Plant and process for producing hydrogen from hydrocarbons with reduced co2 emissions

Abstract

Hydrogen production plant comprising: - At least one reformer (30) for converting a stream comprising a hydrocarbon feedstock (1) through conversion with steam rich in oxygen (02) into a reformed gas stream (32) comprising hydrogen, carbon monoxide, carbon dioxide and at least one hydrocarbon as impurity, said reformer (30) comprising exothermic, oxygen-based autothermal (ATR) or partial oxidation (POX) reforming and a heat recovery section (40), - A fired heater (90) configured to preheat the stream comprising the hydrocarbon feedstock (1) before entry into the at least one reformer (30), - At least one water gas shift (WGS) reactor (50) for converting the carbon monoxide of the reformed gas stream (32) into a shifted gas stream (51) containing additional carbon dioxide and hydrogen, - A Hydrogen and Carbon Dioxide Recovery Unit (60) located downstream of the WGS reactor (50) and configured to remove carbon dioxide and hydrogen from the shifted gas stream (51), and to produce a first product stream (61) enriched in carbon dioxide and a second product stream (62) enriched in hydrogen, a waste stream (63) depleted in both hydrogen and carbon dioxide, - A compressor (70) for compressing a part (65) of the waste gas stream (63) from hydrogen and carbon dioxide recovery unit (60) into a compressed gas stream (71), - A membrane separation system (80) selective for the permeation of hydrogen configured to be fed with the compressed gas stream (71) and to produce a hydrogen- enriched permeate (82) stream and a hydrocarbon-enriched retentate (81) stream, - A passageway for feeding at least part of the hydrogen-enriched permeate (82) to the fired heater (90) to be used as a low-carbon fuel by the fired heater (90), and - A passageway for recycling the hydrocarbon-enriched retentate (81) to the hydrocarbon feed (1) via a pipeline (83) and / or to the reformer (30) via the pipeline (85) and / or to the inlet to the water gas shift reactor (50) via the pipeline (87).

Description

[0001] Title: Plant and process for producing hydrogen from hydrocarbons with reduced CO2emissions

[0002] The present invention relates to a hydrogen plant and a process for producing a hydrogen-comprising product gas from a hydrocarbon feedstock with reduced CO2 emissions implementing this plant.

[0003] In particular, the present invention concerns a plant and process for producing hydrogen from a hydrocarbon feed, in which this hydrocarbon feed is subjected to reforming using a reformer using Autothermal Reforming (ATR) with optional prereforming for generating a synthesis gas, that is subjected to water gas shift conversion to increase the conversion of hydrocarbon feed to hydrogen and carbon dioxide, that are recovered in a hydrogen and carbon dioxide recovery section, producing two products streams, one of which is hydrogen-rich and one of which is carbon dioxide-rich, and a waste stream that is depleted in hydrogen and carbon dioxide, and where all or at least a part of this waste stream are compressed to a membrane separation system selective for hydrogen, such that the hydrogen rich permeate product stream from the membrane separation system is used as low-carbon fuel for a fired heater while the hydrocarbon-enriched retentate stream from the membrane separation system is recycled to at least partly the pre-reformer feed and / or partly to the reformer feed and / or partly to the water gas shift section.

[0004] Plants for producing hydrogen from hydrocarbons, in particular steam methane reforming (SMR) plants, are widely applied in refinery complexes to supply hydrogen for upgrading of several products, for example in hydrocracking, hydrogenation or hydrodesulphurization. Additionally, hydrogen is used as a component of syngas (a mixture comprising hydrogen and carbon monoxide). Syngas is an essential building block to produce for example ammonia, methanol, synthetic fuels and many different chemicals.

[0005] Autothermal Reforming (ATR) is widely applied forthe generation of syngas and, in combination with integrated CO2 capture, is developing as an alternative to SMR based processes for the generation of hydrogen as well. Furthermore, there is growing interest in using hydrogen as such as an alternative to petroleum-based fuel in energy sector as a means to provide seasonal energy storage, in industry to provide high quality heat and in mobility mostly for heavy and long hauls transportation means. It is estimated that about 95% of the global hydrogen supply is produced from fossil fuels; as by-product of the technology, CO2 is produced and emitted to the atmosphere. The CO2 is produced not only by combustion of a carbon-based fuel for heating the feedstock to the temperatures needed to carry out the reforming, but CO2 is also formed as a side-product in the hydrogen production: the steam reforming reaction produces carbon monoxide (with methane as a starting compound: CH4 + H2O CO +3 H2), which is subsequently converted to carbon dioxide via the water gas shift reaction (CO + H2O CO2 + H2). The steam reforming reaction is highly endothermic and requires a significant additional heat input. In processes based on exothermic reforming, such as autothermal reforming (ATR) or partial oxidation reforming, are carried out by a partial combustion of the hydrocarbon feedstock (with methane as starting component: 2 CH4 + 02 2C0 +4 H2) , such that a small-size fired heater is required only for preheating the hydrocarbon feedstock to the reactor inlet temperature but not for supplying the heat of reaction for the reforming reaction. Exothermic reforming processes thereby consume oxygen 02 that is an additional feed to the plant. The formed CO is also subsequently converted to C02 through the water gas shift reaction.

