Low-carbon hydrogen process

Abstract

A process for the production of hydrogen is described which comprises the steps of: (i) subjecting a gaseous mixture comprising a hydrocarbon and steam to steam reforming in a reforming unit comprisin

Description

This invention relates to processes for the conversion of hydrocarbons to hydrogen whilst minimising carbon dioxide production and emission. Processes for generating hydrogen are well-known and generally include a fired steam methane reformer combined with water-gas shift and carbon dioxide (CO2) removal. Such processes create significant volumes of carbon dioxide in flue gases at pressures unsuitable for efficient CO2 capture. A common goal is to increase the rate of progress towards the net zero objective. There is, in the interim, a need for hydrogen production processes that generate lower levels of carbon dioxide effluent and enable more efficient CO2 capture. In support of this, blue hydrogen or low-carbon hydrogen processes are in development. WO2023148469 A1 discloses a process for the production of hydrogen is described comprising the steps of: (i) subjecting a gaseous mixture comprising a hydrocarbon and steam to steam reforming in a gas-heated reformer or adiabatic pre-reformer followed by autothermal reforming with an oxygen-rich gas in an autothermal reformer to generate a reformed gas mixture, (ii) increasing the hydrogen content of the reformed gas mixture by subjecting it to one or more water-gas shift stages in a water-gas shift unit to provide a hydrogen-enriched reformed gas, (iii) passing the hydrogen- enriched reformed gas and an oxygen-rich gas to an oxidation unit containing an oxidation catalyst that converts carbon monoxide present in the hydrogen-enriched reformed gas to carbon dioxide, to form a carbon-dioxide-enriched gas mixture, (iv) cooling the carbon dioxide-enriched gas mixture and separating condensed water therefrom, and (v) passing the carbon dioxide-enriched gas mixture to a carbon dioxide separation unit to provide a carbon dioxide gas stream and a hydrogen product gas stream.. We have developed an improved process where carbon dioxide emissions are reduced. Accordingly, the invention provides a process for the production of hydrogen comprising the steps of: (i) subjecting a gaseous mixture comprising a hydrocarbon and steam to steam reforming in a reforming unit comprising an autothermal reformer to generate a reformed gas mixture; (ii) increasing the hydrogen content of the reformed gas mixture by subjecting it to one or more water-gas shift stages in a water-gas shift unit to provide a hydrogen-enriched reformed gas; (iii) passing the hydrogen-enriched reformed gas and an oxygen-rich gas to an oxidation unit containing an oxidation catalyst that converts carbon monoxide present in the hydrogen-enriched reformed gas to carbon dioxide, to form a carbon dioxide-enriched gas mixture; and (iv) passing the carbon dioxide-enriched gas mixture to a product separation unit to provide a hydrogen gas stream, a carbon dioxide gas stream and a by-product gas stream; wherein at least a portion of the by-product gas stream is compressed and recycled to one or more of the reforming unit, the water-gas-shift unit, the oxidation unit, and the product separation unit. The Applicants have found that the combination of by-product gas recycle and carbon monoxide oxidation allows the added benefit of high feedstock efficiency and high carbon capture, achieving >98% capture rate without the need for other measures such as an electric steam boiler and / or increased steam to carbon ratio. The addition of carbon monoxide oxidation also allows the by-product gas recycle stream to be richer in hydrogen with a very low carbon monoxide content, thereby reducing unwanted methanation in the reforming unit, and lower CO2 emissions in the event a portion of the by-product gas is combusted as a fuel gas. Furthermore, the combination also provides improved flexibility in case of underperformance in the reforming unit or the water-gas shift unit, or when the plant is operating at a reduced rate resulting from an interruption to the flow of hydrocarbon feed or reduced product hydrogen demand. The present invention uses a gaseous mixture comprising a hydrocarbon. The gaseous mixture may comprise any gaseous or low boiling hydrocarbon, such as natural gas, associated gas, LPG, petroleum distillate, diesel, naphtha or mixtures thereof, or hydrocarbon-containing off-gases from chemical processes, such as a refinery off-gas. The gaseous mixture preferably comprises methane, associated gas or natural gas containing a substantial proportion, e.g. over 50% v / v methane. Natural gas is especially preferred. The hydrocarbon may be compressed to a pressure in the range of 10-100 bar abs. The pressure of the hydrocarbon may usefully govern the pressure throughout the process. Operating pressure is preferably in the range of 15-50 bar abs, more preferably 25-50 bar abs as this provides an enhanced performance from the process. If the hydrocarbon contains sulphur compounds, before, or preferably after, compression it may be subjected to desulphurisation. In some arrangement this includes hydrodesulphurisation using C0M0 or NiMo catalysts, and absorption of hydrogen sulphide using a suitable hydrogen sulphide adsorbent, e.g. a zinc oxide adsorbent. An ultra-purification adsorbent may usefully be employed downstream of the hydrogen sulphide adsorbent to further protect the steam reforming or water gas shift catalyst. Suitable, ultra-purification adsorbents may comprise copper-zinc oxide / alumina materials and copper-nickel-zinc oxide / alumina materials. To facilitate hydrodesulphurisation and / or reduce the risk of carbon laydown in the reforming process, a hydrogen-containing gas may be added to the compressed hydrocarbon. The hydrogen-containing gas may be combined with the hydrocarbon upstream and / or downstream of any hydrodesulphurisation stage. The amount of hydrogen in the resulting mixed gas stream, e.g. downstream of the hydrodesulphurisation stage, may be in the range of 1-20% vol, but is preferably in the range 1-10% vol, more preferably in the range of 1-5% vol on a dry gas basis, especially if there are no unsaturated compounds and low level of organic sulphur in the compressed hydrocarbons. In a preferred embodiment, during normal operation, because the by-product gas contains some hydrogen, a portion of the by-product gas from the product separation unit may be compressed and mixed with the compressed hydrocarbon. In other arrangements, desulphurisation may be achieved by treatment of at least a portion of the hydrocarbon by a bulk sulphur removal step including a washing stage with a liquid absorbent in an acid gas recovery unit. This bulk sulphur removal may, if desired, be performed on a portion of the hydrocarbon feed, which is then mixed with the remaining portion after the bulk sulphur removal and the mixed feed is then subjected to hydrodesulphurisation, hydrogen sulphide removal and optionally ultra-purification as described above. If the hydrocarbon contains unsaturated compounds such as olefins, the hydrodesulphurisation and / or H2S absorption vessels may include a recycle arrangement where a portion of hot outlet stream from these vessels is cooled, compressed, and mixed with fresh hydrocarbon to reduce the total level of unsaturates entering the hydrodesulphurisation vessel thereby controlling the heat generated by the saturation reactions. If the hydrocarbon contains other contaminants, such as chloride or heavy metal contaminants, these may be removed, prior to reforming, upstream ordownstream of any desulphurisation, using conventional adsorbents. Adsorbents suitable for chloride removal are known and include alkalised