Process for the production of hydrogen

Abstract

A process is disclosed for the production of hydrogen, the process comprises subjecting a gaseous mixture comprising a hydrocarbon and steam, and having a steam to carbon ratio in a range of from 0.4

Description

Field of the Invention This invention relates to a process for converting hydrocarbons to hydrogen (H?), while also raising high or medium pressure steam that is available for export from the process. Background to the Invention Processes for generating low carbon hydrogen, often referred to as “blue hydrogen”, are well-known. Generally, these processes include a steam reforming stage combined with a water-gas shift reaction stage. In the steam reforming stage, a feed stream of a gaseous mixture of hydrocarbon (predominantly methane) and steam is reformed (CH4 + H2O CO + 3H2) to produce a reformed gas mixture. The steam reforming stage typically comprises an autothermal reformer (ATR) so that most of the CO2 is formed in the reformed gas stream itself, rather than in a low-pressure flue gas stream. The CO2 can thus be removed more efficiently. In an ATR, high purity oxygen is typically combusted sub-stoichiometrically with the feed stream and the combusted gas then undergoes steam reforming reactions adiabatically over a catalyst bed. In the water-gas shift reaction stage, the H2 content of the reformed gas mixture is increased using a water-gas shift reaction (CO + H2O -> CO2 + Hz), to produce a hydrogen-enriched reformed gas mixture, from which carbon dioxide (CO2) is then removed in a CO2 removal stage. The carbon dioxide can then be sequestered. One major use of blue hydrogen is in the decarbonisation of refinery or petrochemicals facilities, which often comprise many fired heaters and fired boilers which use natural gas or other fossil-derived fuels. Capturing carbon dioxide from the low-pressure flue gases of all these sources, may not be an effective solution. Instead, low carbon hydrogen can be used as a fuel in the fired heaters. While low carbon hydrogen could be used in the fired boilers also, it is more efficient if the steam produced in the fired boilers can be replaced with low carbon steam. In these situations, therefore, there is a need for a process to make blue hydrogen, which not only has a high feedstock use efficiency, but which also exports a large quantity of low carbon (blue) steam. The water-gas shift reaction is typically performed adiabatically in two sequential steps, respectively at a high temperature (above 320°C) and at a low temperature (below 230°C but typically above 190°C), to achieve a high overall conversion of the carbon monoxide (CO) in the reformed gas mixture. Thus, unconverted CO in the process is reduced, which also reduces the amount of unconverted CO that has to be recycled to the steam reforming stage and / or is lost from the process, typically as a constituent of a low-pressure fuel stream. In addition to being driven by the steam demand of reforming, the total steam demand of such processes is also determined by the steam demand of the water-gas shift reaction, not only to maximise CO shift to H2 but also, in respect of the high temperature shift in particular, to prevent damage to conventional iron-containing high temperature water-gas shift catalysts through chemical over-reduction by CO and H2. Steam is typically added to the feed upstream of reforming. The amount of steam added is typically characterised by defining a steam to carbon ratio as the mols of steam in the gaseous mixture fed to the reforming divided by the mols of carbon in hydrocarbon compounds in the gaseous mixture. Unreacted steam in the reformed gas mixture will be available to participate in the water-gas shift reaction. Steam can also be added upstream of the water-gas shift reactor but after the reforming. The process steam demand of such processes, specifically for reforming and for the water-gas shift reaction, is typically for medium pressure steam, e.g. steam at a pressure of about 25-50 bar(a) and a temperature of about 225-265°C. This demand is usually comfortably met by using process heat in boilers, predominantly by cooling the reformed gas mixture that is produced in the steam reforming stage through heat exchange with water that is thus converted to steam. Excess high (i.e. at a pressure or greater than 50 bar(a)) or medium pressure steam that is not used in the process can valuably be exported from the process as low carbon medium or high pressure steam, preferably being superheated to be exported as superheated low carbon medium or high pressure steam. One existing such process, that provides for excess medium pressure steam export, involves subjecting a gaseous mixture comprising hydrocarbon and steam to pre-reforming and, thereafter, to autothermal reforming, to produce a reformed gas mixture. The reformed gas mixture is then cooled in a reformed gas boiler that raises high or medium pressure steam, and the cooled reformed gas mixture is subsequently subjected to a water-gas shift reaction in adiabatic reactors, sequential at high and low temperatures, with inter-stage cooling, to produce a hydrogen-enriched reformed gas mixture. The hydrogen-enriched reformed gas mixture is further cooled, raising low pressure steam, typically at a pressure of 3-6 bar(a) in exchangers downstream of the water-gas shift reaction. CO2 in the hydrogen-enriched reformed gas mixture is then removed using a reactive amine wash system, before purification of the gas mixture by pressure-swing absorption that produces purified hydrogen gas and a low-pressure hydrogen-lean CO containing tail gas. Some of the tail gas is used as fuel to superheat excess medium pressure steam that is raised in the process, for export, or as fuel in a fired pre-heater for the feed to the ATR. Some of the tail gas may be recycled to the feed or reforming, however the quantity that can be recycled may be limited by the presence of inert gases. To prevent the build-up of such inert gases, more of the tail gas may need to be purged from the process as a fuel. This can be accomplished by using some of the tail gas as fuel in a fired boiler to raise additional medium pressure steam for export. The amine wash system typically requires a large amount of low temperature heat (for example around 130°C), which demand is conveniently met by heat recovered, either directly or via low-pressure steam, in exchangers downstream of the water-gas shift reaction. The low-temperature heat demand of the amine wash system limits the extent to which the process can be adapted to produce more high or medium pressure steam for export or to reduce the demand for medium pressure steam raised in the process, since the heat recovered in the exchangers downstream of the water-gas shift reaction needs to be used in the amine wash system and is, in any event, too cold for the raising of medium pressure steam. Extra export steam could be created by using more of the tail gas in a fired boiler and superheater to create more export steam, but only at the cost of significantly reducing the process feedstock efficiency. Such a process is described in US2022 / 0194789. Another existing process, that improves on the abovementioned process, avoids the abovementioned demand for low-temperature heat in performing CO2 removal, by employing a cryogenic fractionation CO? removal system instead of an amine wash CO2 removal system. In such a cryogenic fractionation CO2 removal system, the cooled, dewatered hydrogen-enriched reformed gas is first purified using a first pressure-swing absorption step, thereby producing one or more hydrogen gas products of different purities and pressures, and a hydrogen lean CO2 containing tail gas. The tail gas is compressed, dried, chilled to cryogenic temperatures, and sent to a fractionation column, where CO2 is separated as a liquid product. The resulting CO2 lean gas, also containing CO, CH4 and H2, is further purified by a second pressure swing absorption step, producing one or more additional streams of purified hydrogen gas, additional CO2 rich gas, to be recycled to fractionation, and a residual low-pressure gas that is rich in CO and CH4. The residual low-pressure gas is compressed and recycled to the autothermal reformer. The cryogenic process needs power input to provide refrigeration to liquefy and separate CO2 rather than low temperature heat. By avoiding the demand for low-temperature heat in performing CO2 removal, there is excess low-pressure steam, raised in the exchangers downstream of the water-gas shift reaction, that is available for export. Such steam is, however, not as valuable as high or medium pressure steam. While the process with CO2 removal by fractionation does allow for conditions to be optimised to increase medium pressure steam that is available for export from the process, e.g. by reducing the steam to carbon ratio of the gaseous mixture comprising hydrocarbon and steam and therefore reducing the quantity of steam fed to the ATR, the magnitude of this increase is limited. A process including cryogenic CO2 removal is disclosed in US2023 / 0399227. It would be advantageous to increase the net high or medium pressure steam that is available for export from such processes while maintaining a high feedstock efficiency. Summary of the Invention According to the invention, there is provided a process for the production of hydrogen, the process comprising - subjecting a gaseous mixture comprising hydrocarbon and steam, and having a steam to carbon ratio in a range of from 0.4 to 1.8, to autothermal reforming in an autothermal reformer, thus producing a reformed gas mixture; cooling the reformed gas mixture in a reformed gas boiler through heat exchange with water as a coolant, thereby producing a cooled reformed gas mixture and raising high or medium pressure steam; subjecting the cooled reformed gas mixture to a water-gas shift reaction under isothermal conditions in an isothermal steam raising water-gas shift reactor, at a water-gas shift reaction temperature of at least 230°C but below 320°C using a suitably stable and active water-gas shift catalyst and water as a coolant, thus producing a hydrogen-enriched reformed gas mixture and raising medium pressure steam; purifying the hydrogen-enriched reformed gas mixture in a purifier and producing at least a purified hydrogen gas and a hydrogen lean CO2 containing tail gas; and separating CO2 from the hydrogen lean CO2 containing tail gas by subjecting the hydrogen lean CO2 containing tail gas to cryogenic fractionation by cooling it to cryogenic temperatures and recovering a liquid CO2 product and a CO2 lean gas. In this specification, the term “medium pressure steam” is used to mean steam at a temperature that is in a range of from about 225°C to about 265°C and at a pressure that is in a range of from about 25 bar(a) to about 50 bar(a), for example steam that is at a temperature of about 255°C and at a pressure of about 43 bar(a). The term “high pressure steam” is used to mean steam at a temperature of above about 265°C and a pressure above about 50 bar(a). As discussed in the background to the invention, above, in a low carbon hydrogen production process that employs high and low temperature water-gas shift reaction steps, the use of a cryogenic fractionation CO2 removal system, instead of an amine wash CO2 removal system, provides an excess of low-pressure steam that is raised using heat from the water-gas shift reaction that is available for export from the process. Such an excess essentially arises from a reduction in, or removal of, the process demand for low temperature heat by replacement of an amine wash CO2 removal system, which demands such heat, with a cryogenic fractionation CO2 removal system, which does not. Additionally, the applicant has realised that the steam requirements to protect high temperature shift catalysts and the low conversion of adiabatic water-gas shift reactors, which typically achieve single-pass conversions of 80% or less unless a low temperature water-gas shift step is used, constrains the amount of high or medium pressure steam that can be made available for export from the process. Low single-pass conversion levels can be compensated by increasing the size of recycles, but that typically leads to feedstock inefficiency because of the need to purge inert gases from the recycle loops and may also require increased equipment sizes to handle the larger flowrates associated with larger recycles. The process of the invention, by contrast, can achieve both high feedstock efficiency and produce valuable quantities of high or medium pressure steam. In the present invention, it was found that a reduction in, or absence of, the process demand for low-temperature heat, can advantageously be leveraged to reduce the process demand for more valuable medium pressure steam and increase the amount of high or medium pressure steam that can be raised in the process, thus increasing the net medium pressure steam that is available for export from the process. More specifically, this was found to be made possible by performing the water-gas shift reaction under isothermal conditions at a temperature of at least 230°C but below 320°C using water as coolant, instead of in sequential high and low temperature water-gas shift steps. Such an approach reduces the required steam content of the gas entering the water-gas shift reaction, since operating at a temperature of at least 230°C but below 320°C allows for a water-gas shift catalyst to be used that does not require additional steam to be added to avoid over-reduction of the catalyst by CO and H2. Furthermore, in operating isothermally using water as coolant, all of the heat from the water-gas shift reaction can conveniently be used to raise medium pressure steam in the water-gas shift reactor, which is not possible when performing the water-gas shift reaction sequentially at high and low temperatures respectively. The combined reduction in the steam requirement of the water-gas shift reaction and the raising of medium pressure steam in performing the water-gas shift reaction, results in an advantageous overall increase in the net high or medium pressure steam that is available for export from the process of the invention, optionally as superheated high or medium pressure steam as described below. The skilled person will appreciate that, when the term “isothermal” is used herein in relation to the water-gas shift reaction and reactor, this is because the temperature profile in the catalyst is constrained by heat transfer to the boiling water, which is at almost constant temperature. The inlet temperature of the cooled reformed gas to the water-gas shift reactor may be lower or higher than the boiling water temperature and there may be a small increase or decrease in gas temperature between an inlet and an outlet of the water-gas shift reactor, so that the temperature of the hydrogen-enriched reformed gas stream at the outlet of the water-gas shift reactor may be between 0°C and 40°C higher than the boiling water temperature. To operate isothermally, the water-gas shift reactor would be configured to provide for heat exchange in the reactor such that the water-gas shift reaction in the water-gas shift catalyst bed occurs in contact with heat exchange surfaces. Thus, in such a reactor, the water-gas shift cataiyst may be provided in tubes surrounded by coolant, or coolant may be provided in tubes surrounded by the catalyst. A suitable reactor may, for example, be an axial flow reactor in which the water-gas shift catalyst is provided inside the tubes of the reactor and coolant is provided outside the tubes, or an axial or radial flow reactor in which coolant is provided inside the tubes of the reactor and the water-gas shift catalyst is provided outside the tubes. ‘Axial’ flow defines the predominant gas flow in a cylindrical reactor as being parallel to the cylindrical axis, while ‘radial’ flow defines the predominant gas flow in a cylindrical reactor as being in a radial direction, either passing from the outside cylindrical wall of the reactor to an inner collector positioned at the central, cylindrical axis or vice versa. Radial flow reactors may advantageously have a low gas side pressure drop. An example of an isothermal steam raising reactor is described in Operation of the ICI Leading Concept Ammonia Process, I R Barton and K J Elkins, p 4 and 9. The isothermal steam-raising water-gas shift reactor may be coupled with a steam drum. The water and steam mixture, which is produced in the reactor, may then be conveniently separated into steam and boiling water at its bubble point within the drum. Furthermore, makeup boiler feed water to replenish the water removed through steam raising can either be added to the steam drum or to a circulating coolant water line to the isothermal steam raising water-gas shift reactor. This make-up boiler feed water has preferably been heated, for example by process streams, so that it is already at, or close to, such as within 20°C of, its boiling point. The water that is used as the coolant in the isothermal steam raising water-gas shift reactor is preferably passed from the steam drum to the isothermal steam raising water-gas shift reactor. One option is to use a pump to supply the motive power. Another option is to use natural circulation, whereby a density difference (brought about by reduced density of mixed steam and boiling water leaving the isothermal steam raising water-gas shift reactor, compared to the water entering the isothermal steam raising water-gas shift reactor) causes water to flow through the isothermal steam raising water-gas shift reactor. Preferably, the water that is used as coolant in the isothermal steam raising water-gas shift reactor is a mixture of water at its bubble point and make-up boiler feed water. At a shell inlet of the isothermal steam raising water-gas shift reactor the water is preferably at a temperature that is in a range of from about 0°C to 20°C below the temperature at which water boils at the pressure at the shell inlet. The same temperature range may apply where the cooling medium is passed through tubes disposed in a fixed bed of the catalyst The water that is used as coolant in the water-gas shift reactor may be pressurised water, preferably pressurised boiling water. Preferably, as the water is heated it reaches its boiling point and then an increasing fraction becomes steam, such that the water leaves the water-gas shift reactor as a two-phase mixture of medium pressure steam and water. Preferably, the water that is used as coolant in the water-gas shift reactor may, typically at an exit therefor from the reactor, be at a temperature that is in a range of from about 225°C to about 265°C and at a pressure that is in a range of from about 25 bar(a) to about 50 bar(a). The water-gas shift reaction taking place in the water-gas shift reactor is exothermic, and there would therefore continuously be heat available for transfer to the coolant while the reaction is taking place. Such transfer would, in accordance with the invention, at least partially vaporise the coolant, thus producing medium pressure steam. The process may include feeding the cooled reformed gas mixture to the isothermal water-gas shift reactor at a temperature that approximates the temperature at which the water-gas shift reaction takes place. Typically, the reformed gas mixture would leave the autothermal reformer at a much higher temperature than the temperature at which the water-gas shift reaction would take place, e.g. at a temperature in a range of from about 900°C to about 1100°C. The process therefore includes cooling the reformed gas mixture that leaves the autothermal reformer, before feeding the reformed gas mixture to the water-gas shift reactor. Such cooling may be to a temperature that approximates the temperature at which the water-gas shift reaction takes place. Cooling the reformed gas mixture is performed in a reformed gas boiler, through heat exchange with water as a coolant, raising high or medium pressure steam. The reformed gas mixture may optionally also be further cooled by providing heat to preheat other streams such as boiler feed water or the gaseous mixture comprising hydrocarbon and steam). The water-gas shift catalyst would typically not be a high-temperature water-gas shift catalyst. For example, it is preferred that the water-gas shift catalyst is not an iron-based catalyst. More typically, the water-gas shift catalyst would be a medium- or low-temperature water-gas shift catalyst, for example a copper-based catalyst, such as a copper / zinc oxide / alumina catalyst. Such a catalyst is typically reduced from its oxide to metallic form before use, hence it is not affected by the reducing nature of the reformed gas in the same way as an iron-based catalyst and it can advantageously operate at lower temperatures. The water-gas shift catalyst would typically be provided in the water-gas shift reactor in a conventional configuration, i.e. in particulate format as a bed of catalyst particles, either inside of the tubes of the reactor or outside of the tubes of the reactor, if the reactor is configured as described above. In some arrangements a low-temperature water-gas shift reactor is included downstream of the isothermal water-gas shift reactor to increase the hydrogen content of the shifted gas stream. The process may include combining medium pressure steam, raised in the process, with the reformed gas mixture, typically with the cooled reformed gas mixture, upstream of the water-gas shift reactor. Such medium pressure steam may be obtained from the high or medium pressure steam that is raised in the reformed gas boiler and / or from the medium pressure steam that is raised in the water-gas shift reactor. Such combining may advantageously increase the conversion of CO in the isothermal water-gas shift reaction, which may advantageously reduce the amount of CO that needs to be recycled and the power consumption of the process. It may also advantageously reduce the exotherm in the isothermal steam raising water-gas shift reactor. It will be appreciated that such combining may also reduce the amount of medium pressure steam which can be exported from the process, however, the combining may still be advantageous overall in some embodiments according to feedstock, steam and power values and the specific requirement for steam export in a given project. The process may further include producing the gaseous mixture comprising hydrocarbon and steam by using high or, preferably, medium pressure steam that is raised in the process, for example in the isothermal steam raising water-gas shift reactor or in the reformed gas boiler. Such medium pressure steam is preferably fully or partly obtained from the medium pressure steam that is raised in the water-gas shift reactor. If the steam raised in the isothermal steam raising water-gas shift reactor is insufficient to provide the required steam to carbon ratio for the autothermal reforming, it is preferably augmented by using a part of the steam produced in the reformed gas boiler. The steam may be produced in the reformed gas boiler at medium pressure, or it may be produced at high pressure in which case its pressure is preferably reduced before use as process steam to provide the required steam to carbon ratio for the autothermal reforming. More specifically, in producing the gaseous mixture comprising hydrocarbon and steam, a hydrocarbon feed stream may be combined with medium pressure steam from the high or, preferably, medium pressure steam that is raised in the process, more specifically in the reformed gas boiler and / or, preferably, in the water-gas shift reactor. Steam that is added to the gaseous mixture comprising hydrocarbon and steam, or upstream of the isothermal water-gas shift reactor, may either be saturated steam or superheated steam. For instance, medium pressure steam raised in the isothermal water-gas shift reactor may be superheated before mixing with the hydrocarbon feed. If export steam from the process is superheated, superheating for process steam may preferably be carried out in a coil in the same fired heater as the superheating for the export steam. In some arrangements, steam raised in the process, preferably high-pressure steam raised by a reformed gas boiler, may be superheated and sent to a steam turbine generator (STG) to generate power and so reduce the power requirement from the process. The steam stream recovered from the STG may be a superheated medium pressure steam that that may be either exported directly or mixed with other steam streams, e.g. from the isothermal reactor steam drum, before being exported. In some embodiments the cooled reformed gas mixture may be subject to an additional, high-temperature, water-gas shift step upstream of the isothermal water-gas shift reactor. However, such a step is preferably not present. The hydrocarbon may comprise any gaseous or low boiling hydrocarbon, such as natural gas, associated gas, LPG, petroleum distillate, diesel, naphtha, or hydrocarbon-containing offgases from chemical processes, such as a refinery off-gas or a pre-reformed gas. Such gases may contain major or minor impurities, which could negatively affect the process efficiency or could have a deleterious effect on the catalysts utilised in the process. Such gases may be subject to upstream bulk purification steps, however, even after such bulk purification, it may still be advantageous to employ catalytic and absorption steps to remove trace impurities, such as sulphur compounds, to provide a purified hydrocarbon feed to the process and preferably to also react bulk levels of unsaturated hydrocarbons (e.g. ethylene or propylene) with hydrogen to form saturated hydrocarbons. The hydrocarbon feed or mixture may be subjected to desulphurisation comprising hydrodesulphurisation, for example using CoMo or NiMo catalysts, and absorption of hydrogen sulphide using a suitable hydrogen sulphide adsorbent, e.g. a zinc oxide adsorbent A hydrodesulphurisation catalyst may also be suitable to hydrogenate olefins present, provided that there is sufficient hydrogen present. Hydrodesulphurisation may be carried out before, or preferably after, compression of the hydrocarbon. An ultra-purification adsorbent may usefully be used downstream of the hydrogen sulphide adsorbent to protect the steam reforming catalyst further. Suitable ultra-purification adsorbents may comprise copper-zinc oxide / alumina materials and copper-nickel-zinc oxide / alumina materials. To facilitate hydrodesulphurisation, hydrogen is preferably present in the hydrocarbon feed or mixture. It may already be present in sufficient quantity as part of one or more of the individual hydrocarbon feeds, or it may be added via a recycled hydrogen containing process stream. The amount of hydrogen in the resulting mixed gas stream may be in the range 1-20 vol%, but is preferably in the range 1-10 vol%, more preferably in the range 1-5 vol% on a dry gas basis. This may be the amount of hydrogen at the exit of the hydrodesulphurisation vessel. If olefins are present in the feed or mixture, the hydrogen content is preferably augmented to provide at least enough additional hydrogen for stoichiometric conversion of the olefin content. The hydrocarbon feed or mixture may be pre-heated in one or more stages. It may conveniently be pre-heated after compression and before desulphurisation. Where the hydrocarbon feed or mixture is desulphurised, it may after desulphurisation be further heated before being mixed with steam. Upstream of subjecting the, optionally preheated, gaseous mixture comprising hydrocarbon and steam to autothermal reforming, the process may include subjecting the mixture to adiabatic pre-reforming in a pre-pre-reformer. The, optionally pre-heated, gaseous mixture comprising hydrocarbon and steam is either fed to the autothermal reformer or preferably subjected to a step of adiabatic steam reforming in one or more pre-reformer vessels (“pre-reforming”) before being subjected to autothermal reforming in the autothermal reformer. The pre-reformer and autothermal reformer are, in such an embodiment, preferably operated in series. In pre-reforming, the gaseous mixture comprising hydrocarbon and steam is preferably passed at an inlet temperature in the range of 300-650°C, preferably 380-450°C, adiabatically through a bed of a steam reforming catalyst, usually a steam reforming catalyst having a high nickel content, for example above 40% by weight. During such an adiabatic pre-reforming step, hydrocarbons higher than methane (e.g. ethane, propane) react with steam to give a mixture of methane, carbon oxides and hydrogen. The use of such an adiabatic steam reforming step, commonly termed pre-reforming, is desirable to ensure that the feed to the autothermal reformer contains no hydrocarbons higher than methane and also contains some hydrogen. The pre-reformed gas, which preferably comprises methane, hydrogen, steam and carbon oxides, is then fed to the autothermal reformer in which it is subjected to autothermal reforming. According to the design conditions of the ATR, this advantageously reduces or eliminates the potential to form soot in the combustion of the gaseous mixture comprising hydrocarbon and steam with oxygen. Preferably, all of the pre-reformed gaseous mixture comprising hydrocarbon and steam is fed to the autothermal reformer. If desired, the temperature and / or pressure of the pre-reformed gaseous mixture comprising hydrocarbon and steam may be adjusted before feeding it to the autothermal reformer. In a preferred embodiment, the pre-reformed gaseous mixture comprising hydrocarbon and steam recovered from the adiabatic pre-reforming step is heated before feeding it to the autothermal reformer, for example by passing it through a fired heater fuelled by at least a portion of a hydrogen-rich fuel gas produced in the process. This may be the same fired heater used to pre-heat the hydrocarbon and / or to superheat the high or medium pressure steam. Desirably, the pre-reformed gas is heated to 450-700°C, preferably 550-650°C. The autothermal reformer may comprise a burner disposed at the top of the reformer, to which the pre-reformed gaseous mixture comprising hydrocarbon and steam and an oxygen-rich gas are fed, a combustion zone beneath the burner through which a flame extends, and a fixed bed of particulate steam reforming catalyst disposed below the combustion zone. In autothermal reforming, the heat for the endothermic steam reforming reactions is provided by combustion of a portion of hydrocarbon in the pre-reformed gaseous mixture comprising hydrocarbon and steam in an autothermal reformer fired heater. The pre-reformed gaseous mixture comprising hydrocarbon and steam is typically fed to the top of the reformer and the oxygen-rich gas is fed to the burner, mixing and combustion occur downstream of the burner generating a heated gas mixture the composition of which is brought to equilibrium as it passes through the steam reforming catalyst. The autothermal steam reforming catalyst may comprise nickel supported on a refractory support such as rings or pellets of calcium aluminate, magnesium aluminate, alumina, titania, zirconia and the like. In a preferred embodiment, the autothermal steam reforming catalyst comprises a layer of a catalyst comprising Ni and / or Ru on zirconia over a bed of a Ni on alumina catalyst to reduce catalyst support volatilisation that can result in deterioration in performance of the autothermal reformer. The oxygen-rich gas may comprise at least 90% vol O2, still more preferably at least 95% vol □2, most preferably at least 98% vol O2, or at least 99% vol O2, e.g. a pure oxygen gas stream, which may be obtained using a vacuum pressure swing adsorption (VPSA) unit or, more typically, an air separation unit (ASU). When an ASU is used, the ASU may be electrically driven and is desirably driven using renewable electricity to further minimise CO2 emissions. Preferably, the amount of oxygen added is such that the reformed gas mixture leaves the autothermal reforming catalyst at a temperature in the range 900-1100°C. The gaseous mixture comprising hydrocarbon and steam may be formed from one hydrocarbon containing feed or it may be formed from a mixture of various different hydrocarbon containing feeds. Examples include refinery off gases from treating fossil and renewable feedstocks (on refineries), cracker off gases (on petrochemical facilities), LPG, renewable natural gas, biogases, LNG and natural gas. For example, the process may be employed in a refinery and the gaseous mixture comprising hydrocarbon and steam then comprises refinery off gases. The hydrocarbon feeds preferably comprise, in respect of the hydrocarbon portion thereof, methane or methane with a lesser quantity of higher hydrocarbons. The hydrocarbon feeds may comprise other components, in particular hydrogen. Preferably the hydrogen content of the hydrocarbon feed is sufficient to meet the process requirement for hydrodesulphurisation and any required olefin conversion. In such embodiments it may be that the hydrocarbon feed is advantageously not mixed with any recycle gas or recycled, purified hydrogen product upstream of the hydrodesulphurisation. In other words, preferably the hydrocarbon feed, preferably refinery off gas, comprises hydrogen and no hydrogen is added upstream of the hydrodesulphurisation. The hydrocarbon feed may be at a pressure in the range 10-100 bar(a). The pressure of the hydrocarbon may usefully govern the pressure throughout the process. Operating pressure is preferably in the range 15-50 bar(a), more preferably 25-50 bar(a) as this advantageously provides optimum overall economics for the process. The hydrocarbon feed and / or the offgases may be compressed, if necessary, to the desired pressure using conventional compression equipment. Introducing steam into the hydrocarbon feed stream, to produce the gaseous mixture comprising hydrocarbon and steam, may be performed by direct injection of steam, which may be steam that is produced as hereinbefore described, i.e. in the reformed gas boiler and / or in the isothermal steam raising water-gas shift reactor. The amount of steam introduced is sufficient to give a steam to carbon ratio (defined as the steam to hydrocarbon carbon ratio at the inlet to the autothermal reformer) in the abovementioned range of from 0.4 to 1.8, preferably 0.4 to 1.6, preferably 0.6 to 1.4 and more preferably 0.8 to 1.2. The medium pressure steam demand to produce the gaseous mixture comprising hydrocarbon and steam, and to perform the water-gas shift reaction, would typically, as a feature of the invention, be met fully by the high or medium pressure steam that is generated in the reformed gas boiler and / or the medium pressure steam that is generated in the isothermal steam raising water-gas shift reactor, and would leave an excess of such high or medium pressure steam. At least a major portion, e.g. above 70% vol or 90 vol%, or all, of the excess medium pressure steam that remains after utilisation of medium pressure steam in producing the gaseous mixture comprising hydrocarbon and steam, may be exported from the process, i.e. is not utilised in the process. It will be appreciated that there are often a number of other minor flow uses of medium pressure steam within the hydrogen production process. Examples include (i) preheating of feed or regeneration gas streams and (ii) stripping of organic contaminants from process condensate streams. A small proportion of the excess medium pressure steam is preferably therefore used for such minor flow uses, with the remainder being exported. The process preferably includes superheating the excess high or medium pressure steam that would be exported from the process, such that the exported excess high or medium pressure steam is superheated high or medium pressure steam. Superheating the excess high or medium pressure steam that would be exported from the process may include using purified hydrogen gas and / or an impure hydrogen fuel gas, produced in the purifier, as a fuel gas in a fired heater. For example, an impure hydrogen fuel gas produced in the purifier may be combusted in a fired heater that comprises a superheater coil to superheat the high or medium pressure steam using heat produced by such combustion. Preferably the superheating uses a hydrogen containing gas from the process that contains higher levels of inerts compared to the purified hydrogen gas and / or is only available at a lower pressure than the purified hydrogen gas. The hydrogen fuel may be supplemented with another fuel source, e.g. a portion of the hydrocarbon feed, if desired. The process may include cooling the hydrogen-enriched reformed gas mixture from the water-gas shift reactor and separating condensed water therefrom, to provide a de-watered hydrogen-enriched reformed gas mixture. Thus, it would be the de-watered hydrogen-enriched reformed gas mixture that would be subjected to purification. Cooling of the hydrogen-enriched reformed gas mixture in order to produce the de-watered hydrogen enriched reformed gas mixture, may involve cooling it to a temperature below its dew point, so that the steam content thereof condenses to produce a liquid water condensate, which is then separated from the hydrogen-enriched reformed gas mixture in a gas-liquid separator, to produce the de-watered hydrogen enriched reformed gas mixture. To cool the hydrogen-enriched reformed gas mixture, any coolant may be used in one or more stages. Preferably, heat from cooling is used to heat process water streams (such as boiler feed water, process condensate and demineralised water), which are subsequently used in producing steam, for example in the reformed gas boiler and / or in the isothermal water-gas shift reactor. Cooling may also be accomplished by rejection of heat to the atmosphere, for example by using an air cooler or by using cooling water. In separating the liquid water condensate from the hydrogen-enriched reformed gas mixture to produce the de-watered hydrogen enriched reformed gas mixture, one or more, preferably two or three, stages of condensate separation may be used. If desired, a portion or all of the condensate, preferably treated as described below, may be used to generate steam for the process, e.g. in the reformed gas boiler and / or in the water-gas shift reactor. Preferably, some or all of the condensate is treated by contacting with medium pressure steam to volatilise compounds dissolved in the condensate. One such dissolved compound may be methanol, which may form as a byproduct in the water-gas shift reactor. The medium pressure steam, together with the volatilised compounds may then be added to the hydrocarbon feed. In this way, organic compounds, dissolved process gases and other impurities in the condensate may be returned to the process and thus increase the process efficiency and reduce the burden on any aqueous effluent treatment. The treated condensate may be used, as boiler feed water, for the generating of steam. Any condensate not used to generate steam may be sent to water treatment as effluent. If desired, the de-watered hydrogen-enriched reformed gas mixture may be treated by washing it with a pure water stream. Such washing may remove trace gas impurities that have formed in the reforming or water-gas shift sections and which are not desirable in downstream operations. Such trace gas impurities may include methanol, ammonia, or methylamines. The purifier may suitably comprise a hydrogen selective membrane unit, a temperature swing adsorption system, or a pressure swing adsorption system, or a combination thereof. Such systems are commercially available and are not described in detail. The purifier preferably comprises one or more pressure swing adsorption units. Such systems comprise regenerable porous adsorbent materials that selectively absorb gases other than hydrogen and thereby produce products enriched in hydrogen. The purified hydrogen gas may have a purity of 95% or higher. The purified hydrogen gas produced by the purifier preferably has a purity greater than 98 vol%, more preferably greater than 99.5 vol%, even more preferably greater than 99.9 vol% or yet more preferably greater than 99.99 vol%. An impure hydrogen fuel gas, as has been referenced earlier herein, may also be produced in the purifier. The impure hydrogen fuel gas is preferably enriched in inert gas content (such as nitrogen and argon) and optionally comprises such content of CO, CH4 and CO2 as is acceptable to allow a combustion product of the impure hydrogen fuel gas to pass to atmosphere while still meeting the carbon footprint specification of the low carbon hydrogen process or the target carbon capture efficiency. The process may include compressing the purified hydrogen gas, for example using an electrically driven compressor, preferably being powered by renewable electricity. The purified hydrogen gas may be used in downstream power or heating processes, for example by using it as fuel in a gas turbine or by injecting it into a domestic or industrial networked gas piping system. The purified hydrogen gas may also be used as a fuel for vehicles using fuel cells. The purified hydrogen gas can also be stored in a convenient form as an energy carrier and optionally transported. For example, it can be stored as compressed hydrogen or liquid hydrogen, or it can be reacted with another compound (such as reaction with benzene to form cyclohexane) to enable easier transport to another geographic location. The purified hydrogen gas may alternatively be used in a downstream chemical synthesis process. Thus, the purified hydrogen gas may be used to produce, for example, ammonia by reaction with nitrogen in an ammonia synthesis unit. It is also possible for the purified hydrogen gas to be used to upgrade hydrocarbons, e.g. by hydro-treating or hydro-cracking hydrocarbons in a hydrocarbon refinery, or in any other process where pure hydrogen may be used. Such hydrocarbons can be fossil fuel derived or may be biogenic, such as in the hydrogenation of vegetable oils. In such cases, it could be advantageous to use the off gas and / or a low value stream from the hydrocarbon upgrading process as one or more of the hydrocarbon feeds to the process of this invention. A portion of the purified hydrogen gas may be recycled to the hydrocarbon feed stream, if desired for hydrodesulphurisation or hydrogenation of unsaturated compounds. The composition of the impure hydrogen fuel gas that is produced in the purifier depends on the extent of the purification in the purifier. The impure hydrogen fuel gas may, for example, comprise 80-90 vol% hydrogen, with the balance comprising methane, carbon monoxide, carbon dioxide and inert gases. The methane content may be in the range 1-5vol%, preferably 2-5 vol%. The carbon monoxide content may be in the range 2-10 vol%, preferably 2-8 vol%. The carbon dioxide content may be in the range 0-1.5 vol%. There may also be traces of steam and inerts (such as nitrogen and argon) in the range 0-5 vol%. The production of the impure hydrogen fuel gas may also usefully purge inert species from the process, including gaseous nitrogen and argon. The hydrogen lean CO2 containing tail gas that is produced in the purifier and from which CO2 is separated by cryogenic fractionation, would typically comprise CO and CH4, in addition to hydrogen and CO2 and at least a fraction of the inerts content (typically nitrogen and argon) in the feed to the purifier. Thus, the CO? lean gas would typically comprise a major portion of CO and CH4 and a minor portion of hydrogen and CO2. Upstream of the cryogenic fractionation, the hydrogen lean CO2 containing tail gas may be compressed and dried. This advantageously increases the pressure sufficiently high to enable CO2 to condense at a viable cryogenic temperature and advantageously reduces the water content so that solid water ice does not form in the process. Thus, it would be compressed and dried hydrogen lean CO2 containing waste gas that would be subjected to cryogenic fractionation. The process may also include - further purifying the CO2 lean gas, produced by cryogenic fractionation, in a further purifier and producing a CO and CH4 rich gas with a lower total flowrate of inert gases than in the hydrogen enriched reformed gas; and recycling the CO and CH4 rich gas to form part of the gaseous mixture comprising the hydrocarbon and steam, for example by combining it with the hydrocarbon feed before or after purification thereof, if the hydrocarbon feed is purified before adding steam. An additional purified hydrogen gas and / or a CO2 rich gas may also be produced in the further purifier. Purifying the CO2 lean gas in the further purifier may, as in the case of purifying the de-watered hydrogen-enriched reformed gas, be suitably performed using a membrane system, a temperature swing adsorption system, or a pressure swing adsorption system, or a combination thereof. Preferably, the total flowrate of inert gases in the CO and CH4 rich gas is at least 10%, or 15%, or 20%, or 30% lower than the total flowrate of inert gases in the hydrogen enriched reformed gas. The reduction in the total flowrate of inert gases may permit recycling of a larger stream of CO and CH4 rich gas, thus reducing the conversion required in the reforming and / or water-gas shift reactions and increasing the process efficiency and net high or medium pressure steam export. The additional purified hydrogen gas that is optionally produced in purifying the CO2 lean waste gas may be combined with the purified hydrogen gas produced in purifying the de-watered hydrogen-enriched reformed gas or it may form an additional hydrogen product. The CO2 rich gas that is optionally produced in further purifying the CO2 lean gas may be recycled to mix with the hydrogen lean CO2 containing tail gas to be subjected to CO2 separation through cryogenic fractionation. Detailed Description of an Embodiment of the Invention The invention is now described by way of illustrative example only, with reference to the accompanying diagrammatic drawing, in which Figure 1 is a diagrammatic flowsheet of one embodiment of the process of the invention, Figure 2 is a diagrammatic flowsheet of another embodiment of the process of the invention, and Figure 3 is a diagrammatic flowsheet of another embodiment of the process of the invention. It will be understood by those skilled in the art that the drawings are diagrammatic and that further items of equipment such as heat exchangers utilising process heat from process streams, boilers, reflux drums, pumps, vacuum pumps, temperature sensors, pressure sensors, pressure relief valves, control valves, flow controllers, level controllers, holding tanks, storage tanks, and the like may be required in a commercial plant. The provision of such items of equipment forms no part of the present invention and would be provided in accordance with conventional chemical engineering practice. In the drawing of Figure 1, reference numeral 10 generally indicates a process for the production of hydrogen, in accordance with the invention. In the process 10, a natural gas stream comprising >85 vol% methane is fed along feed line 12 to a gas purification stage 14. Upstream of the gas purification stage 14, the natural gas is combined with a part of a recycled CO and CH4 rich gas fed along recycle line 98, in order to give, in this embodiment, a hydrogen concentration leaving the downstream gas purification stage 14 of 2 vol%. The recycled CO and CH4 rich gas is produced in the process 10 as hereinafter described. In the gas purification stage 14, the natural gas is subjected to desulphurisation, e.g. hydrodesulphurisation in a hydrodesulphurisation vessel (not illustrated) containing a bed of hydrodesulphurisation catalyst. Hydrodesulphurisation converts organic sulphur compounds in the natural gas to hydrogen sulphide. The natural gas, thus treated, is then subjected to hydrogen sulphide and trace sulphur compound removal in desulphurisation and ultra-purification vessels (not illustrated) containing a bed of zinc oxide adsorbent and a bed of copper-zinc-alumina ultra-purification adsorbent respectively. The desulphurised natural gas is then passed along line 15 to a reforming stage 16, being combined with the remainder of the recycled CO and CH4 rich gas fed along recycle line 100 upstream of the reforming stage 16. In the steam reforming stage 16, the desulphurised natural gas is combined with medium pressure steam that is raised in the process 10, as hereinafter described, and that is provided along line 18, to provide a gaseous mixture of hydrocarbon (predominantly methane) and steam. Steam addition is set in this embodiment to give a steam to carbon ratio, as measured at the inlet of the downstream autothermal reformer, of 1.1. The gaseous mixture is then subjected to adiabatic pre-reforming, followed by autothermal reforming, in the reforming stage 16. Adiabatic pre-reforming is performed in a pre-reformer (not illustrated), in the reforming stage 16, which contains a bed of pelleted nickel-based steam reforming catalyst. As the gaseous mixture passes over the pre-reforming catalyst, higher hydrocarbons are converted to methane and methane is partially steam reformed to produce a pre-reformed gas mixture containing hydrogen. The pre-reformed gas mixture is then fed from the pre-reformer to a fired heater (not illustrated) where it is heated in a coil to the required inlet temperature for autothermal reforming. The heated pre-reformed gas mixture is then fed from the fired heater to the burner region of an autothermal reformer (not illustrated) in the reforming stage 16, where it is partially combusted with oxygen that is provided along line 20 and that has been produced in an air separation unit (not illustrated) and has been pre-heated in a heat exchanger (not illustrated). An impure hydrogen fuel, produced in the process 10, as hereinafter described, may optionally be fed to the autothermal reformer, along line 22, to aid combustion. The hot combusted gas mixture is brought towards equilibrium, in this embodiment at a pressure of about 30 bar(a) and a temperature of 1000°C, exit the autothermal reformer over a fixed bed of a pelleted nickel-based secondary reforming catalyst disposed below the combustion zone in the autothermal reformer. The resulting hot reformed gas mixture is passed from the autothermal reforming stage 16, and more specifically from the autothermal reformer, along line 24, to a reformed gas boiler stage 26. The reformed gas boiler stage 26 in this embodiment comprises a high-pressure steam-raising boiler (not illustrated) coupled to a steam drum (not illustrated). More specifically, the reformed gas mixture is passed to the tube-side of the reformed gas boiler in the boiler stage 26. Through heat exchange, the hot reformed gas mixture then vaporises boiler feed water that is fed to the coupled boiler and steam drum, along line 28, thus raising high pressure steam. High pressure steam that is recovered from the reformed gas boiler stage 26, and more specifically from the steam drum, along line 30, is divided, and a small fraction of it is depressurised and used in the process 10 as hereinafter described, while some of it is passed as stream 31 to superheater 66 and then exported from the process 10 in stream 68, as excess high pressure steam. The reformed gas mixture from the reforming stage 16 is cooled as it passes through the reformed gas boiler stage 26, and thus a cooled reformed gas mixture is obtained from the reformed gas boiler stage 26. Further cooling of the reformed gas mixture is conducted in this embodiment in heat exchangers to preheat boiler feed water and the gaseous mixture comprising hydrocarbon and steam (not illustrated) to a temperature that approximates the temperature at which the reformed gas mixture will subsequently be subjected to the isothermal water-gas shift reaction, typically being about 245°C. The cooled reformed gas mixture is then passed to an isothermal steam-raising water-gas shift reactor, or vessel, 32 along line 34, to be subjected to a water-gas shift reaction under isothermal conditions using pressurised boiling water as coolant, to increase the H2 content thereof by water-gas shift of CO to CO2 and H2 in the presence of steam. Upstream of the water-gas shift reactor 32, the cooled reformed gas mixture in line 34 may optionally be combined with additional medium pressure steam, which is introduced along line 36. Such additional medium pressure steam may be supplied for the purpose of increasing the shift of CO to H? and to reduce the CO that needs to be recycled to the feed stream that is ultimately subjected to autothermal reforming. In this embodiment the water-gas shift reactor 32 is an axial flow reactor, coupled with a steam drum, in which a water-gas shift catalyst is provided inside the tubes of the reactor and boiling water, as a coolant, is provided outside the tubes. Preheated boiler feed water is supplied to the coupled steam drum and reactor 32 along line 38. The preheated boiler feed water is at a pressure of about 40 bar(a) and, after mixing with water in the stream drum, forms a coolant, which circulates to the water-gas shift reactor shell at a temperature of about 230°C. The water-gas shift reactor 32 contains a particulate bed of copper-based mediumtemperature shift catalyst. The water-gas shift reaction whereby the hydrogen content of the reformed gas mixture is increased, and the carbon monoxide thereof converted to carbon dioxide, occurs as the cooled reformed gas mixture passes through the bed. The water-gas shift reaction taking place in the water-gas shift reactor 32 is exothermic, and heat exchange with the boiling water coolant thus vaporises the water and raises medium pressure steam at about 35 bar(a) in the water-gas shift reactor 32, which is separated in the stream drum and withdrawn along line 40. At the conditions of the embodiment all of the medium pressure steam thus raised in the water-gas shift reactor is used in the process 10, e.g. in the autothermal reforming stage 16 to produce the gaseous mixture of hydrocarbon and steam as stream 18, or for other minor uses. It is a feature of the process 10 that the medium pressure steam demand of the process 10 is fully met by the high-pressure steam that is raised in the boiler stage 26 and the medium pressure steam raised in the water-gas shift reactor 32, leaving excess high pressure steam that is raised in the boiler stage 26, which is superheated and exported as described. The water-gas shift reaction in the water-gas shift reactor 32 produces a hydrogen-enriched reformed gas mixture, which is withdrawn from the reactor 32 along line 42. The hydrogen-enriched reformed gas mixture is then subjected to de-watering. More specifically, the hydrogen-enriched reformed gas mixture is cooled to condense out water in a heat exchangers 44. Heat from the cooling gas is used to preheat boiler feed water, hydrocarbon feed and demineralised water and then finally heat is rejected to atmosphere by cooling with cooling water. The condensed water is subsequently separated from the hydrogen-enriched reformed gas in a gas-liquid separator 48, to which the cooled hydrogen-enriched reformed gas mixture carrying condensed water is passed along line 46. The condensate is recovered from the separator 48 along line 50, and the hydrogen-enriched reformed gas mixture, as a de-watered hydrogen-enriched reformed gas mixture, is withdrawn from the separator 48 along line 52. The condensate is conveniently used in the process 10. For example, after steam stripping to remove contaminants such as methanol, it may be used in raising high pressure steam in the boiler stage 26 and / or medium pressure steam in the water-gas shift reactor 32and / or may be used to form part of the steam added to the reforming section 16 via line 18. The de-watered hydrogen-enriched reformed gas mixture is then passed, along line 52, to a purifier in the form of a pressure swing absorption unit 54. The pressure swing absorption unit 54 contains porous adsorbents that, through multiple pressurisations, equalisations, depressurisations and purges, absorbs non-hydrogen molecules, including carbon oxides and methane, contained in the de-watered hydrogen-enriched reformed gas mixture, thereby producing a purified hydrogen gas which is withdrawn along line 56. In this embodiment, the pressure swing absorption unit 54 is also configured to produce a low-pressure impure hydrogen fuel gas and a hydrogen lean CO2 containing tail gas. These are withdrawn from the pressure swing absorption unit 54 respectively along lines 58 and 60. The purified hydrogen gas is passed along line 56 to a compressor 62, in which the purified hydrogen gas is compressed, thus producing a compressed purified hydrogen gas product of the process 10, which is withdrawn along line 64. The impure hydrogen fuel gas is passed along line 58 to the superheater 66 previously referenced, where it is used as a fuel. More specifically, in the superheater 66 the impure hydrogen fuel gas is combusted to superheat excess high-pressure steam raised in and supplied from the reformed gas boiler stage 26, thereby to produce superheated high pressure steam that is withdrawn from the superheater 66 as a superheated high pressure steam product of the process 10, along line 68. Some of the impure hydrogen fuel is recovered from line 58 along line 70 upstream of the superheater 66. This may be fed as fuel to the fired heater in the reforming section 16. The hydrogen lean CO? containing tail gas is further treated for CO2 removal. More specifically, the hydrogen lean CO2 containing tail gas is passed, along line 60, to a compressor 72, which compresses it, and then, along line 73, to a dryer 74, in which it is dried. The dried and compressed hydrogen lean CO2 containing tail gas is then passed along line 76 to a fractionation system 78 comprising a refrigeration unit 80. In the fractionation system 78, the dried and compressed hydrogen lean CO2 containing tail gas is chilled and subjected to cryogenic fractionation thus producing liquid CO2, which is recovered as a liquid CO2 product of the process along line 82, and a CO2 lean gas, which is recovered along line 84. The CO2 lean gas is passed along line 84 to a further pressure swing absorption unit 86, which produces additional purified hydrogen gas, a CO2 rich gas, and CO and CH4 rich gas comprising mainly CO and CH4 and hydrogen. These are recovered from the pressure swing absorption unit 86 respectively along lines 88, 90, and 92. The additional purified hydrogen gas recovered along line 88 is combined with the purified hydrogen gas that is recovered from the pressure swing absorption unit 54 along line 56, to be subjected to compression by the compressor 62 and thus form part of the compressed purified hydrogen gas product of the process 10 that is withdrawn along line 64. The CO2 rich gas is recycled along line 90 and is combined with the hydrogen lean CO2 containing tail gas that is recovered from the pressure swing absorption unit 54 along line 60, to be resubjected to fractionation in the fractionation system 78. The CO and CPU rich gas that is recovered along line 92 is compressed in a compressor 94, to produce compressed CO and CH4 rich gas 96. Such compressed CO and CH4 rich gas 96 is recycled along line 98 to the natural gas in feed line 12 and along line 100 to the desulphurised natural gas in line 15. In accordance with the invention, in providing for removal of CO2 contained in the hydrogen-enriched reformed gas mixture by cryogenic fractionation as opposed to using a reactive wash, such as an amine wash, the low-temperature heat demand of the latter is avoided in the process 10, in contrast to existing processes that utilise reactive wash systems. In contrast to existing processes using cryogenic fractionation for CO2 removal, the process of the invention uses isothermal steam-raising water-gas shift reactor 32 to shift the CO at a high enough temperature that is operated to produce medium pressure steam. In addition, by operating the isothermal steam-raising water-gas shift reactor 32 at a moderate water-gas shift temperature, e.g. at about 245°C, non-iron-based catalysts can be used, which do not constrain the minimum steam content in the cooled reformed gas mixture entering the water-gas shift reactor 32. This allows less steam to be added upstream of the ATR and steam addition upstream of the water-gas shift reactor 32 may advantageously be avoided. The combined reduction in the steam content required in the reformed gas passing to the water-gas shift reaction and the raising of additional medium pressure steam using all of the heat from the isothermal water-gas shift reaction therefore increases the net high or medium pressure steam that may be exported from the process 10, as compared to a similar process that employs high and low temperature shifts combined with the cryogenic fractionation CO2 removal system. In Figure 2, like equipment has like numbers to Figure 1 and is not redescribed. In Figure 2, not all of the high-pressure steam raised in the reformed gas boiler 26 is passed to the superheater 66. Some is mixed with medium pressure steam 40 coming from the water-gas shift reaction 32, optionally also with some of the superheated steam from superheater 66. The resulting steam can be used in the process 10 as medium pressure steam 130 and for minor steam uses 131, such as heating of feed streams. Such an arrangement may be useful where the process steam requirements of the process 10 exceed the quantity of medium pressure steam 40 raised in the water-gas shift reaction section 32. Additionally, in the embodiment of Figure 2, a further impure hydrogen fuel gas 158 is recovered from the further pressure swing adsorption unit 86. This further impure hydrogen fuel gas 158 is combined with the impure hydrogen fuel gas 58 for use in the superheater 66 and the fired heater in the reforming section 16. In Figure 3, like equipment has like numbers to Figures 1 and 2 and is not redescribed. The further pressure swing adsorption unit 86 is replaced by a membrane unit 186 from which an impure hydrogen fuel gas 158 and a pressurised CO and CH4 rich gas 96 are recovered. Example The following is a computer simulation example of the invention, using the embodiment of Figure 3. The skilled person will appreciate that the use of simulations is established practice in the chemical engineering field, where it is not feasible to build multiple test plants. The simulation is set up to produce 100 kNm3 / hr of pure hydrogen product at >99.9% purity and 50 bar(a) while capturing -98% CO2 and maximising the export of superheated medium pressure steam. Medium pressure steam is raised in the reformed gas boiler. The purification steps and CO2 removal are according to figure 3 in US2023 / 0399227. The pressure swing adsorption unit 54, the CO2 separation 78 and the membrane unit 186 perform as set out in US2023 / 0399227. Table A Stream 12 20 24 42 52 82 bo 22 95 18 30 40 58 Mass flow (kg / hr) 29,830 33,610 106,1'77 106,177' 96,920 76,960 9,063 2,217' 8,311 34,002 83,820 28,060 70,510 Molar flow (kmci / hr) 1655.2 1049.0 7455.4 7407.0 5910.9 1756.9 4451.5 315.8 356.9 1886.9 4553.3 lob / .o 3914.1 Temp C'C) 20.0 20.0 1040.0 257.5 42.2 38.2 45.0 42.2 192.0 399.7 254.7 254.7 399.8 Pressure (bar(a)) 50.0 45.0 38.0 34.9 32.7 187.0 50 0 1.5 45.0 42.4 43.0 43 0 42 5 Components (mo! fraction) Methane 0.8900 0.0000 0.0079 0.0079 0.0085 0.0049 0.0000 0.0123 0.1297 0.0000 0.0000 0.0000 0.0000 Ethane 0.0700 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Propane 0.0100 0.0000 0 0000 0.0000 0.0000 0 0000 0.0000 0.0000 0.0000 0.0000 0 0000 0.0000 0.0000 n-butane 0.0010 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 n-Pentane 0.0001 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0 0000 0.0000 0.0000 0.0000 0.0000 Hydrogen 0.0000 0.0000 0.4690 0.6451 0.6913 0.0000 0.9995 0.8230 0.1547 0.0000 0.0000 0.0000 0.0000 CO2 0.0200 0.0000 0.0589 0.2389 0.2559 0.9909 0.0000 0.0235 0.0580 0.0000 0.0000 0.0000 0.0000 CO 0.0000 0.0000 0 2144 0.0329 0.0353 0 0037 0.0000 0.0853 0.5895 0.0000 0 0000 0.0000 0.0000 Oxygen 0.0000 0.9950 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 N2 + A 0.0039 0.0050 0.0055 0.0055 0.0059 0.0000 0.0005 0.0559 0.0581 0.0000 0.0000 0.0000 0.0000 H2O 0.0000 0.0000 0 2443 0.0663 0.0029 0 0000 0.0000 0.0000 0.0000 1.0000 1 0000 1.0000 1.0000 Methanol 0.0000 0.0000 0.0000 0.0033 0.0002 0.0005 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 0.0000 Ammoma 0.0000 0.0000 0.0001 0.0001 0.0000 0.0000 0.0000 0.0000 0 0000 0.0000 0.0000 0.0000 0.0000 The following can be seen from the output numbers in the Table A. Medium Pressure steam is generated as follows: 83.8 tes / hr in the steam drum of the reformed gas boiler 26 and 28.1 tes / hr in the steam drum of the isothermal water-gas shift reactor 32. Process steam addition to the hydrocarbon feed in line 18 is 34 tes / hr to give a steam to carbon ratio at the autothermal reformer inlet of 1.1. There is about 7.4 tes / hr steam used for minor uses in the process. About 68% of the process steam addition in line 18 and the minor uses can be provided by the steam raised in the isothermal water-gas shift reaction section 32. This leaves 70.5 tes / hr net steam available to be exported as superheated medium pressure steam at "-400°C and 42.5bar(a) at a rate of 7.78 tes medium pressure steam / te hydrogen product. The process carbon capture efficiency is about 97.9%. Of the inert gas (nitrogen and argon) in the feed to the pressure swing adsorption unit 54, about 43% is purged in the impure hydrogen fuel gases 58 and 158, about 5.5% ends up in the purified hydrogen gas 56. Therefore, overall 48.5% is removed and 51.5% is recycled back, directly or indirectly, to the autothermal reformer. By contrast, a similar process using adiabatic high and low temperature water-gas shift reactors and an amine wash system, produces just 4.04 tes medium pressure steam / te hydrogen product and a similar process using adiabatic high and low temperature water-gas shift reactors and cryogenic CO2 removal produces 4.14 tes medium pressure steam / te hydrogen product. The combination of an isothermal water-gas shift reactor and cryogenic CO2 removal thus enables a significant increase in medium pressure steam production whilst maintaining satisfactory carbon capture efficiency and hydrogen product purity.

Claims

1. A process for the production of hydrogen, the process comprising -subjecting a gaseous mixture comprising a hydrocarbon and steam, and having a steam to carbon ratio in a range of from 0.4 to 1.8, to autothermal reforming in an autothermal reformer, thus producing a reformed gas mixture;cooling the reformed gas mixture in a reformed gas boiler through heat exchange with water as a coolant, thereby producing a cooled reformed gas mixture and raising high or medium pressure steam;subjecting the cooled reformed gas mixture to a water-gas shift reaction under isothermal conditions in an isothermal steam raising water-gas shift reactor, at a water-gas shift reaction temperature of at least 230°C but below 320°C using a suitably stable and active water-gas shift catalyst and water as a coolant, thus producing a hydrogen-enriched reformed gas mixture and raising medium pressure steam;purifying the hydrogen-enriched reformed gas mixture in a purifier and producing at least a purified hydrogen gas and a hydrogen lean CO2 containing tail gas; andseparating CO2 from the hydrogen lean CO2 containing tail gas by subjecting the hydrogen lean CO2 containing tail gas to cryogenic fractionation by cooling it to cryogenic temperatures and recovering a liquid CO2 product and a CO2 lean gas.

2. The process according to claim 1, wherein the water-gas shift reactor isconfigured to provide for heat exchange in the reactor such that the water-gas shift reaction in the water-gas shift catalyst bed occurs in contact with heat exchange surfaces.

3. The process according to claim 1 or claim 2, wherein the water-gas shift reactoris configured such that the water-gas shift catalyst is provided in tubes surrounded by coolant, or such that coolant is provided in tubes surrounded by the catalyst.

4. The process according to any one of claims 1 to 3, wherein the water-gas shiftreactor is an axial flow reactor in which either the water-gas shift catalyst is provided inside the tubes of the reactor and coolant is provided outside the tubes or the water-gas shift catalyst is provided outside the tubes of the reactor and coolant is provided inside the tubes, or wherein the water-gas shift reactor is a radial flow reactor in which coolant is provided inside the tubes of the reactor and the water-gas shift catalyst is provided outside the tubes.

5. The process according to any one of ciaims 1 to 4, wherein the water that isused as coolant in the water-gas shift reactor is pressurised boiling water, at a pressure that is in a range of from about 25 bar(a) to about 50 bar(a).

6. The process according to any one of claim 5, wherein the water that is used ascoolant in the water-gas shift reactor is supplied to the water-gas shift reactor at a temperature that is 0°C to 20°C lower than the boiling point of water at the pressure.

7. The process according to any one of claims 1 to 6, wherein heat exchange withthe coolant in the water-gas shift reactor maintains the water-gas shift reaction temperature and wherein the water-gas shift reaction temperature therefore approximates the temperature of the coolant.

8. The process according to any one of claims 1 to 7, wherein the water-gas shiftreaction temperature is at least 250°C but less than 300°C.

9. The process according to any one of claims 1 to 8, wherein the reformed gasmixture is cooled to a temperature that approximates the temperature at which the water-gas shift reaction takes place.

10. The process according to any one of claims 1 to 9, wherein the water-gas shiftcatalyst is a medium- or low-temperature water-gas shift catalyst.

11. The process according to claim 10, wherein the water-gas shift catalyst is acopper-based catalyst.

12. The process according to any one of claims 1 to 11, wherein the gaseousmixture comprising the hydrocarbon and steam has a steam to carbon ratio in a range of from 0.6 to 1.4, or from 0.8 to 1.2.

13. The process according to any one of claims 1 to 12, which includes producingthe gaseous mixture comprising hydrocarbon and steam using high or medium pressure steam raised in the process.

14. The process according to claim 13, wherein the high or medium pressure steamcomprises medium pressure steam that is raised in the water-gas shift reactor.

15. The process according to any one of claims 13 or 14, wherein the high ormedium pressure steam demand to produce the gaseous mixture comprising hydrocarbon and steam to perform the water-gas shift reaction is met fully by the high or medium pressure steam that is generated in the reformed gas boiler and / or the medium pressure steam that is generated in the water-gas shift reactor, leaving an excess of such high or medium pressure steam.

16. The process according to claim 15, which includes exporting from the processat least a major portion, e.g. above 90 vol%, or all, of the excess high or medium pressure steam that remains after utilisation of high or medium pressure steam in producing the gaseous mixture comprising hydrocarbon and steam.

17. The process according to claim 16, which includes superheating the excesshigh or medium pressure steam that is exported from the process, such that the exported excess high or medium pressure steam is superheated high or medium pressure steam.

18. The process according to claim 17, wherein superheating the excess mediumpressure steam that would be exported from the process includes using one or more of the purified hydrogen gas or an impure hydrogen fuel gas produced in the purifier.

19. The process according to any one of claims 1 to 18, wherein the purifiercomprises a pressure swing adsorption system.

20. The process according to any one of claims 1 to 19, wherein the process furthercomprises cooling and dewatering the hydrogen enriched reformed gas mixture downstream of the water-gas shift reactor to produce a cooled, dewatered hydrogen enriched reformed gas mixture that is passed to the purifier.

21. The process according to any one of claims 1 to 20, wherein the process furthercomprises: further purifying the CO2 lean gas to produce a CO and CH4 rich gas with a lower total flowrate of inert gases than in the hydrogen enriched reformed gas; and recycling the CO and CH4 rich gas to form part of the gaseous mixture comprising the hydrocarbon and steam.

22. The process according to claim 21, wherein the further purifying is carried outin a pressure swing adsorption system or a membrane system or a combination thereof.

23. The process according to claim 21 or 22, wherein the total flowrate of inertgases in the CO and CFU rich gas is at least 10%, or 15%, or 20%, or 30% lower than the total flowrate of inert gases in the hydrogen enriched reformed gas.

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