[0006] In recent years an increasing industrial focus on environmental emissions and reduction of the carbon footprint challenges the design of hydrogen production facilities (a.k.a. Hydrogen Production Units: HPU's) to reduce carbon footprint as well.

[0007] Traditionally, hydrogen plants were integrated with a refinery or industrial complex where the generated excess steam could be used, since the hydrogen plant is considered as an efficient steam producer, in order to recover as much low grade heat as possible, the excess steam generated was optimized to meet the external steam demand, making the excess steam a valuable by-product of the facility. Consequently, the firing and thus C02 emissions were not the main design parameters but this changes with the evolving regulations on greenhouse gas emissions. Many different solutions have been proposed to reduce the steam production of the hydrogen plant by reducing the firing demand. The options include preheating the combustion air (up to typically 600°C) with flue gas or another indirect heat source, preheating fuel and / or tail gas, applying an adiabatic (pre)reforming step. All these solutions reduce the required firing demands and thereby also the export steam flowrate and CO2 emissions. However, the majority of the CO2 emitted through the flue gas comes from the tail gas comprising methane, and further residual hydrogen, carbon monoxide and carbon dioxide produced as part of the reforming reaction. This tail gas is typically generated as the waste stream of the final product purification, carried out by an Adsorption-based process (e.g., Pressure Swing Adsorption, PSA). In recent years, with more focus on CO2 emissions, a commonly applied solution is to capture CO2 from the syngas, downstream of the (final) shift reactor. Hydrogen product that is this obtained is generally referred to as blue hydrogen. The total reduction of CO2 emissions is limited by the CH4 and CO slip from the reformer system and shift section, respectively, on the process side and the amount of additional fossil fuel fed as make-up fuel to meet the heat duty requirement of the reforming process. The methane slip may be reduced by operating at high autothermal reformer outlet temperature, which requires more oxygen and firing.

[0008] So, a continuing need exists to provide alternative processes and equipment to reduce greenhouse gas emissions (carbon dioxide, methane), e.g. to allow refurbishment of existing HPU's as well as new HPU's. In particular, a continuing need exists to provide an efficient way to produce hydrogen from hydrocarbon feeds by reforming processes, whereby greenhouse gas emissions (carbon dioxide and / or methane) are (further) reduced and / or whereby global carbon footprint is (further) reduced. In particular, a primary objective is to reduce the hydrocarbon consumption (i.e. use of hydrocarbon for other purposes than generating hydrogen product gas from it). Enabling re-use of unconverted carbon in the process instead of using it as a fuel will further reduce the direct carbon dioxide emissions from the process. Oxygen-based reforming processes require the addition of a stream that is rich in oxygen 02. Processes used for the generation of this stream may be Pressure Swing Adsorption, Vacuum Pressure Swing Adsorption, Membrane separation or cryogenic fractionation of air, all of which require ambient air to be compressed to elevated pressure prior to the actual separation step. Oxygen 02 can also be generated by electrolysis of water. All processes available for the generation of high-purity oxygen are therefore high in electrical power consumption, typically ranging from 0.38 to 0.75 kWh / kg 02, depending on purity, pressure, capacity and selected technology. A secondary objective, in particular for oxygen-based reforming processes, is therefore the reduction of the electrical power consumption and its associated CO2 emissions, or, in case low-carbon intensity renewable electrical power is already sufficiently available, utilize electrical power to reduce the total hydrocarbon consumption of the reforming processes. The present invention addresses both objectives.

[0009] As the oxygen-based reforming process provides the majority of the heat of reaction required for the catalytic conversion, it is found that the heat input provided by the tail gas from the final product purification often contains too much calorific value. When a CO2 removal system is included, the tail gas stream is generally high in hydrogen content (50 to 75 % mole) and concentrates the unconverted carbon molecules in the form of methane CH4 and carbon monoxide CO as the balancing impurities, together with the inert gases from the feed streams (nitrogen, argon and / or helium) and water vapour. In case no CO2 removal is applied, the tail gas stream will contain between 20 and 50 % mole of CO2, correspondingly lowering the concentrations of the other components without affecting the stream's heating value. To avoid excessive firing not required by the process when using the stream as fuel in the fired heater, another use needs to be found for this stream. The document W02022 / 038089 describes a plant in which the tail gas is therefore, at least partly, recycled back to the inlet of the ATR reformer, optional pre-reforming and / or water gas shift section. The document WO2021 / 073834 describes a plant in which the tail gas is submitted to an additional hydrogen recovery step to reduce the calorific value of the tail gas to the fired heater. This can be considered a more generic version of the document WO2020 / 221642 that describes a plant in which the tail gas is compressed and submitted to a membrane separation section, recycling the hydrogen-rich product stream from the membrane to the hydrogen purification unit and using the hydrocarbon-enriched product stream from the membrane as fuel to the fired heater. The document EP 3988502 describes a plant based on SMR reforming that uses part of the waste gas as fuel in the fired heater, recycles part of the waste gas as feed to the reformer and finally exports the excess waste gas as a fuel stream to be used outside the battery limits of the hydrogen plant.

[0010] Recycling high concentrations of a reactant to a reaction section generating this reactant as a product always has an adverse effect on the reaction equilibrium. As the reforming and water gas shift reactions are highly equilibrium-controlled, recycling the waste gas stream, high in reactant H2, to the reforming and / or water gas shift reaction sections, will adversely affect the reaction equilibrium, although the recycling of the carbon molecules will benefit the overall feed consumption. The efficiency of the processes described in the documents WO2020 / 221642 and WO2021 / 073834 therefore benefits from the additional hydrogen recovery steps on the waste gas stream as hydrogen reactant is removed from the stream recycled to the reaction sections.

[0011] The excess heat from the process gas and flue gas is recovered in the hydrogen plant by the generation of steam. The reforming and water gas shift reactions consume steam as a reactant. As the generation of steam requires heat that is provided by the conversion of hydrocarbon feedstock, reducing the steam consumption as required by the reforming and water gas shift reactions will increase the efficiency of the process. Typically, the steam required for the reforming and water gas shift reactions is added to the hydrocarbon feed at the feed inlet of the plant, thereby sending sometimes larger than required steam quantities through the upstream sections of the plant. The document WO2020 / 221642 describes a plant in which the steam addition to the hydrocarbon feed is split per section, adding steam to the reforming section and water gas shift sections as required by the individual sections. This reduces the steam consumption by the process and also reduces the heat available for generating steam. An optimized plant therefore only produces the steam it requires for the process and does not produce export steam that is not consumed in the process and aims to inject the steam into the process at the location where it is required. An optimized steam generation system utilizes the available heat with maximum efficiency and without temperature pinches. A further optimization that can optionally be included in particular embodiments of this invention, in particular when electrical power has a high cost or associated carbon intensity, can be achieved by increasing the pressure in the steam system to maximize the heat intake from the process gas and flue gas, recovering the energy stored in the steam as pressure by expanding the steam over a turbine to the pressure required by the process. This approach reduces the energy losses in the heat recovery section and reduces the net imported electrical power required by the process.

[0012] For plants in which the end user has no destination for any excess generated export steam, a balanced steam production can be targeted (i.e. no steam export and only generating the steam consumed within the HPU). Steam production can be reduced via the process changes already described (combustion air preheat, fuel preheat, prereforming) but these steps typically do not allow a reduction to zero export steam. In order to achieve zero export steam, heat normally used for generating export steam is used for additional reforming instead of firing extra duty or combusting additional feedstock . This invention achieves this target by the inclusion of a heat exchanger reformer, as described in the document WO2018 / 104526. This heat exchanger reformer can be installed in parallel similar to the description of the document WO2018 / 104526 or in series to the main reformer as described in the document W02011 / 077107 or the document WO2012 / 057922. A heat exchanger reformer utilizes the heat of the hot process gas flowing through the shell of a shell and tube heat exchanger reactor as the heat source to colder hydrocarbon feedstock flowing over reforming catalyst contained in the tubes of the shell and tube heat exchanger reactor, thereby providing additional hydrocarbon feedstock conversion and reducing the amount of heat available for generating steam until the quantity of steam generated by process matches the process steam demands, realizing a zero export steam plant.

[0013] Steam export and feed conversion efficiency are only two of the important key performance indicators of a modern hydrogen production unit. The carbon dioxide emissions from the stack of the hydrogen production unit (scope 1), the associated carbon dioxide emissions generated during the generation of the electrical power consumed by the unit (scope 2), and the associated carbon dioxide emissions from residual carbon remaining in the hydrogen product or the carbon dioxide emissions in the life cycle of the feedstock of the hydrogen production unit (scope 3) are another important key performance indicator. Typically, the carbon intensity is used as a parameter, expressing the total equivalent CO2 emissions per kilogram of hydrogen generated by the plant. The carbon intensity can be expressed for scope 1, scope 1+2, scope 1+3 or scope 1+2+3 emissions, depending on the accurateness and availability of data to calculate the scope 2 and scope 3 emissions from the plant. To reduce the scope 1 emissions, more CO2 should be captured in the plant's CO2 capture unit and methane and carbon monoxide slip into the fuel for the fired heater should be minimized. To reduce the scope 2 emissions, less power consumption is favoured. To reduce the scope 3 emissions, a low residual content of carbon molecules (methane, carbon monoxide and carbon dioxide) in the hydrogen product are preferred. Using a Pressure Swing Adsorption (PSA) based hydrogen purification step is required to achieve this target. PSA processes will generate the waste gas stream containing the residual carbon molecules as described earlier. To reduce scope 1 emissions, the process should utilize only the hydrogen from the waste stream as fuel while recycling the carbon molecules as feed for the reformer. The processes described by the documents WO2021 / 073834 and WO2020 / 221642 utilize a hydrogen-depleted, hydrocarbon-enriched stream obtained from the waste gas of the PSA unit as fuel to the fired heater, thereby utilizing the unconverted carbon from the hydrocarbon feed as fuel instead of recycling it as feed to reduce the overall feed consumption. The present invention therefore extracts the hydrogen from the waste gas of the PSA unit to use it as low-carbon fuel, reducing the scope 1 CO2 emissions from the combustion, while recycling the carbon molecules back to the reformer, improving the feed conversion.

[0014] The oxygen 02 required for the internal partial combustion reaction of the oxygen-based reforming process needs to be supplied externally and typically requires high-purity oxygen to avoid contamination of the produced syngas with excess Nitrogen. High-purity oxygen generation by air separation is the most commonly used process for this purpose. While different technologies exist for this process (Pressure Swing Adsorption, Vacuum Swing Adsorption, Membranes, Cryogenic Fractionation), they all require compression of ambient air as the source of oxygen, which makes the air separation process very energy consuming. While a reduction in fuel consumption is beneficial for the direct C02 emissions from the stack of the fired heater (so-called scope 1 emissions), the substantial increase in electrical power consumption for the oxygen generation comes with a potential larger increase in C02 emissions associated with the generation of the electrical power (so-called scope 2 emissions), depending on the source of the electrical power. For this reason, the total CO2 emissions from oxygen- based reforming processes are favoured by regions with high contribution of electricity production from renewable sources.

[0015] In regions where the carbon intensity of generated electricity is relatively high, any reduction of the net imported power consumption will benefit the overall carbon intensity of the process. Plant optimization to generate electrical powerwithin the plant by the generation of higher pressure steam and recovering the power in a steam turbine while expanding the steam to the process pressure as described by the present invention can be beneficial under these conditions.

[0016] Another recent approach at reducing the CO2 emissions, or more specifically, the carbon intensity (the total amount of CO2 generated by the process, direct or indirect, per amount of hydrogen generated of the reforming processes), is the use of non-fossil or e-hydrocarbons as feed to the process. Non-fossil sources of hydrocarbon feed can be obtained from the waste hydrocarbon containing streams generated by the hydroprocessing of vegetable oils and / or other sources of long hydrocarbon-chain containing organic molecules. E-Hydrocarbons are hydrocarbon molecules generated by a process combining captured CO2 with H2 generated from renewable power, that are used as an energy-carrier to be converted back into H2 when storage and / or transportation of H2 are considered impractical at the location where the H2 is generated. As the carbon contained in the hydrocarbon molecules used as feedstock to the reforming process is of non-fossil or re-captured origin, it can be considered neutral when converted into CO2 by the reforming process in terms of CO2 emissions. Capturing the CO2 generated by the process would enable even a negative carbon-intensity of the reforming processes.

[0017] The present invention aims to achieve the objectives of increasing reforming efficiency, reducing CO2 emissions from stack, reducing fuel consumption and reducing electrical power consumption by making several improvements to the existing art of hydrogen production using oxygen-based reforming, in particular auto-thermal reforming (ATR). The main object of the invention is the addition of a hydrogen-selective membrane separation system on the waste stream of the Hydrogen and Carbon Dioxide Recovery Unit and in particular the use of the hydrogen-rich permeate stream as low- carbon fuel, while recovering the retentate stream rich in carbon molecules as feed to the reformer. Reformer in this text refers to a reforming reactor, utilizing exothermic, oxygen-based reforming, preferably Autothermal Reforming (ATR) or Partial Oxidation (POX). This enables lowering CO2 stack emissions by eliminating the carbon in the fuel via said membrane separation system selective to hydrogen and the use of the resulting hydrogen-rich permeate stream as low-carbon fuel. It also improves reforming efficiency by the recycle (as feed to the reformer) of the carbon otherwise used as fuel in the fired heater. Additional fuel consumption and export steam production reduction is achieved by the use of heat-exchanger reforming in addition to the auto-thermal reforming with optional pre-reformer. The heat recovery and additional hydrocarbon conversion achieved by the heat exchanger reformer reduces the fuel requirements and increases reforming efficiency. This further reduces the oxygen consumption of the ATR and thereby the power requirements associated with the oxygen generation. Optimization of the heat recovery system reducing temperature pinches while generating process steam at higher pressure, enables power recovery by steam expansion and reduces the net imported electrical power consumption. It is an object of the present invention to address one or more of said needs. One or more alternative or additional objects which may be addressed follow from the description below.

[0018] An object of the invention is a hydrogen production plant comprising:

[0019] At least one reformer 30 for converting a hydrocarbon feedstock 1 through conversion with steam rich in oxygen 02 into a reformed gas stream 32 comprising hydrogen, carbon monoxide, carbon dioxide and at least one hydrocarbon as impurity, said reformer 30 comprising oxygen-based reforming and a heat recovery section 40,

[0020] At least one water gas shift (WGS) reactor 50 for converting the carbon monoxide of the reformed gas stream 32 into a shifted gas stream 51 containing additional carbon dioxide and hydrogen,

[0021] A Hydrogen and Carbon Dioxide Recovery Unit 60 located downstream of the WGS reactor 50 and configured to remove carbon dioxide and hydrogen from the shifted gas stream 51, and to produce a first product stream 61 enriched in carbon dioxide and a second product stream 62 enriched in hydrogen, a waste stream 63 depleted in both hydrogen and carbon dioxide,

[0022] A compressor 70 for compressing a second part 65 of the waste gas stream 63 from hydrogen and carbon dioxide recovery unit 60 to a compressed gas stream 71,

[0023] A membrane separation system 80 selective for the permeation of hydrogen configured to be fed with the compressed gas stream 71 and to produce a hydrogen- enriched permeate 82 stream and a hydrocarbon-enriched retentate 81 stream, A passageway for feeding at least part of the hydrogen-enriched permeate 82 to the fired heater 90, and

[0024] A passageway for recycling the hydrocarbon-enriched retentate 81 to the hydrocarbon feed 1 via a pipeline 83 and / or to the reformer 30 via the pipeline 85 and / or to the inlet to the water gas shift reactor 50 via the pipeline 87.

[0025] Depending on the embodiment, the hydrogen plant according to the present invention can comprise one or more of the following features: the reformer 30 comprises exothermic, oxygen-based autothermal (ATR) or partial oxidation (POX) reforming, there is provided a fired heater 90 configured to preheat the stream comprising the hydrocarbon feedstock 1 before entry into the at least one reformer 30, there is provided at least one fired heater 90 configured to preheat the stream comprising the hydrocarbon feedstock 1 before entry into the at least one reformer 30, the passageway is arranged to feed at least part of the hydrogen-enriched permeate 82 to the fired heater 90 to be used as a low-carbon fuel by the fired heater 90, the plant comprises a passageway for feeding a fired heater 90 with a first part 64 of the waste gas stream 63 from hydrogen and carbon dioxide recovery unit 60. the Hydrogen and Carbon Dioxide Recovery Unit 60 is configured to produce a flash gas stream 69, and / or a stream 68 depleted in carbon dioxide and rich in hydrogen, the fired heater 90 receiving its fuel from at least one of the following streams: a) At least a part 64 of the waste gas stream 63 from the hydrogen and carbon dioxide recovery unit 60; b) At least a part of the hydrogen-enriched permeate 82 produced by the hydrogen-permeating membrane separation system 80; c) At least a part of the hydrogen-enriched product 62 from the hydrogen and carbon dioxide recovery unit 60; d) At least a part of the hydrocarbon feedstock stream 1; e) A make-up fuel stream imported from a battery limit; f) At least a part of the flash gas stream 69 and g) At least a part of the optional stream 68. the heat recovery section 40 is configured to generate a steam stream 41 and the plant comprises means for routing at least a part of this steam stream 41: a) As process steam 45 to the inlet of the reformer 30 and / or b) As process steam 43 to the inlet of the water gas shift reactor 50 c) As export steam 46 to a battery limit. the hydrogen plant comprises a feed purification section 10 upstream of the reformer 30, configured to produce a treated hydrocarbon stream 11. the heat recovery section 40 is configured to generate a steam stream 41 and the hydrogen plant comprises: At least one pre-reformer reactor 20 upstream of the reformer 30, configured to produce a pre-reformed syngas stream 21 from at least one or both of the treated hydrocarbon stream 11 and a part of the hydrocarbon- rich stream 81 via means 84, and Means for routing at least a part 44 of steam stream 41 to the inlet of the pre-reformer 20. Preferably, the hydrogen plant comprises means for recycling a part 44 of the steam stream 41 produced by the heat recovery section 40 and / or means for recycling a part 44 of the steam stream 41 produced by the heat recovery section 40. the heat recovery section 40 is configured to generate a steam stream 41 and the plant comprises a heat exchanger reformer 35 installed in series to the reformer 30 to receive the at least one of the hydrocarbon feed 11, the pre-reformer feed 23 or part of the hydrocarbon-rich stream 81 provided via means 86, mixed with steam 42 to the tube side inlet in order to produce a reformed stream 37 from the tube side outlet; means for feeding this reformed stream 37 to the reformer 30 to produce a reformed stream 32, with the reformed stream 32 configured to enter the shell side inlet of heat exchanger reformer 35 and provide the heat of reaction to the tube side of the heat exchanger reformer 35 and to produce a reformate stream 36 from the shell side outlet of the heat exchanger reformer 35, means for sending the reformate stream 36 to the heat recovery section 40 and means for mixing at least a part 42 of steam stream 41 to the tube inlet of the heat exchanger reformer 35. the heat recovery section 40 is configured to generate a steam stream 41; The plant comprises means for dividing the reformer feed stream 21 into a first part 23 and a second part 22; The reformer 30 is configured to receive the second part 22 and to produce the reformed gas stream 32; The plant comprises a heat-exchanger reformer 35 installed in parallel to the reformer 30 and configured to receive at least one or both of the part 23 of the reformer feed stream 21 and at least part of the hydrocarbon-rich stream 81 through means 86, as well as the reformed gas stream 32, to produce a reformate stream 36, and means for mixing the reformed feed gas stream 32 from the shell side of the heat exchanger-reformer 35 with the outlet of the heat exchanger reformed gas from the tube side of the heat exchanger-reformer 35, inside or outside of the heat exchanger-reformer 35, and means for mixing at least a part 42 of steam stream 41 to the tube inlet of the heat exchanger reformer 35; and the heat recovery section 40 is configured to receive reformate stream 36.

[0026] In an embodiment, there is provided a means for mixing the reformed feed gas stream 32 into the shell side of the heat exchanger-reformer 35 with the outlet of the heat exchanger reformed gas from the tube side of the heat exchanger-reformer 35, inside or outside of the heat exchanger-reformer 35, and means for mixing at least a part 42 of steam stream 41 to the tube inlet of the heat exchanger reformer 35; and the heat recovery section 40 is configured to receive reformate stream 36. The hydrogen production plant comprises multiple heat exchanger reformers 35. the heat recovery section 40 is configured to produce a steam stream 41 and the plant comprises a steam turbine 47 configured to receive at least a part of the steam stream 41 and to produce a lower pressure stream 48. Preferably, the hydrogen plant comprises a passageway for recycling at least a part of the lower pressure stream 48 to the heat exchanger reactor 35, and / or a passageway 43 for recycling at least a part of the lower pressure stream 48 to the water gas shift reactor 50, and / or a passageway 44 for recycling at least a part of the lower pressure stream 48 to the pre-reformer 20, and / or a passageway 45 for recycling at least a part of the lower pressure stream 48 to the reformer 30, and / or a passageway 46 for exporting at least a part of the lower pressure stream 48 outside the plant. the membrane separation system 80 comprises membrane elements based polysulfone, poly-imid, poly-aramid, cellulose acetate, any combination thereof, or other polymeric material, or Palladium sheets, exhibiting a selectivity to preferentially permeate hydrogen to a lower pressure. the hydrogen plant comprises a passageway for importing hydrogen and hydrocarbon containing off-gas 72 from a battery limit as feed to the membrane separation system 80, recovering hydrogen from said off-gas to the hydrogen- enriched permeate 82 and hydrocarbons to the hydrocarbon-enriched retentate 81. the hydrogen plant comprises a passageway to route at least part of the hydrogenrich permeate stream 81 as fuel to the fired heater 90 and means for recycling the remaining part of the hydrogen-rich permeate stream 81 to the inlet of the hydrogen recovery unit inside the Hydrogen and Carbon Dioxide Recovery Unit 60. at least a part 88 of the hydrogen-rich stream 82 from the membrane separation system 80 is sent as a product stream to the plant battery limits,

[0027] The hydrogen production plant further comprises a hydrogen recovery unit positioned between the reformer 30 and the water gas shift reactor 50, the hydrogen recovery unit configured to remove hydrogen from the reformed gas stream 32 before entry into the water gas shift reactor 50,

[0028] The fired heater 90 is provided as a separate component to the at least one reformer 30,

[0029] The reformer may include a reactor vessel that contains a packed bed of catalyst (e.g. reformer catalyst),

[0030] The reactor vessel may contain a packed bed of catalyst installed on the base of the vessel, the hydrocarbon feed stream is obtained by at least one of the following : a) A fossil-based hydrocarbon source, such as natural gas, Liquefied Petroleum Gas, Naphtha or any combination thereof b) Hydrogen-rich off-gases from processing treating fossil-based hydrocarbon streams c) Liquid or vapour product streams generated by the treatment of biogenic streams such as but not limited to Vegetable Oil, Cooking Oil and other similar sources, with a hydrocarbon content similar to Liquefied Petroleum Gas, Naphtha or any combination thereof d) Hydrogen-rich off-gases generated by the processes intended under c) of this claim e) Off-gases generated by fermentation processes containing hydrogen, methane and / or carbon monoxide f) Methane rich gases obtained from biological sources or from landfill gases, g) E-Hydrocarbons obtained through synthesis of captured CO2 with hydrogen obtained from renewable or other sources, such as e-methane, and their derivatives or by-products obtained from synthesizing their derivatives by Fisher-Tropsch synthesis or similar reactions.

[0031] Figure 1 shows a flow scheme of the present invention. Figure 2 comprises the elements of Figure 1, adding a heat-exchanger reformer in parallel. Figure 3 comprises the elements of Figure 1, adding a heat-exchanger reformer in series. Figure 4 comprises the elements of Figure 1, adding a steam turbine for power generation.

[0032] Advantageously the waste gas stream 63 comprises H2, CO and usually methane, preferably at least 70% of H2, preferably up to 90% H2, and at least 2 to 4% of CO and 2 to 3% of methane.

[0033] Advantageously, particular embodiments of the invention may split the membrane separation system 80 into several membrane stages:

[0034] A. As a first example, the membrane separation system 80 comprises a first membrane separation unit able to receive its feed gas streams including the gas flow 71 coming from the compressor 70 and to supply a first permeate and a first retentate; a second membrane separation unit able to receive the first retentate and to supply a second permeate and a second retentate; with the second retentate being hydrocarbon-enriched retentate 81 and the mixture of first permeate and second permeate being the hydrogen-enriched permeate 82; B. As a second example, the membrane separation system 80 comprises a first membrane separation unit able to receive its feed gas streams including the gas flow 71 coming from the compressor 70 and to supply a first permeate and a first retentate; a second membrane separation unit able to receive the first retentate and to supply a second permeate and a second retentate; with the second retentate being hydrocarbon-enriched retentate 81, the first permeate being hydrogen-enriched permeate 82 and second permeate recycled to the suction of compressor 70;

[0035] C. As a third example, the membrane separation system 80 comprises a first membrane separation unit able to receive its feed gas streams including the gas flow 71 coming from the compressor 70 and to supply a first permeate and a first retentate; a second membrane separation unit able to receive the first retentate and to supply a second permeate and a second retentate; with the second retentate being hydrocarbon-enriched retentate 81, the second permeate being hydrogen-enriched permeate 82 and first permeate exported as a byproduct 88 to the plant battery limit or recycled to the Hydrogen and Carbon Dioxide Recovery Unit 60; a compressor may be required to correct the pressure of the first permeate to enter the Hydrogen and Carbon Dioxide Recovery Unit 60;

[0036] D. Preferably, the first membrane separation unit and second membrane unit operate at the same temperature and / or permeate pressure.

[0037] E. Preferably, the first membrane separation unit and second membrane unit operate at a different temperature and / or permeate pressure.

[0038] F. Preferably, the first membrane separation unit operates at a lower temperature than the second membrane unit.

[0039] G. Preferably, the first membrane separation unit operates at a higher permeate pressure than the second membrane unit.

[0040] Depending on the embodiment, the hydrogen plant according to the present invention can comprise one or more of the following features (also depicted in Figure 1):

[0041] A. The hydrogen plant includes a feed purification section 10 that removes all impurities that could be present in the hydrocarbon feedstock that could be harmful to the proper operation of the catalysts used in the downstream sections, such as (but not limited to) sulphur species (H2S, mercaptans and / or other sulphur containing species), halides (chlorides and other halogen containing species), metals (mercury, arsenic, and / or others), hydrogenate olefins, and / or other species that would reduce or inhibit the performance of the catalysts.

[0042] B. Optionally the hydrogen plant includes a pre-reformer 20 upstream of reformer 30. Pre-reformer 20 is an adiabatic reforming reactor that does not use any external heat input and that is typically used to convert all higher hydrocarbons into methane, which is preferred for the operation of the downstream reformer 30. In the specific embodiments in which this configuration is chosen, the addition of steam for the reforming reaction may be performed only upstream of the prereformer 20 (stream 44), only upstream of the reformer 30 (stream 45), or may use any combination of upstream pre-reformer 20 (stream 44), upstream reformer 30 (stream 45) and / or upstream WGS reactor 50 (stream 43). Similarly, the recycling of the hydrocarbon-enriched retentate 81 may be done to one or more points in the flow scheme, i.e. a first part 83 entering the feed purification section 10, and / or a second part 84 entering the pre-reformer 20, and / or a third part 85 entering the reformer 30, and / or a fourth part 87 entering the water gas shift reaction section 50.

[0043] In some embodiments the flow rate of the export steam to the battery limit stream 46 may be zero, meaning that the plant is not exporting any steam and all steam generated inside the plant is used in the hydrogen generation process.

[0044] Some embodiments of the present invention may furthermore provide the relevant passageway to enable the use at least or more of the following streams as fuel in fired heater 90: a) A part 64 of the hydrogen and CO2 depleted waste stream 63 from Hydrogen and Carbon Dioxide Recovery unit 60, b) a part of the high-pressure, CO2 depleted stream generated inside the Hydrogen and Carbon Dioxide Recovery unit 60 after removal of CO2 from gas stream 51 fed to the Hydrogen and Carbon Dioxide Recovery Unit, c) A part of the waste stream available at intermediate pressure from the Hydrogen and Carbon Dioxide Recovery unit 60, generated by reducing the pressure on a solvent stream used to capture CO2 from gas stream 51, d) A part of the pure hydrogen product stream 62 generated by the Hydrogen and Carbon Dioxide Recovery Unit 60, e) A part of Hydrogen-rich permeate 82 from the membrane separation system 80. f) A part of the hydrocarbon feed stream 1 to the plant g) A dedicated fuel stream to be used in the fired heater 90

[0045] In these embodiments, Hydrogen and CO2 Recovery Unit 60 comprises one or more units utilizing amine-based solvent wash, Hydrogen-selective membranes, CO2- selective membranes, Pressure Swing Adsorption (PSA) or Vacuum Pressure Swing Adsorption (VPSA), electrochemical compression, sorption enhanced water gas shift, and / or cryogenic fractionation.

[0046] In these embodiments, Hydrogen and CO2 Recovery Unit 60 comprises a CO2 Capture Unit followed by a Hydrogen Purification Unit, with the CO2 capture unit preferably being an amine-based solvent wash system and the Hydrogen Purification Unit preferably a Pressure Swing Adsorption Unit.

[0047] In the generation of stream 82 from the membrane separation system 80, some embodiments of the present invention may provide the necessary passageways in the hydrogen plant to feed the membrane separation system 80 with at least one or more of the following streams: a) A Part 65 or the entirety of the waste stream 63 from the Hydrogen and Carbon Dioxide Recovery Unit 60, after compression by compressor 70 into stream 71 feeding the membrane separation system 80, b) A part of the high-pressure, CO2 depleted stream generated inside the Hydrogen and Carbon Dioxide Recovery unit 60 after removal of CO2 from gas stream 51 fed to the Hydrogen and Carbon Dioxide Recovery Unit, c) A part of the waste stream available at intermediate pressure from the Hydrogen and Carbon Dioxide Recovery unit 60, generated by reducing the pressure on a solvent stream used to capture CO2 from gas stream 51. d) In these embodiments, Hydrogen and CO2 Recovery Unit 60 comprises one or more units utilizing amine-based solvent wash, Hydrogen-selective membranes, CO2-selective membranes, Pressure Swing Adsorption (PSA) or Vacuum Pressure Swing Adsorption (VPSA), electrochemical compression, sorption enhanced water gas shift, and / or cryogenic fractionation, e) In these embodiments, Hydrogen and CO2 Recovery Unit 60 comprises a CO2 Capture Unit followed by a Hydrogen Purification Unit, with the CO2 capture unit preferably being an amine-based solvent wash system and the Hydrogen Purification Unit preferably a Pressure Swing Adsorption Unit. f) In these embodiments, membrane separation system 80 comprises membrane elements based on polysulphone, poly-imid, poly-aramid, cellulose-acetate, or other polymer that exhibits a selectivity for permeating hydrogen to a lower pressure. g) Alternatively, in a specific embodiment, membrane separation system 80 could utilize the selective permeation of hydrogen through a palladium-based membrane operating at elevated temperatures (300-350 °C). Even though the process gas would require additional heating to the membrane operation temperature, the membrane products would be delivered at this same temperature. Having a hot permeate fuel stream is beneficial for the fuel consumption of fired heater 90 as less heat is consumed in heating the fuel to the flame temperature. Similarly, the palladium membrane operating temperature is very similar to the operating temperature of the feed purification, or would reduce the amount of heat to be delivered by the fired heater (90) to pre-heat the cold hydrocarbon feed gas to the reformer 30 inlet temperature. Overall, this means that pre-heating the waste gas to the high membrane operating temperature would not increase the overall hydrocarbon feed consumption required to generate the hydrogen fuel as the heat delivered can be recovered later in the process. A particular advantageous effect of the present invention is the generation of hydrocarbon-depleted, Hydrogen-rich permeate stream 82 that is used as low-carbon fuel to the fired heater 90. While the part 64 of the waste gas stream 63 typically contains approximately 50-90% H2, 1-7% CH4 and 1-5% CO, the hydrogen-rich permeate stream 82 can contain 98.0-99.5% H2 and less than 1 % CH4 + CO. Typically 90% or more of the H2 contained in compressed stream 71 can be recovered to the hydrogen-rich permeate stream 82. Not only does this reduce the CO2 emissions from the stack of the fired heater, it also reduces the concentration of hydrogen and the volumetric flow rate of the hydrocarbon-enriched recycle stream 81 by 60-80 %, thereby saving on electrical compression power on compressor 70 when compressing part 65 of the waste gas stream 63 and furthermore positively contributing to converting feedstock to hydrogen and carbon dioxide that is subsequently captured. This combined effect lowers the hydrocarbon feed consumption and CO2 emissions because less firing will be required, as less H2 is recycled towards the feed of the plant (requiring pre-heating), and less H2 as fuel needs to be generated.

[0048] Furthermore, by optimizing the flow fraction between streams 64 and 65, the plant can be tailored to target a specific CO2 capture rate, as the flow fraction governs the amount of carbon recycled with the fuel to the fired heater. The hydrogen-rich permeate stream 82 is typically available at low pressure (between 0.01 and 0.1 MPa) which is sufficient for use as fuel to the fired heater 90.

[0049] In another embodiment of this invention, the reforming section includes a heatexchanger reformer 35. In such an embodiment, the plant may include the reformer 30 (e.g. an ATR or partial oxidation POX reformer) and a heat-exchanger reformer 35 downstream of the reformer 30. The heat-exchanger reformer 35 may be positioned between the reformer 30 and the heat recovery section 40. A plurality of heat-exchanger reformers 35 may be provided. The heat-exchanger reformer 35 may not be present in alternative embodiments.

[0050] The heat-exchanger reformer 35 can be described as a shell and tube exchanger with the tubes filled with reforming catalyst. The reformed gas stream 32 from the reformer 30 can be mixed into the shell side of the heat-exchange reformer. In such an arrangement, the plant may include a means for mixing the reformed gas stream 32 into the shell side of the heat-exchanger reformer 35, e.g. a passageway or conduit. The plant may be arranged to mix the reformed gas stream 32 into the shell side inside the heatexchanger reformer 35 or outside of the heat-exchanger reformer 35 (e.g. upstream of the shell side inlet). The heat supplied by the reformed gas stream (e.g. hot gas effluent) 32 of reformer 30 into the shell side of the vessel is transferred to the mixture of syngas stream 23 and additional steam 42 flowing through the tubes. When the heat-exchanger reformer 35 is installed as a parallel reformer, shell side effluent and tube side effluent mix at the tube outlet inside the heat exchanger reformer 35 into syngas stream 36 that is the inlet to the heat recovery section 40 and water gas shift reactor 50.

[0051] Figure 2 represents the flow scheme for this embodiment. All variations to the flow scheme applied to Figure 1 without any modification be applied as well to the flow scheme of Figure 2 including the heat-exchanger reformer 35.

[0052] The heat exchanger reformer 35 can also be installed in series, as indicated in Figure 3. In this embodiment,

[0053] A. Hydrocarbon feed 1 and optional hydrocarbon-rich recycle 83 are pre-treated in purification section 10,

[0054] B. Pre-treated hydrocarbon feed 11 and optional hydrocarbon-rich recycle 84 is mixed with process steam 44 and fed to optional pre-reformer 20,

[0055] C. Pre-reformer effluent 21 and optional hydrocarbon-rich recycle 86 is mixed with process steam 42 and fed to the tube side of heat-exchanger reformer 35 producing effluent stream 37,

[0056] D. Effluent stream 37 and optional hydrocarbon-rich recycle 85 is mixed with process steam 45 and fed to reformer 30, utilizing stream 31 containing high purity oxygen 02 to yield reformer effluent 32,

[0057] E. