plant for separating a mixture of gases into a relatively condensable first component and a relatively non condensable second component wherein (1) the first component comprises one or more gases selected from the group consisting of carbon dioxide, carbonyl sulphide and hydrogen sulphide and (2) the second component comprises one or more gases selected from the group consisting of hydrogen, methane, ethane, carbon monoxide, nitrogen, oxygen and synthesis gas wherein it comprises;
Preferably the expansion system for warming and expanding the scrubbed second component fraction comprises at least one turbo expander for recovering mechanical work and means for supplying cooling capacity to at least one of the heat exchangers in the cooling and compression system.
The expansion system may be at least partly arranged upstream and/or downstream of the scrubber.
The separation plant described about can be stand-alone or part of a larger complex for example a hydrogen power plant, a synthesis gas generating unit or a natural gas offshore platform.
Where the first component includes hydrogen sulphide, preferably and the method includes separating hydrogen sulphide from the mixture of gases. In preferred arrangements, preferably the separation takes place prior to the forming of the two-phase gas-liquid mixture.
By separating out the hydrogen sulphide upstream of the two-phase mixture separation, the separation of H2S from the mixture can be carried out separately from the separation of CO2 from the mixture. This can be desirable in many applications, in particular where high purity carbon dioxide is sought, and also can lead to efficiencies as well as flexibility of choice of the sulphur removal system.
The upstream sulphide separation step, for example hydrogen sulphide separation, preferably is carried out at relatively low pressure, that is prior to the compression step. In some applications, there could be substantially no compression upstream of the hydrogen sulphide separation step. In other arrangements, one or more upstream compression steps could be used. For example, one or more compressors could be arranged upstream of the H2S absorption apparatus, and/or one or more compressors arranged downstream of the H2S absorption apparatus. Thus in some examples. The H2S may be carried out at an elevated pressure. Preferably the elevated pressure is less than the pressure of the phase separation unit, although it is contemplated that the H2S apparatus and phase separation unit could operate at similar pressures.
Any appropriate sulphur removal apparatus could be used, for example Selexol®, Rectisol, or other. Preferably the separation method is one which does not involve an elevated temperature; preferably the gas mixture is not heated significantly prior or during the sulphur removal step. This can lead to efficiencies, in particular in view of the cooling step which is carried out after the sulphur removal step.
However, it is envisaged that a high temperature sulphur removal process could be used. If so, preferably it is carried out upstream of any significant cooling of the stream to reduce potential inefficiencies in cooling, reheating, and re-cooling of the stream.
In a preferred arrangement, the sulphur removal process involves a biological process. For example the hydrogen sulphide separation step may include a biological process. For example, the sulphur removal process may include the Paques process (Shell).
This feature is of particular importance and is provided independently. Thus a further aspect of the invention provides a method of separating a mixture of gases to separate a relatively condensable first component from a relatively non-condensable second component wherein the first component comprises carbon dioxide and the second component comprises one or more gases selected from the group consisting of hydrogen, methane, ethane, carbon monoxide, nitrogen, oxygen and synthesis gas wherein the mixture of gases further includes hydrogen sulphide and the method includes the steps of:
(a) treating the gas mixture to remove hydrogen sulphide;
(b) compressing and cooling the treated gas mixture in at least one compressor and at least one heat exchanger to a temperature and elevated pressure at which the first component condenses and a two-phase gas-liquid mixture is formed;
(c) separating the two phase mixture into separate liquid first and gaseous second component fractions;
(d) extracting residual first component from the separated gaseous second component fraction by scrubbing the second component fraction at elevated pressure with a solvent in a scrubber and
wherein the method further includes the steps of warming and expanding the second component fraction using at least one heat exchanger to exchange heat with the mixture of step (b) and at least one expander capable of recovering mechanical work.
The solvent is preferably a physical solvent. For example, the solvent may comprise an alcohol, for example methanol.
The warming and/or expanding of the second component fraction may each be carried out upstream and/or downstream of the scrubbing step. In an example, expansion and heat exchange of the second component fraction occurs both upstream and downstream of the solvent separation step.
The separation of the sulphur-containing component may be carried out using a Paques process.