alumina materials. Similarly, adsorbents for heavy metals such as mercury or arsenic are known and include copper sulphide materials. The hydrocarbon may be pre-heated. It may conveniently be pre-heated after compression and before desulphurisation. Various hot gas sources are provided in the present process that may be used for this duty. However, in a preferred embodiment, the hydrocarbon is heated by passing it through a heat exchanger heated by the carbon dioxide-enriched gas recovered from the oxidation unit. The hydrocarbon is mixed with steam. The steam introduction may be performed by direct injection of steam, including injection of a stripped process condensate steam, and / or by saturation of the hydrocarbon by contact with a stream of heated water. In some arrangements, a gaseous mixture comprising the hydrocarbon and steam may be formed by directly mixing the hydrocarbon with steam, for example steam generated in one or more fired heaters and / or steam generated by cooling the reformed gas mixture with water and / or steam recovered from a stripping unit that has been used to strip process condensate. In other arrangements, the hydrocarbon may be saturated in a saturator fed with hot water to form a saturated gas mixture. The water fed to the saturator may comprise one or more of the condensate streams produced in the process. The steam content of the saturated gas mixture may, if desired, be increased by the direct addition of steam. In some arrangements, a saturated hydrocarbon gas may be combined with steam generated in the water gas shift unit. In the present invention, at least a portion of the by-product gas is compressed and recycled to one or more of the reforming unit, the water-gas-shift unit, the oxidation unit, and the product separation unit. Hence, a portion of the by-product gas may be combined with the hydrocarbon feedstock, before or after any desulphurisation, or the gaseous mixture of hydrocarbon and steam upstream, of the reforming unit. Alternatively, a portion of the by-product gas may be combined with a pre-reformed or partially reformed gas mixture generated within the reforming unit. Additionally or alternatively, if desired, a portion of the compressed by-product gas may also be recycled to the process by combining it with the reformed gas upstream of the water gas shift unit and / or the hydrogen-enriched gas upstream of the oxidation unit, and / or the carbon dioxide-enriched gas upstream of the product separation unit. If desired, the portion recycled to the product separation unit may be fed to one or more stages of product separation within the unit. The reforming unit comprises an autothermal reformer. Autothermal reformers are known and typically may comprise an elongate vessel arranged vertically, having a burner disposed at the top of the reformer, to which a feed gas containing hydrocarbon 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 therefore provided by combustion of a portion of hydrocarbon in the feed gas. The feed gas 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 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 50% vol O2 and may be an oxygen-enriched air mixture. However, in the present invention the oxygen-rich gas preferably comprises at least 90% vol O2, more preferably at least 95% vol O2, 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, an air separation unit (ASU) or by the electrolysis of water. The ASU may be electrically driven and is desirably driven using renewable electricity to further improve the efficiency of the process and minimise CO2 emissions. The amount of oxygen-rich gas added to the autothermal reformer is preferably such that 40 to 60 moles of oxygen are added per 100 moles of hydrocarbon carbon in the hydrocarbon fed to the process. Preferably the amount of oxygen added is such that the reformed gas leaves the catalyst in the autothermal reformer at a temperature in the range 800-1100°C, more preferably 900-1100°C, most preferably 970-1070°C. In a preferred embodiment, a small purge of steam may be added to the oxygen-rich gas to protect against reverse flow at low plant rates or if the plant trips. The flow of steam purge can vary depending on the plant rate. Whereas the gaseous mixture of hydrocarbon and steam may be fed directly to the autothermal reformer, preferably the feed gas to the autothermal reformer is a partially reformed gas mixture, in particular a partially reformed gas mixture generated upstream of the autothermal reformer by adiabatic pre-reforming in a pre-reformer or steam reforming in a gas-heated reformer that comprises catalyst-containing tubes heated by the reformed gas mixture recovered from the autothermal reformer. The pre-reformer or gas-heated reformer and the autothermal reformer are therefore operated in series. The partially reformed gas mixture, comprising methane, hydrogen, steam and carbon oxides, may be fed without any dilution or heat exchange, directly to the autothermal reformer in which it is subjected to autothermal reforming to generate the reformed gas mixture. In arrangements using a gas-heated reformer and autothermal reformer in series, the amount of steam introduced is desirably sufficient to give a steam to carbon ratio at the inlet to the reforming unit operations of at least 1.5:1, i.e. at least 1.5 moles of steam per mole of hydrocarbon carbon in the gaseous mixture. In a preferred operation, the steam to carbon ratio is at least 2:1 and more preferably at least 2.5:1. The steam to carbon ratio can be adjusted in a preferred range of 1.5:1 to 3.5:1 to provide an optimal balance of feedstock efficiency and carbon capture rate. In arrangements using an adiabatic pre-reformer and autothermal reformer in series, the amount of steam introduced is desirably sufficient to give a steam to carbon ratio (defined as the steam to hydrocarbon carbon ratio at the inlet to reforming unit operations) of at least 0.4:1, i.e. at least 0.4 moles of steam per mole of hydrocarbon carbon in the gaseous mixture, with a preferred range of 0.4:1 to 5:1. Where the steam to carbon ratio at the inlet to the reforming unit operations is in the range 0.4:1 to less than 2.4:1, it is possible to add additional steam to the reformed gas upstream of the water-gas shift stage. Operating the reforming section at a steam to carbon ratio in the range of 0.4:1 to less than 2.4:1 has the advantage that the heating requirement and oxygen demand for the reforming stages is reduced and that the front-end equipment (e.g. fired heater, pre-reformer, and autothermal reformer) will be smaller and lower in cost. Where the steam to carbon ratio is in the range of 2.4:1 to 5:1, no further steam addition upstream of the water-gas shift unit is necessary, which may be useful in circumstances where steam addition to the reformed gas is impractical. The gaseous mixture comprising hydrocarbon and steam is desirably pre-heated prior to reforming. The preheating may be performed using heat generated downstream in the process. In some arrangements, the invention comprises a stage of adiabatic pre-reforming upstream of the autothermal reformer. One or more pre-reformers may be used. In these arrangements, the gaseous mixture comprising the hydrocarbon and steam may be preheated by passing through one or more heat exchangers downstream of a reformed gas boiler or water gas shift unit. The preheated gas mixture may then be subjected to a step of adiabatic steam reforming in one or more pre-reformer vessels containing a fixed bed of a pre-reforming catalyst. In such a process, the gaseous mixture comprising the hydrocarbon and steam, typically at an inlet temperature in the range of 380-650°C, is passed 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, any hydrocarbons higher than methane 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, can be desirable to ensure that the feed to the autothermal reformer contains no hydrocarbons higher than methane and also contains some hydrogen. In other arrangements, the gaseous mixture comprising the hydrocarbon and steam is subjected to steam reforming in a gas-heated reformer. In a preferred embodiment using a gas-heated reformer and autothermal reformer, the gaseous mixture may be preheated by passing it through a reformed gas interchanger, where it is heated by the reformed gas mixture. Desirably, the mixed stream is heated to 400-500°C, preferably 420-460°C. Different types of gas-heated reformer may be used. In one type of gas-heated reformer, the catalyst is disposed in tubes extending between a pair of tube sheets through a heat exchange zone. Reactants are fed to a zone above the upper tube sheet and pass through the tubes and into a zone beneath the lower tube sheet. The heating medium is passed through the zone between the two tube sheets. The heating medium is typically a hot reformed gas recovered from the autothermal reformer. Gas-heated reformers of this type are described in GB1578270 and WO97 / 05 947. The gaseous mixture comprising the hydrocarbon and steam is passed through the catalyst-filled tubes in the gas-heated reformer. During passage through the reforming catalyst, the endothermic steam reforming reaction takes place with the heat required for the reaction supplied by a hot reformed gas recovered from the autothermal reformer, which flows past the exterior surface of the tubes. The steam reforming catalyst used in the gas-heated reformer may comprise nickel supported on a particulate refractory support such as rings or multi-holed pellets of calcium aluminate, magnesium aluminate, alumina, titania, zirconia and the like. Alternatively, a combination of nickel and a precious metal such as ruthenium, may be used. In place of, or in addition to, the particulate steam reforming catalyst, the steam reforming catalyst may comprise one or more structured catalyst units, which may be in the form of metal or ceramic monoliths or folded metal structures on which a layer of nickel and / or precious metal steam reforming catalyst has been deposited. The temperature of the autothermally reformed gas used to heat the gas-heated reformer is preferably sufficient that the gas undergoing steam reforming leaves the catalyst tubes at a temperature in the range 600-850°C, preferably 650-750°C, more preferably 680-720°C. In arrangements comprising a pre-reformer and autothermal reformer, the autothermally reformed gas may be cooled in heat exchange with water, e.g., in a waste-heat boiler, to generate steam. This steam may be used as process steam added to the feed to the reforming unit or water gas shift unit. The steam may also be used for heating and / or for power generation in a steam turbine. In arrangements comprising a gas-heated reformer and an autothermal reformer, the reformed gas produced by the autothermal reformer is used to provide the heat required for the steam reforming step by using it as the hot gas flowing past the tubes in the gas-heated reformer. During this heat exchange, the reformed gas cools by transferring heat to the gas undergoing steam reforming. Preferably, the reformed gas cools by several hundred degrees Centigrade but it will leave the gas-heated reformer at a temperature somewhat above the temperature at which the gaseous mixture comprising hydrocarbon and steam mixture is fed to the gas-heated reformer. Preferably the reformed gas leaves the gas-heated reformer at a temperature in the range 450-650°C, more preferably 450-580°C. After leaving the gas-heated reformer, the reformed gas is desirably then further cooled in one or more steps of heat exchange. Heat recovered during this cooling may be employed for reactants pre-heating and / or for heating water used to provide the steam employed in the steam reforming step. In some arrangements, the reformed gas mixture exiting the shell side of the gas-heated reformer may be used to heat water fed to a saturator. Heat recovered from the reformed gas may additionally, or alternatively, be used in the carbon dioxide separation step. The reformed gas comprises hydrogen, carbon monoxide, carbon dioxide, steam, and a small amount of unreacted methane, and may also contain small amounts of inert gases such as nitrogen and argon. Preferably, the hydrogen content of the reformed gas is in the range of 30-50% vol on a wet gas basis and the carbon monoxide content in the range 5-15% vol on a wet gas basis. In the present invention, the hydrogen content of the reformed gas mixture is increased by subjecting it to one or more water-gas shift stages in a water-gas shift unit thereby producing a hydrogen-enriched reformed gas and at the same time converting carbon monoxide in the reformed gas to carbon dioxide. The reaction may be depicted as follows; CO + H2O^CO2 + H2 Whereas steam is present in the reformed gas, supplemental steam may be added before the one or more water gas shift stages, e.g. by direct addition to the reformed gas, if desired. The one or more water-gas shift stages may include stages of high-temperature shift, mediumtemperature shift, isothermal shift and low-temperature shift. High-temperature shift is operated adiabatically in a shift vessel with inlet temperature in the range of 300-400°C, preferably 320-360°C, over a bed of a reduced iron catalyst, such as chromia-promoted magnetite. Alternatively, a promoted zinc-aluminate catalyst may be used. Medium-temperature shift and low-temperature shift stages may be performed using shift vessels containing supported copper catalysts, particularly copper / zinc oxide / alumina compositions. In low-temperature shift, a gas containing carbon monoxide (preferably <6% vol CO on a dry basis) and steam (at a steam to total dry gas molar ratio in the range of 0.3 to 1.5) may be passed over the catalyst in an adiabatic fixed bed with an outlet temperature in the range 200 to 300°C. The outlet carbon monoxide content may be in the range of 0.1 to 1.5%, preferably under 0.5% vol on a dry basis if additional steam is added. Alternatively, in mediumtemperature shift, the gas containing carbon monoxide and steam may be fed to the catalyst at an inlet temperature in the range of 200 to 240°C although the inlet temperature may be as high as 280°C. The outlet temperature may be up to 300°C but may be as high as 360°C. Whereas one or more adiabatic water-gas shift stages may be employed, such as a high-temperature shift stage, optionally followed by a low-temperature shift stage, the reformed gas may be subjected to a stage of isothermal water-gas shift in a cooled shift vessel, optionally followed by one or more adiabatic medium- or low-temperature water-gas shift stages in uncooled vessels as described above. Using an isothermal shift stage, i.e. with heat exchange in the shift converter such that the exothermic reaction in the catalyst bed occurs in contact with heat exchange surfaces that remove heat, offers the potential to use the reformed gas stream in a very efficient manner. Whereas the term “isothermal” is used to describe a cooled shift converter, there may be a small increase in the temperature of the gas between the inlet and outlet, so that the temperature of the hydrogen-enriched reformed gas stream at the exit of the isothermal shift converter may be between 1 and 25 degrees Celsius higher than the inlet temperature. The coolant conveniently may be water under pressure such that partial, or complete, boiling takes place. The water can be in tubes surrounded by catalyst or vice versa. The resulting steam can be used, for example, for heating, or to drive a turbine, e.g. for electrical power, or to provide process steam for supply to the process. The hydrogen-enriched reformed gas contains steam. In some arrangements it may be desirable, following the one or more water-gas shift stages, to cool the hydrogen-enriched reformed gas to a temperature below the dew point so that at least a portion of the steam condenses. The liquid water condensate may then be separated using one or more gas-liquid separators, which may have one or more further cooling stages between them, to form a dewatered hydrogen-enriched reformed gas. Any coolant may be used. In some arrangements, the cooling includes heat exchange with one or more of the process streams such as boiler feed water or hydrocarbon feed. Cooling may also be performed in heat exchange with one or more liquids in the product separation unit, for example in absorbent regeneration. In a preferred arrangement, the hydrogen-enriched reformed gas is fed to a boiler to generate steam. The steam may be used for heating, used a as a process stream, or used for electricity generation in a steam turbine. In some arrangements, cooling of the hydrogen-enriched reformed gas is carried out in heat exchange with the process condensate. The process condensate may be the condensate recovered from the cooled hydrogen-enriched reformed gas and any other aqueous process condensate recovered from the process, such as process condensate recovered downstream of the oxidation unit and process condensate recovered from knock-out pots coupled to compression equipment. As a result, a stream of heated condensate is produced, which may be fed to a saturator or process condensate stripper and used to supply some or all of the steam required for reforming and / or water-gas shift stages. Thus, in some arrangements, condensates recovered from the hydrogen-enriched reformed gas and the carbon dioxideenriched gas mixture are used to provide at least a portion of steam fed to the reforming unit and / or to supplement the steam present in the reformed gas fed to the water gas shift unit. Because the condensate may contain ammonia, methanol, hydrogen cyanide and CO2, returning the condensate to form steam offers a useful way of returning hydrogen and carbon to the process. One, two or more stages of cooling and condensate recovery may be included upstream and / or downstream of the oxidation unit. However, in some arrangements, it may not be necessary for condensate recovery from the hydrogen-enriched gas mixture upstream of the oxidation unit. In the present invention, the hydrogen-enriched reformed gas, optionally after cooling and separation of condensate, is fed to an oxidation unit. In the oxidation unit, carbon monoxide present in the hydrogen-enriched reformed gas is oxidised to carbon dioxide. An oxygen-rich gas is added to the hydrogen-enriched reformed gas. The oxygen-rich gas may be generated by vacuum pressure swing adsorption (VPSA) unit, an air separation unit and / or by electrolysis of water. Some hydrogen is also oxidised to form water. Accordingly, the oxidation catalyst is preferably a CO-selective oxidation catalyst to minimise hydrogen losses. The oxidation reactions may be depicted as follows; 2 CO + O2 2 CO2 2 H2 + O2 2 H2O Other reactions e.g. oxidation of residual methane to CO2 and H2O may also occur. The oxidation catalyst is preferably a supported precious metal catalyst. For example, the catalyst may comprise one or more of Pt, Pd, Rh, Ir or Ru, desirably on an oxidic support such as alumina, titania, zirconia or silica. The amount of precious metal may be in the range of 0.1 to 5% by weight. The oxidation catalyst may be in the form of pellets or extrudates, a foam, monolith or coating on an inert support. Precious metal oxidation catalysts suitable for CO-oxidation preferably consist of alumina-supported platinum promoted with an oxide of a metal selected from the group consisting of manganese, iron, cobalt, copper, nickel and mixtures thereof. Particularly suitable catalysts are alumina-supported platinum catalysts promoted with iron oxide and / or copper oxide. The loading of platinum, if present, on a particulate support material should be in the range of from about 1 to 5 weight percent, preferably about 1 to 3 weight percent. The copper loading, if present, should be about 2-12 weight percent, and preferably 4-8 weight percent. The iron loading, if present, is desirably from about 0.1-2 weight percent, and preferably from about 0.2-1 weight percent. Such catalysts are described in US2006276332A1, US6559094 and US3088919. Alternatively, the oxidation catalyst may consist of a supported copper oxide. For example, CN102407123A discloses CuO supported on ceria as a preferential CO-oxidation catalyst. The oxidation reaction is exothermic, and the reaction may be performed adiabatically in a fixed bed oxidation vessel. The flow through the bed may be axial and / or radial flow. The inlet temperature in such an arrangement may be in the range of 20 to 200°C, and the exotherm in the bed is desirably kept below 75 degrees Celsius. Lower temperatures in the catalyst bed generally favour greater selectivity to CO-oxidation and are preferred. Therefore, it is preferable to locate an adiabatic oxidation vessel downstream of one or more stages of cooling and condensate recovery. In this way a higher selectivity can be achieved, which limits the temperature rise across the adiabatic catalyst bed and affords a longer catalyst lifetime. The oxidation unit may comprise two or more reactors in parallel. In some arrangements, it is preferred to operate the oxidation unit with staged oxygen addition. In such arrangements, the oxidation unit may comprise two or more adiabatic oxidation reactors in series. A first oxygencontaining gas stream may be fed to the first oxidation reactor to carry out a first stage of selective CO oxidation. The effluent from the first oxidation reactor may be cooled in indirect heat exchange, for example with cooling water or with a process stream, to form a first cooled effluent stream. A second oxygen-containing gas stream may be added to the first cooled effluent stream and the resulting mixture fed to a second oxidation reactor to carry out a second stage of selective CO oxidation. The sequence may be repeated with additional reaction, cooling and oxygen addition stages. The oxidation step may alternatively be operated with cooling of the catalyst bed by passing a gaseous or liquid coolant, such as a suitable process stream, preferably steam or a boiling water-steam mixture, through one or more tubes disposed within the catalyst bed. Alternatively, the catalyst can be in tubes surrounded by boiling water. Where a cooled oxidation reactor is used, it is possible to locate the oxidation unit directly downstream of the water-gas shift unit, especially where the water gas shift unit comprises an isothermal shift vessel, with no or minimal cooling stage in between. In this arrangement, the gas mixture fed to the oxidation unit is the hydrogen-enriched reformed gas containing steam and the inlet temperature for the oxidation unit will be close to the exit temperature from the water-gas shift unit and may be in the range of 200 to 320°C. Where the coolant is water under pressure such that partial, or complete, boiling takes place, the water pressure is preferably the same as in the isothermal shift converter, so that a single steam drum can be shared by the isothermal water-gas shift and oxidation units. Alternatively, the water pressure in the cooled oxidation unit may be different from the water pressure in the isothermal water-gas shift unit, preferably lower. A cooled oxidation unit may also be used in arrangements where an isothermal water-gas shift converter is not used. In such arrangements, the water gas shift unit may comprise one or more adiabatic reactors. The water pressure in the cooled oxidation unit may range from atmospheric to 50 bar. Additionally, to limit the peak temperature in a cooled oxidation unit, the oxidation unit may comprise two or more cooled catalyst beds in series, and part of the oxygen-rich gas may be added in stages between successive catalyst beds. The two or more cooled catalyst beds may be contained within the same pressure vessel or within different pressure vessels. In the present invention the oxidation unit is fed with an oxygen-rich gas, preferably a portion of the same oxygen-rich gas fed to the autothermal reformer. Thus, the oxygen-rich gas fed to the oxidation unit preferably comprises at least 90% vol O2, more preferably at least 95% vol O2, most preferably at least 98% vol O2, or at least 99% vol O2, e.g. a pure oxygen gas stream. In some arrangements the oxygen may be provided by electrolysis, using renewable sources of electricity, or electricity produced by a turbine powered by the process, e.g., from combustion of a portion of the by-product gas or steam provided by combustion of a portion of the by-product gas. In order to ensure high conversion of the residual carbon monoxide, the oxygen may be added in stoichiometric excess of the carbon monoxide, but too high an excess may cause unwanted side reactions. Consequently, the oxygen in the gas mixture is preferably less than 150% in excess of the stoichiometric amount. The oxidation unit produces a carbon dioxide-enriched gas mixture. Following the oxidation step, the carbon dioxide-enriched gas mixture is preferably cooled to below the dew point to cause condensation of steam present in the oxidised gas mixture. The resulting liquid water condensate may then be separated using one or more gas-liquid separators, which may have one or more further cooling stages between them, to form a de-watered carbon dioxide-enriched reformed gas. The cooling may be performed in heat exchange in one or more stages using water, air, process streams, ora combination of these. As set out above, some or all of the condensate recovered downstream of the oxidation unit may be used to generate steam for the steam reforming and / or water-gas shift stages. The carbon dioxide-enriched gas mixture, optionally after cooling and condensate recovery, is fed to a product separation unit to separate hydrogen from the other components and generate a hydrogen gas product stream, The separation unit further provides a carbon dioxide stream and a by-product stream. The separation unit may comprise two distinct arrangements; in a first separation arrangement it may comprise a hydrogen separation unit coupled to a downstream cryogenic carbon dioxide separation unit (as well as ancillary units such as one or more compressors, dehydration units, and by-product separation units), or in a second separation arrangement it may comprise a carbon dioxide removal unit coupled to a downstream hydrogen purification unit. Each type of separation unit arrangement provides distinct advantages. The first separation arrangement produces a hydrogen gas stream, a carbon dioxide gas stream, a by-product gas stream and a separate fuel gas stream. The second separation arrangement produces a hydrogen gas stream, a carbon dioxide gas stream and a by-product gas stream, a portion of which may be used as a fuel gas if desired. In both cases, a portion of the by-product gas stream is recycled to the process. In the first separation arrangement, the hydrogen separation unit provides a purified hydrogen product gas and a hydrogen depleted gas. The hydrogen separation unit may suitably comprise a membrane system, a temperature swing adsorption system, or a pressure swing adsorption (PSA) system. Such systems are commercially available. The hydrogen separation unit can produce a pure hydrogen stream preferably with a purity greater than 99.5% vol, more preferably greater than 99.9% vol. The hydrogen separation unit also provides a hydrogen-depleted gas stream. The hydrogen-depleted gas stream comprises inert gases such as nitrogen and argon, methane, carbon monoxide, hydrogen and carbon dioxide. Because the process includes a carbon monoxide oxidation step, the carbon monoxide content of the hydrogen-depleted gas will be lower than arrangements without the oxidation unit, which permits higher levels of by-product recycle gas to be generated from the hydrogen-depleted gas than previously considered possible because the unwanted methanation reaction is suppressed. The hydrogen-depleted gas in the present invention preferably has a carbon monoxide content of less than 100 ppmv or 50 ppmv, preferably less than 20 ppmv, more preferably less than 10 ppmv. In some arrangements, the hydrogen separation unit can also provide a hydrogen-rich fuel gas that can be combusted to generate steam or pre-heat process streams. The hydrogen-depleted gas stream recovered from the hydrogen separation unit is preferably compressed, e.g. to between 50 and 80 barabs, and dehydrated,, e.g. to below lOOppmv water using a suitable absorbent, before being fed into the cryogenic carbon dioxide separation unit to prevent freezing due to lower temperature expected in the cryogenic carbon dioxide separation unit. The cryogenic carbon dioxide separation unit may comprise of a series of cooling or refrigeration devices configured to liquefy the hydrogen-depleted gas streams, a fractionation unit to separate carbon dioxide, and desirably ancillary equipment such as a carbon dioxide pump to increase the pressure of separated carbon dioxide, and a closed or open loop refrigeration unit. The separated carbon dioxide stream provided by this arrangement may be recovered from the bottom of the fractionation column and typically comprises carbon dioxide, with low levels of other compounds such as methane, inerts and water. Preferably the amount of carbon dioxide present in the carbon dioxide stream is in excess of 95% mol on a wet basis, preferably in excess of 99% mol on a wet basis, more preferably in excess of 99.5% mol on a wet basis. The carbon dioxide-depleted stream may be recovered from the top of the fractionation column. The carbon dioxide-depleted stream may be processed directly, e.g. to form a by-product gas stream that may be recycled as a recycle gas to the process, or may be sent to a secondary hydrogen separation unit downstream of the cryogenic carbon dioxide separation unit to increase the recovery of hydrogen into product hydrogen. The secondary hydrogen purification unit may suitably comprise a membrane system, a temperature swing adsorption system, or a pressure swing adsorption (PSA) system. Such systems are commercially available. The secondary hydrogen purification unit produces a pure hydrogen stream preferably with a purity greater than 99.5% vol, more preferably greater than 99.9% vol, and is preferably mixed with pure hydrogen stream recovered from the hydrogen separation unit upstream of the cryogenic carbon dioxide separation unit. The secondary hydrogen separation unit may also generate a carbon dioxide recycle stream. In some arrangements, the carbon dioxide recycle stream may be combined with the hydrogen-depleted gas recovered from the hydrogen separation unit and the resulting mixture compressed and fed to the cryogenic carbon dioxide separation unit to increase carbon dioxide recovery. The secondary hydrogen separation unit may also generate a by-product gas stream that may be recycled as a recycle gas to the process or further purified with subsequent unit operations, such as membrane unit, to generate a carbon-containing recycle gas and a further pure hydrogen stream that may be combined with the pure hydrogen product stream recovered from the hydrogen separation unit upstream of the cryogenic carbon dioxide separation unit. The by-product gas stream generated by this arrangement that may be recycled or further processed typically comprises hydrogen, methane, carbon monoxide, carbon dioxide, inerts such as nitrogen and argon, and water. Preferably, the hydrogen content of the by-product gas is in the range of 40-80% vol on a wet gas basis and the methane content in the range 5-25% vol on a wet gas basis. In addition, Preferably, the carbon monoxide content of the by-product gas is lower than 10%vol on a wet gas basis, more preferably lower than 1%vol on a wet gas basis. The secondary hydrogen separation unit may also generate a fuel gas stream that may be combusted to generate steam or pre-heat process streams, The fuel gas typically comprises hydrogen and inert gases with very low levels of other compounds such as methane, carbon monoxide, carbon dioxide and water. Preferably, the hydrogen content of the fuel gas is in the range of 80-99% vol on a wet gas basis and the inert content in the range 2-5% vol on a wet gas basis. In the second separation arrangement, the carbon dioxide removal unit may comprise a physical wash system or a reactive wash system, preferably a reactive wash system, especially an amine wash system. The carbon dioxide may be separated by an acid gas recovery (AGR) process. In the AGR process the carbon dioxide-enriched gas stream is contacted with a stream of a suitable absorbent liquid, such as an amine, particularly methyl diethanolamine (MDEA) solution so that the carbon dioxide is absorbed by the liquid to give a laden absorbent liquid and a gas stream having a decreased content of carbon dioxide. The laden absorbent liquid is then regenerated by heating and / or reducing the pressure, to desorb the carbon dioxide and to give a regenerated absorbent liquid, which is then recycled to the carbon dioxide absorption stage. Alternatively, methanol or a glycol may be used to capture the carbon dioxide in a similar manner as the amine. If the carbon dioxide separation step is operated as a single pressure process, i.e. essentially the same pressure is employed in the absorption and regeneration steps, only a little recompression of the recycled carbon dioxide will be required. The recovered carbon dioxide may comprise 90-95% vol carbon dioxide with the balance comprising water, hydrogen, inerts gases and traces of other volatile compounds such as methane, carbon monoxide, methanol and ammonia. The recovered carbon dioxide may be further purified if desired and used for the manufacture of chemicals such as methanol, sent to storage or sequestration, or used in enhanced oil recovery (EOR) processes. In cases where the CO2 is to be compressed for storage, transportation, use in EOR processes or conversion to other chemical products, the CO2 may be dried to prevent liquid water present in trace amounts, from condensing. For example, the CO2 may be dried to a dew point <-10°C by passing it through a bed of a suitable desiccant, such as a zeolite, or contacting it with a glycol in a glycol drying unit. The CO2 may also be subjected to a purification step, for example by washing, to remove contaminants such as methanol and ammonia, depending on the required specification. The washing stage may be carried out before product separation, compression or drying. The CO2 may if desired also be subjected to a liquefaction step using known liquefaction schemes after product separation. Upon the separation of the carbon dioxide, a crude hydrogen product gas stream is produced The crude hydrogen stream may comprise 95-99% vol hydrogen with the balance comprising methane, carbon monoxide, carbon dioxide and inert gases. The methane content may be in the range of 0.25-1.5% vol, preferably 0.25-0.6% vol. The carbon monoxide content may be less than 100 ppmv or 50 ppmv, preferably less than 20 ppmv, more preferably less than 10 ppmv. The carbon dioxide content may be in the range of 0.01-0.5% vol, preferably 0.01-0.1% vol. The balance may be made up of nitrogen and residual water vapour. In the second separation arrangement, the crude hydrogen product gas stream is passed to a hydrogen purification unit to provide a purified hydrogen product gas and a by-product gas. The hydrogen purification unit may suitably comprise a membrane system, a temperature swing adsorption system, ora pressure swing adsorption (PSA) system. Such systems are commercially available. The hydrogen purification unit can produce a pure hydrogen stream preferably with a purity greater than 99.5% vol, more preferably greater than 99.9% vol. The hydrogen purification unit also provides a by-product gas stream. The by-product gas stream comprises inert gases such as nitrogen and argon, methane, carbon monoxide, hydrogen and carbon dioxide. Preferably the by-product gas may comprise 80-90% vol hydrogen with the balance comprising inert gases, water, carbon oxides, and methane. Because the process includes a carbon monoxide oxidation step, the carbon monoxide content of the by-product gas will be lower than arrangements without the oxidation unit, which permits higher levels of by-product gas recycle to the process than previously considered possible because the unwanted methanation reaction is suppressed. The by-product gas in the present invention preferably has a carbon monoxide content of less than 100 ppmv or 50 ppmv, preferably less than 20 ppmv, more preferably less than 10 ppmv. Whereas all of the by-product gas may be recycled to the process, in some arrangements it is advantageous to divide the by-product gas and consume part as a fuel gas to prevent the buildup of inert gases in the process. Thus, in some arrangements, the by-product gas is divided into a first portion and a second portion, the first portion is recycled to the process and the second portion is fed to one or more fired heaters used to heat feed steams or provide steam for the process. The first portion may be 50 to 90% by volume of the by-product gas. A portion of the hydrogen product, with or without purification, may be compressed if necessary and recycled to the hydrocarbon feed if desired for desulphurisation and to reduce the potential for carbon formation in the reforming unit. The hydrogen product gas may be compressed and used in downstream power or heating processes, for example, by using it as fuel in a gas turbine (GT) or by injection into a domestic or industrial networked gas piping system. The hydrogen product may also be used in a downstream chemical synthesis process. Thus, the hydrogen product may be used to produce ammonia by reaction with nitrogen in an ammonia synthesis unit. Alternatively, the hydrogen product may be used with a carbon dioxide-containing gas to manufacture methanol, in a methanol production unit or substitute natural gas in a substitute natural gas production unit. Alternatively, the hydrogen product may be used with a carbon-monoxide containing gas to synthesise hydrocarbons in a Fischer-Tropsch production unit. Alternatively, the hydrogen may 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. The invention is illustrated by reference to the accompanying drawing in which: Figure 1 depicts a flowsheet of one embodiment of the invention, and Figure 2 depicts of a flowsheet of another embodiment 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 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 ancillary items of equipment forms no part of the present invention and is in accordance with conventional chemical engineering practice. In Figure 1, a natural gas stream 10 is combined with a first portion of a by-product gas stream supplied via line 12, and the resulting mixture heated and passed to a desulphurisation unit 14 comprising a hydrodesulphurisation vessel containing a hydrodesulphurisation catalyst followed by two vessels containing zinc oxide hydrogen sulphide absorbents and ultra-purification adsorbents. Sulphur compounds present in the natural gas are converted to hydrogen sulphide which, along with hydrogen sulphide already present in the natural gas, is removed. The resulting desulphurised hydrocarbon 16 is combined with a second portion of the by-product gas provided by line 18 and the resulting gaseous mixture fed to a saturator 20 where it is saturated with steam. The saturator 20 is fed with a water stream comprising a heated process condensate provided via line 22 and a recirculated water stream recovered from the saturator via line 24. The resulting saturated hydrocarbon is combined with steam provided by line 26 to increase the steam to carbon ratio and the resulting mixture heated and fed via line 28 to a reforming unit 30 comprising a gas-heated reformer (GHR) and autothermal reformer (ATR) in series. The gaseous mixture 28 is subjected to a first step of steam reforming in the gas-heated reformer over a particulate nickel steam reforming catalyst disposed in a plurality of externally heated tubes followed by autothermal reforming with an oxygen gas stream provided by line 32 in an autothermal reformer containing a bed of nickel steam reforming catalyst. Reformed gas recovered from the autothermal reformer is used to heat the tubes of the gas-heated reformer thereby producing a partially cooled reformed gas 34, which, without further steam addition, is fed to a water-gas shift unit 36 comprising an isothermal water-gas shift vessel containing a bed of water-gas shift catalyst cooled by boiling water under pressure, such as demineralised boiler feed water. The exothermic water-gas shift reaction generates steam, a portion of which is recovered and fed from the water gas shift unit 36 via line 26 to the saturated hydrocarbon feed. Within the water gas shift unit 36, shifted gas recovered from the isothermal water-gas shift vessel is cooled and fed to an adiabatic low-temperature water-gas shift vessel containing a bed of low-temperature shift catalyst to reduce the carbon monoxide level in the resulting hydrogen-enriched gas. A hydrogen-enriched gas mixture is recovered from the water-gas shift unit 36 and fed via line 38 to a heat recovery unit 40 where it is cooled in heat exchange with boiler feed water, saturator feed water and process condensate to below the dew point. Process condensate is recovered in two gas-liquid separators in series optionally with intermediate cooling in heat exchange with liquid amine absorbent from a downstream carbon dioxide separation unit 58, to produce a dewatered hydrogen-enriched gas. Process condensate recovered from the gas-liquid separators in the heat recovery unit 40 is passed to the saturator feed 22 via line 42. The dewatered hydrogen-enriched gas is passed from the heat recovery unit 40 via line 44 to an oxidation unit 46 comprising an oxidation vessel containing a fixed bed of a precious metal CO-oxidation catalyst. A further oxygen-rich gas stream is added via line 48 to the oxidation vessel. Residual carbon monoxide is oxidised to form a carbon dioxide-enriched gas mixture. The carbon dioxide-enriched gas mixture recovered from the oxidation unit 46 via line 50 is subjected to heat recovery in a heat recovery unit 52 in heat exchange with demineralised water, process condensate and the feed gas to the desulphurisation unit to cool it to below the dew point. Process condensate is recovered in two gas-liquid separators in series optionally with intermediate cooling in heat exchange with ambient temperature air and water, to produce a dewatered carbon dioxide-enriched gas. Process condensate recovered from the gas-liquid separators in the heat recovery unit 52 via 54 is mixed with the process condensate recovered in line 42 from heat recovery unit 40, heated, and passed to the saturator via line 22. The de-watered carbon dioxide-enriched gas recovered from the heat recovery unit 52 is fed via line 56 to a product separation unit 57 comprising a carbon dioxide removal unit 58 and a hydrogen purification unit 64. The carbon dioxide removal unit 58 operates by means of reactive absorption with an amine absorbent, which removes carbon dioxide from the gas mixture, thereby producing a carbon dioxide stream 60 which is compressed, preferably dehydrated, and sent for storage or conversion into chemicals. By removing carbon dioxide from the feed gas 56, the carbon dioxide removal unit 58 produces a crude hydrogen gas stream which is recovered from the carbon dioxide removal unit via line 62. The crude hydrogen stream 62 is passed to a hydrogen purification unit 64 comprising a pressure-swing absorption unit that purifies the hydrogen by removing trace amounts of methane and carbon oxides as well as any residual inert gases such as nitrogen and argon to provide a high-purity hydrogen product stream that is recovered from the purification unit 64 by line 66. The hydrogen purification unit 64 produces a by-product gas stream recovered via line 68 that contains inert gases, methane and carbon oxides as well as some hydrogen. The by-product gas stream 68 is optionally compressed then divided; part is fed via line 70 to a steam boiler 72 to generate steam for the process via line 74. The remaining portion of by-product gas is compressed and recycled to the process via line 76 by dividing it into streams 12 and 18 that are combined with the natural gas fed to the reforming unit. In Figure 2, the process is similar to that depicted in Figure 1, except: (i) the reforming unit 30 comprises an adiabatic pre-reformer and autothermal reformer in series with direct steam addition and with by-product stream 24 recycled to between the pre-reformer and the autothermal reformer in the reforming unit 30; (ii) the water gas shift unit 36 comprises a high-temperature shift vessel and low-temperature shift vessel in series with addition of steam to the reformed gas mixture upstream of the water-gas shift unit; and (iii) the product separation unit 57 comprises a hydrogen separation unit coupled to a downstream cryogenic carbon dioxide separation unit. In Figure 2, a natural gas stream 10 is combined with a first portion of a by-product gas stream supplied via line 12, and the resulting mixture heated and passed to a desulphurisation unit 14 comprising a hydrodesulphurisation vessel containing a hydrodesulphurisation catalyst followed by two vessels containing zinc oxide hydrogen sulphide absorbents and ultra-purification adsorbents. Sulphur compounds present in the natural gas are converted to hydrogen sulphide which, along with hydrogen sulphide already present in the natural gas, is removed. The resulting desulphurised hydrocarbon 16 is fed to a reforming unit 30 comprising an adiabatic pre-reformer and autothermal reformer in series. Upstream of the adiabatic pre-reformer, the desulphurised hydrocarbon 16 is combined with steam fed via line 80. The resulting gaseous mixture of hydrocarbon and steam is passed adiabatically through the pre-reformer containing a bed on nickel pre-reforming catalyst to convert higher hydrocarbons to methane and form a pre-reformed gas mixture comprising methane, steam, hydrogen and carbon oxides. The prereformed gas mixture is combined with a second portion of the by-product gas stream fed to the reforming unit 30 via line 18 and the combined mixture fed to the autothermal reformer containing a bed of nickel steam reforming catalyst, where it is autothermally reformed with an oxygen gas stream provided by line 82. A reformed gas recovered from the autothermal reformer is fed via line 84 to a waste-heat boiler 86. Demineralised boiler feed water fed to the boiler 86 via line 88 is converted to steam which is fed via line 90 to form part of the steam stream 80 fed to the reforming unit 30. The cooled reformed gas recovered from boiler 86 is combined with steam, suitably a portion of steam from line 80 provided by line 91, and the steam-enriched reformed gas mixture fed via line 92 to a water-gas shift unit 36 comprising an adiabatic high-temperature shift vessel containing a bed of high-temperature water-gas shift catalyst. A shifted gas recovered from the high-temperature water-gas shift vessel is cooled and fed to an adiabatic low-temperature water-gas shift vessel containing a bed of low-temperature shift catalyst to reduce the carbon monoxide level in the resulting hydrogen-enriched gas. A hydrogen-enriched gas mixture is recovered from the water-gas shift unit 36 and fed via line 38 to a heat recovery unit 40 where it is cooled in heat exchange with prereformed feed, purified feed, and boiler feed water, to below the dew point. Process condensate is recovered in one, two or more gas-liquid separators in series to produce a dewatered hydrogen-enriched gas. Process condensate recovered from the gas-liquid separators in the heat recovery unit 40 is passed via line 42 to a steam stripping unit 92 fed with steam via line 94. In the stripping unit 92 the condensate is stripped with the steam 94 to form an enriched steam stream containing organic compounds that were dissolved in the condensate. The enriched steam stream is fed via line 96 to form part of the steam 80 fed to the reforming unit 30. If desired, a portion of the steam 94 may be combined with the enriched steam stream 96. The dewatered hydrogen-enriched gas is passed from the heat recovery unit 40 via line 44 to an oxidation unit 46 comprising an oxidation vessel containing a fixed bed of a precious metal CO-oxidation catalyst. A further oxygen-rich gas stream is added via line 48 to the oxidation vessel. Residual carbon monoxide is oxidised to form a carbon dioxide-enriched gas mixture. The carbon dioxide-enriched gas mixture recovered from the oxidation unit 46 via line 50 is subjected to heat recovery in a heat recovery unit 52 in heat exchange with demineralised or cooling water, process condensate, the feed gas to the desulphurisation unit, and cooling air to cool it to below the dew point. Process condensate is recovered in one, two or more gas-liquid separators in series to produce a dewatered carbon dioxide-enriched gas. Process condensate recovered from the gas-liquid separators in the heat recovery unit 52 via line 54, is mixed with the process condensate recovered in line 42 from heat recovery unit 40 and passed to the steam stripping unit 92. The de-watered carbon dioxide-enriched gas recovered from the heat recovery unit 52 is fed via line 56 to a separation unit 57 comprising a hydrogen separation unit 100 coupled to a downstream cryogenic carbon dioxide separation unit 102. The hydrogen separation unit operates by pressure-swing adsorption (PSA) to produce a hydrogen product stream recovered via line 104 and a hydrogen-depleted gas, which is potentially compressed, dried and passed via line 106 to the cryogenic carbon dioxide separation unit 102. The cryogenic carbon dioxide separation unit cools and condenses a liquid carbon dioxide stream, which is recovered from the unit 102 via line 108. The cryogenic carbon dioxide separation also produces a carbon dioxide-depleted gas stream, which is recovered from the unit 102 via line 110. Rather than recycle the carbon-dioxide depleted gas from the cryogenic carbon dioxide separation unit 102 to the process directly, in this arrangement, the carbon dioxide-depleted gas stream 110 is fed to a downstream secondary purification unit 112 to increase hydrogen recovery and make the recycle stream richer in carbon compounds. Hence the separation unit comprises the hydrogen separation unit 100, the cryogenic carbon dioxide separation unit 102 and the secondary purification unit 112. The secondary purification unit may be a further pressure-swing adsorption unit that separates the carbon dioxide-depleted gas 110 into a further hydrogen product stream 114, which is combined with the hydrogen product stream 104 from the hydrogen separation unit 100 to form a combined hydrogen product stream 116. The secondary purification unit may also produce a recycle stream 111, which may be fed to the cryogenic carbon dioxide separation unit 102 to increase carbon dioxide recovery. The secondary purification unit 112 also produces a carbon-containing by-product gas stream, which is recycled to the process via line 118. The byproduct stream 118 is compressed, divided into the by-product streams 12 and 18 and recycled to the process. In this arrangement, the hydrogen separation unit 100 is operated to produce a fuel gas stream, which is recovered from the unit 100 and fed via line 120 to a steam boiler 72 to generate steam for the process via line 74. The invention is further illustrated by reference to the following calculated example of a process in accordance with the flowsheet depicted in Figure 1. Stream Number 10 12 16 18 22 24 26 28 Mass Flow t / h 56.91 0.47 57.38 0.95 634.8 531.8 44.2 209.6 Temperature °C 17 119 241 119 246 171 255 440 Pressure bara 39.5 40.0 37.9 40.3 42.5 37.5 43.0 36.9 Molar Composition Unit Methane mol% 91.79 4.71 89.62 4.71 0.11 0.24 25.41 Ethane mol% 3.00 2.93 0.83 Propane mol% 0.59 0.58 0.16 Butanes mol% 0.14 0.14 0.04 Pentanes mol% 0.04 0.04 0.01 Hexanes mol% 0.02 0.02 0.01 Hydrogen mol% 84.25 2.00 84.25 1.19 Carbon Dioxide mol% 0.77 0.07 0.64 0.07 0.19 Carbon Monoxide mol% 0.11 0.03 Oxygen mol% Nitrogen mol% 3.64 10.09 3.8 10.09 1.15 Argon mol% 0.50 0.01 0.50 0.01 Water mol% 0.34 0.12 0.34 99.86 99.75 100.0 70.91 Methanol mol% 0.01 0.01 0.02 0.01 0.05 Ammonia mol% 0.03 0.03 0.01 0.01 Stream Number 32 34 38 44 48 50 56 60 Mass Flow t / h 47.58 257.2 257.2 206.4 1.19 207.6 169.9 146.6 Temperature °C 244 257 204 137 15 175 45 45 Pressure bara 34.2 29.1 26.9 25.7 38.0 24.8 23.7 1.3 Molar Composition Unit Methane mol% 0.30 0.30 0.36 0.36 0.41 0.01 Ethane mol% Propane mol% Butanes mol% Pentanes mol% Hexanes mol% Hydrogen mol% 43.39 54.32 64.15 63.91 73.81 0.70 Carbon Dioxide mol% 6.57 17.57 20.74 20.98 24.23 94.59 Carbon Monoxide mol% 11.24 0.20 0.24 0.00 0.00 Oxygen mol% 97.66 99.50 Nitrogen mol% 0.74 0.74 0.88 0.88 1.01 0.01 Argon mol% 0.49 0.04 0.04 0.05 0.50 0.05 0.06 Water mol% 1.85 37.71 26.75 13.52 13.76 0.44 4.57 Methanol mol% 0.05 0.06 0.06 0.03 0.11 Ammonia mol% 0.01 0.01 0.01 0.01 0.01 0.01 Stream Number 62 66 68 70 Mass Flow t / h 25.41 18.58 6.83 5.84 Temperature °C 45 45 45 40 Pressure bara 23.3 22.0 1.3 3.73 Molar Composition Unit Methane mol% 0.54 4.63 4.63 Ethane mol% Propane mol% Butanes mol% Pentanes mol% Hexanes mol% Hydrogen mol% 97.83 99.78 82.23 83.23 Carbon Dioxide mol% 0.01 0.07 0.07 Carbon Monoxide mol% Oxygen mol% Nitrogen mol% 1.34 0.19 9.96 9.96 Argon mol% 0.08 0.03 0.49 0.49 Water mol% 0.18 1.57 1.57 Methanol mol% 0.01 0.01 Ammonia mol% 0.03 0.03 The Example provides a carbon capture of 98.4% and feedstock efficiency of 82.7% (on LHV basis) compared to typical values of 95.4% and 80.8% (on LHV basis) without by-product gas recycle and the oxidation unit. This corresponds to a total carbon intensity (Scope 1+2+3) of 12.64 g CO2 / MJ H2 (LHV) based on v3 (Dec 2023) based on UK Low Carbon Hydrogen Standard compared to a typical value of 15.31 g CO2 / MJ H2 (LHV) without by-product gas recycle and the oxidation unit. The combination of by-product gas recycle and oxidation of carbon monoxide also provides flexibility in case of upstream reforming or water-gas shift operating underperformance, or deviations in the feedstock composition that could lead to higher carbon monoxide slip from the water-gas shift section.

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