Thus the plant according may further include a hydrogen sulphide removal system upstream of the fractionation unit for removing a sulphur-containing component from the gas mixture. The hydrogen sulphide removal system may for example include a Paques apparatus.
In examples, the Paques process removes hydrogen sulphide form the synthesis gas stream in an absorber using a caustic solution. Hydrogen sulphide is then removed from the solvent in a biological reactor which produces elemental sulphur and regenerates the solvent which is recycled to the absorber. Advantages of such a process over for example a Selexol™ acid gas remover include lower energy consumption. Capital cost reductions can also be realized in some applications due to the need for fewer items of equipment. For example there is no need for a Claus plant for sulphur recovery, or for apparatus for the removal of the tail gas (tail gas treatment unit (TGTU)). Energy consumption may also be lower because little energy is required to regenerate the solvent in the biological reactor. Furthermore, where hydrogen produced by the process is used in a power island, the steam balance of the power island can be positively impacted.
In examples described herein, to increase the amount of carbon dioxide separated, to reach high levels of carbon capture rates, for example in excess of 90%, a system is provided in which a phase separation apparatus for separating carbon dioxide from a gas mixture is used in combination with a solvent separation apparatus. For example, a methanol absorber is used in combination with the phase separation process and removes carbon dioxide from the incondensable hydrogen rich gas at high pressure and low temperature. The carbon dioxide rich methanol is then sent to a regeneration plant where for example through combination of pressure reduction and temperature increase the carbon dioxide is liberated fro the methanol. The carbon dioxide released can then be captured, and compressed to export pressure.
A further aspect of the invention provides a system for separating a relatively condensable first component from a mixture of gases including the first component and a relatively non-condensable second component wherein the apparatus includes:
(a) a phase separation apparatus including
Preferably in some examples, the system is arranged such that substantially all of the first component separated by the solvent separation apparatus is fed to the phase separation apparatus.
Preferably at least 50% of the first component is separated from the two-phase mixture in step aii.
Thus some or all of the first component, for example liquid carbon dioxide, may be recycled to the phase separation apparatus.
In this way, the system may have a single take-off point for the separated first component. Preferably this single take-off point is from the outlet of the phase separation apparatus.
Thus in examples, carbon dioxide liberated from the solvent, for example methanol, of the solvent system is recycled to the front end of the phase separation unit and then joins the fresh feed from the hydrogen sulphide removal process (where present). The result would be an increase in the flow rate of the carbon dioxide capture stream in the phase separation apparatus. If for example the solvent system is designed to remove substantially all of the carbon dioxide from the hydrogen stream then the only carbon dioxide capture stream in the process is that from the phase separation apparatus and the capture level might be in excess of 99%.
Also by not capturing carbon dioxide separately at the solvent separation apparatus, separate compression and pumping for carbon dioxide liberated by the solvent will not be required.
A broad aspect of the invention provides a method of separating carbon dioxide from a mixed gas using a system including a plurality of carbon dioxide separation units, wherein the carbon dioxide stream separated by a first carbon dioxide separation unit is fed to a second carbon dioxide unit, such that a separated carbon dioxide stream is withdrawn from the system from a single region of the system.
Preferably the carbon dioxide stream is withdrawn from the second unit.
In examples described the first component comprises carbon dioxide. Preferably the solvent unit includes an alcohol, preferably methanol. Other solvents could be used as appropriate, for example including those described herein. Preferably the solvent is one having a suitable affinity to carbon dioxide so that the carbon dioxide in the gas mixture can be separated from the mixture, and then subsequently released to form a carbon dioxide product stream which is then recirculated to the phase separation apparatus. The carbon dioxide product stream (or other first component product stream) may include other components for example impurities. In such a case, the stream could be fed to the H2S absorber to remove impurities.
The phase separation apparatus may generate a gas product stream comprising the second component of the two-phase gas-liquid mixture, the solvent separation apparatus being arranged downstream of the phase separation apparatus such that at least a portion of gas of the gas product stream including the second component is fed to the inlet of the solvent separation apparatus.
In some examples, the initial gas mixture further includes hydrogen, a hydrogen-rich gas forming the second gaseous component of the two-phase gas-liquid mixture in the phase separation apparatus.
The system may be arranged to feed a liquid stream including the condensed first component from the separator of the phase separation apparatus to a heat exchanger for exchanging heat within the system. Thus the condensed first component may be used as an indirect refrigerant within the system, for example exchanging heat with another process stream of the system.
The system may be adapted such that the condensed first component at least partly evaporates at or upstream of the heat exchanger. By evaporation of the first component, for example carbon dioxide, additional cooling can be provided within the system. For example, the first component stream may be flashed, for example across a valve, upstream of or at a heat exchanger. In other arrangements, the second component stream may be used as a coolant in liquid form. This feature of use of the first component stream as an internal coolant may be provided as a part of any of the examples described herein and may be provided in relation to any one of the aspects herein. At least a part, or all, of the first component stream, for example carbon dioxide stream, may be used as an internal coolant.
The system preferably includes an expander system arranged for receiving separated second component, the expander system including at least one heat exchanger for exchanging heat within the system and at least one expander capable of recovering mechanical work.
Thus work of the compressor and/or cooling apparatus in increasing the pressure and decreasing the temperature for the phase separation can be recovered from the second component stream.
Preferably the expander system is arranged to receive separated second component from the solvent separation system.
Alternatively, or in addition, an expander system is arranged to receive separated second component from the phase separation apparatus.
In some arrangements, some expansion can be carried out upstream of the solvent system, with further expansion carried out downstream of the solvent system.
In preferred arrangements, the solvent separation apparatus is arranged downstream of the phase separation apparatus and the solvent separation is carried out at elevated pressure and/or low temperature, prior to work being recovered using the expander system.
In other arrangements, the expander system may be arranged between the phase separation apparatus and the solvent separation apparatus, the recovery of some or all of the pressure and temperature being carried out upstream of the solvent separation apparatus. However, in some cases this will not be preferred because the first component separated by the solvent separation apparatus and recirculated to the phase separation apparatus will have lost pressure and more work will be required by the compressor of the phase separation system to repressurise it.
Preferably the system has a single outlet point for separated second component. Preferably that outlet is downstream of the expander system. Preferably where the second component stream includes hydrogen, it may be fed directly or indirectly to the inlet of a hydrogen power plant. For example, hydrogen rich gas may form a feed gas for a combustor of a gas turbine. In other applications, the separated second component, for example a hydrogen rich gas, may be removed for storage or further processing.
Preferably the system has a single outlet point for the separated first component. The separated first component may be passed through for example one or more heat exchangers and/or expanders after separation to recover work. However, in some cases, preferably the elevated pressure of the first component is maintained. High pressure carbon dioxide, preferably in liquid phase, can preferably be suitable for direct sequestration and/or use in enhanced oil recovery (EOR).
In other applications, in particular where a component for example hydrogen is required at relatively high pressure, less, or no, expansion of the second component may be provided. In this case, the expanders may not be included.
Preferably the system further includes a sulphur removal apparatus upstream of the phase separation apparatus. As for example discussed in more detail below, the sulphur removal apparatus may include a biological system, for example a Paques system.
One or more features of other aspects of the invention described herein may be combined with features of the present aspect. For example, the compressor of the phase separation apparatus may include a plurality of compressors, for example arranged in series. Preferably the recirculated first component passed from the solvent separation apparatus to the phase separation apparatus is introduced to an appropriate point relative to the series of compressors. For example the recirculated first component stream may be introduced between adjacent compressors of the series.
According to another aspect of the invention there is provided a method of separating a relatively condensable first component from a mixture of gases including the first component and a relatively non-condensable second component wherein the method includes the steps of:
compressing and cooling a gas mixture such that the first component condenses and a two-phase gas-liquid mixture is formed;
separating the condensed first component stream and a second gaseous component stream from the two-phase mixture
feeding the second gaseous component stream to a solvent separation apparatus, the solvent separation apparatus including a solvent and being adapted for separating a first component stream from the received gas mixture; and
feeding first component stream from the solvent separation apparatus to the phase separation apparatus.
Preferably substantially all of the first component separated by the solvent separation apparatus is fed to the phase separation apparatus.
Preferably the first component stream is fed to the compressing and cooling stage of the solvent separation apparatus.
The two-phase mixture may be fed to the separator is at a pressure of above 60 bar, preferably at least 80 bar, preferably above 120 bar. In some examples, the pressure of the two-phase mixture will be 150 bar or above.
The pressure of the CO2 solvent separation apparatus may be greater than 35 bar, preferably at least 50 bar, preferably greater than 60 bar, preferably at least 80 bar, preferably greater than 100 bar. The solvent stage pressure may be from about 80 bar to about 400 bar.
By carrying out the separation at high pressure, efficiencies may be made because the first component which is fed back from the solvent separation apparatus to the phase separation apparatus will be of elevated pressure so that the compression required in the phase separation apparatus will be relatively less. The compression for the subsequent phase separation may therefore be carried out using some or all of the compression apparatus used in relation to the compression of the initial feed stream.
The pressures of operation of the separator of the phase separation apparatus and of the solvent separation apparatus may in some cases be substantially the same. While there may be some inherent pressure loss in the system between the phase separation system and the solvent separation system, preferably there is little or no active drop in pressure effected. Preferably the pressure of operation of the solvent unit is more than 30%, preferably more than 40% or more than 50% of the pressure of operation of the phase separation apparatus. In some examples the pressure of operation of the solvent unit is more than 70%, for example more than 80%, for example more than 90% of the pressure of the separator of the phase separation apparatus.
Preferably the first component includes carbon dioxide, and the second component includes hydrogen.
The method may further include warming and expanding a separated second component fraction using at least one heat exchanger to exchange heat and at least one turbo-expander capable of recovering mechanical work.
Preferably the heat exchanger exchanges heat with an internal stream that is with another stream of the separation system. Alternatively or in addition external coolants and/or refrigerants could be used.
Features described in relation to one aspect of the invention may be applied to other aspects of the invention in any appropriate combination. Features of method aspects may be applied to apparatus aspects and vice versa.
The invention also provides a method and/or apparatus being substantially as herein described, preferably having reference to one or more of the accompanying figures.
Embodiments of aspects of the present invention will now be illustrated purely by way of example having reference to the following figures:
The low pressure gas stream 1D that exits compressor C1 is at a pressure of 76 bar and a temperature of 68° C., the increase in temperature arising from the heat of compression. Stream 1D is then cooled in interstage heat exchanger E1 against a cold water stream and passed to compressor C2 thereby generating a stream 2D having a pressure of 100 bar and a temperature of 69° C. Stream 2D is then cooled in heat exchanger E2 against a cold water stream and passed to the first HP compressor C3 thereby generating a high pressure gas stream 3D having a pressure of 132 bar and a temperature of 69° C. Stream 3D is then in turn cooled in heat exchanger E3 against a cold water stream and passed to compressor C4 thereby forming a gas stream 4D having a pressure of 175 bar and a temperature of 70.° C. Stream 4D is then finally cooled in heat exchanger E4 against a cold water stream thereby generating a high pressure shifted synthesis gas stream S1 having a temperature of about 40° C. It will be appreciated that the number of compression stages used has been chosen to minimise power consumption and that it would be possible to achieve this degree of compression using three or even two compression stages. It will also be appreciated that stream 4D could be cooled further in heat exchanger E4 by using a specially designed refrigerant system employing, for example, propane, propylene or ammonia. This would have the advantage of reducing the cooling demand on the next heat exchanger multichannel heat exchanger LNG-100.
The cooled high pressure stream S1 exiting E4 is next passed through multichannel an array of diffusion bonded heat exchangers ex Heatric UK (shown schematically as LNG-100) where it is cooled against a plurality of cold process streams (see below derived from the decompression system (see below) thereby generating a two phase gas/liquid stream S2 having a pressure of 172 bar and a temperature of about −50° C. This stream is passed directly to gas-liquid separator vessel F360 where a hydrogen rich gas is separated from a liquid phase containing hydrogen sulphide and carbon dioxide. The hydrogen rich gas is removed overhead as stream S2V and fed at the same temperature and pressure via line 23 to a scrubber A3. In A3 the hydrogen rich gas is contacted counter currently with liquid methanol delivered via line 25 at a temperature of −50° C. in order to remove any residual carbon dioxide and hydrogen sulphide contained therein. The treated hydrogen-rich gas is removed from A3 via line 1N and then routed to turbo expander EX1 where it is expanded isentropically to lower pressure. The person skilled in the art will understand that isentropic expansion of this gas stream results in it being cooled. Accordingly the hydrogen-rich gas exits EX1 at a pressure of 112 bar and a temperature of −53° C. and is routed through multichannel heat exchanger LNG-100 where it is heat exchanged with the high pressure gas stream S1 up to a temperature of −30° C. and then passed to turbo expander EX2 (via line 2N) where it is expanded yet again to form stream 2T at a pressure of 75 bar and a temperature of −53° C.
At the same time a liquid stream containing carbon dioxide and hydrogen sulphide S2L is withdrawn from the bottom of F360 and is flashed across valve VLV-109 thereby generating a further two phase stream 18 that is passed to flash vessel F150. A hydrogen-rich gas stream S2LV that is withdrawn from the top of this vessel and combined with stream 2T to form combined vapour stream 2™ at point M1. Gas stream 2™ is then passed through multichannel heat exchanger LNG-100 thereby again cooling stream S1. The hydrogen-rich gas stream 3N that exits the multichannel heat exchanger LNG-100, now at a temperature of −30° C. is then passed to turbo expander EX3 where it is expanded once again to a pressure of 48 barg and a temperature of −53° C. (line 3T). This stream is warmed yet again to −30° C. in LNG-100 before being expanded one last time in turbo expander EX4 to a pressure of 31 bar and a temperature of −53C. Finally the gas stream is passed through LNG-100 one more time and optionally a series of other heat exchangers (not shown) where it is warmed to a final temperature of 25° C. The warm hydrogen-rich gas stream then leaves the plant via line 29.
Meanwhile a liquid stream consisting of carbon dioxide and hydrogen sulphide is withdrawn from the bottom of F150 and passed through LNG-100 where it too is warmed against stream S1 and optionally other heat exchangers before being exported offsite.
Depending on the cold heat requirement of LNG-100, the liquid stream may be flashed through a valve (not shown in
Turning now to a consideration of the methanol solvent in the solvent system, after contact with the gaseous second component in scrubber A3, the methanol solvent now rich in carbon dioxide and hydrogen sulphide, is removed an fed via line 26 to the head of a stripper column A2 in which the hydrogen sulphide and carbon dioxide and the methanol are separated. A hydrogen sulphide and carbon dioxide rich gas stream is then removed overhead via line 20 for optional further separation prior to treating the hydrogen sulphide in e.g. a Claus Plant. Lean methanol is then returned to A3 via line 25 and A2 is provided with a reboiler serviced by lines 21 and 22 to maintain the methanol at the correct temperature.
In the arrangement of
The low pressure gas stream 1D that exits compressor C1 is at a pressure of 76 bar and a temperature of 68° C., the increase in temperature arising from the heat of compression. Stream 1D is then cooled in interstage heat exchanger E1 against a cold water stream and passed to compressor C2 thereby generating a stream 2D having a pressure of 100 bar and a temperature of 69° C. Stream 2D is then cooled in heat exchanger E2 against a cold water stream and passed to the first HP compressor C3 thereby generating a high pressure gas stream 3D having a pressure of 132 bar and a temperature of 69° C. Stream 3D is then in turn cooled in heat exchanger E3 against a cold water stream and passed to compressor C4 thereby forming a gas stream 4D having a pressure of 175 bar and a temperature of 70° C. Stream 4D is then finally cooled in heat exchanger E4 against a cold water stream thereby generating a high pressure shifted synthesis gas stream S1 having a temperature of about 40° C. It will be appreciated that the number of compression stages used has been chosen to minimise power consumption and that it would be possible to achieve this degree of compression using three or even two compression stages. It will also be appreciated that stream 4D could be cooled further in heat exchanger E4 by using a specially designed refrigerant system employing, for example, propane, propylene or ammonia. This would have the advantage of reducing the cooling demand on the next heat exchanger multi-channel heat exchanger LNG-100.
The cooled high pressure stream S1 exiting E4 is next passed through multi-channel an array of diffusion bonded heat exchangers ex Heatric UK (shown schematically as LNG-100) where it is cooled against a plurality of cold process streams (see below derived from the decompression system (see below) thereby generating a two phase gas/liquid stream S2 having a pressure of 172 bar and a temperature of about −50° C. This stream is passed directly to gas-liquid separator vessel F360 where a hydrogen rich gas is separated from a liquid phase containing carbon dioxide. The hydrogen rich gas is removed overhead as stream S2V and fed at substantially the same temperature and pressure as the feed 112 to the solvent separation apparatus 118. Any appropriate arrangement may be used. For example, the arrangement of
The treated hydrogen-rich gas 122 is removed from A3 and is then routed to the expander apparatus, initially to turbo expander EX1 where it is expanded isentropically to lower pressure. The person skilled in the art will understand that isentropic expansion of this gas stream results in it being cooled. Accordingly the hydrogen-rich gas exits EX1 at a pressure of 112 bar and a temperature of −50° C. and is routed through multi-channel heat exchanger LNG-100 where it is heat exchanged with the high pressure gas stream S1 up to a temperature of −33° C. and then passed to turbo expander EX2 (via line 2N) where it is expanded yet again to form stream 2T at a pressure of 75 bar and a temperature of −53° C.
At the same time a liquid stream containing carbon dioxide S2L is withdrawn from the bottom of F360 and is flashed across valve VLV-109 thereby generating a further two phase stream 18 that is passed to flash vessel F150. A hydrogen-rich gas stream S2LV that is withdrawn from the top of this vessel and combined with stream 2T to form combined vapour stream 2™ at point M1. Gas stream 2™ is then passed through multi-channel heat exchanger LNG-100 thereby again cooling stream S1. The hydrogen-rich gas stream 3N that exits the multi-channel heat exchanger LNG-100, now at a temperature of −30° C. is then passed to turbo expander EX3 where it is expanded once again to a pressure of 48 barg and a temperature of −53° C. (line 3T). This stream is warmed yet again to −30° C. in LNG-100 before being expanded one last time in turbo expander EX4 to a pressure of 31 bar and a temperature of −50° C. Finally the gas stream is passed through LNG-100 one more time and optionally a series of other heat exchangers (not shown) where it is warmed to a final temperature of 25° C. The warm hydrogen-rich gas stream then leaves the plant via line 122′.
Meanwhile a liquid stream consisting of carbon dioxide is withdrawn from the bottom of F150 and passed through LNG-100 where it too is warmed against stream S1 and optionally other heat exchangers before being exported offsite as the CO2 output 114.
Meanwhile in the solvent system 118, after contact with the gaseous second component in scrubber A3, the methanol solvent now rich in carbon dioxide, is removed an fed via line 26 (see
Lean methanol is then returned to A3 via line 25 and A2 is provided with a reboiler serviced by lines 21 and 22 to maintain the methanol at the correct temperature.
It will be noted that there is a single take-off point for carbon dioxide in the system of
By integrating the phase separation process, expander apparatus and solvent system, an arrangement in which all of the carbon dioxide capture occurs within the phase separation process can be achieved. The integrated solvent system recovers incondensable carbon dioxide from the syngas as a recycle stream. The high pressure and low temperature conditions of the phase separation process are synergistic with physical solvent systems resulting in improved energy efficiency and reduced plant footprint.
By separating the hydrogen sulphide removal process from the carbon dioxide separation system, the best available state of the art technology can be employed and further simplifications can be achieved. For example, the introduction of a biological sulphur removal process, for example Paques, can eliminate the need for a sulphur recovery unit (SRU) and Tail Gas Treatment Unit (TGTU) which may otherwise be required. Also, by decoupling the hydrogen sulphide and carbon dioxide removal allows future low energy options for hydrogen sulphide removal to be employed.
It will be understood that features of the invention have been described above only by way of example, an variations can be made within the scope of the aspects of the invention.
| Number | Date | Country | Kind |
|---|---|---|---|
| PCT GB2009 001810 | Jul 2009 | GB | national |
| PCT GB2009 002383 | Oct 2009 | GB | national |
| PCTGB2009 002895 | Dec 2009 | GB | national |
| 10250096.4 | Jan 2010 | EP | regional |
| 1001710.1 | Feb 2010 | GB | national |
| 1012110.1 | Jul 2010 | GB | national |
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/GB2010/001405 | 7/23/2010 | WO | 00 | 1/24/2012 |
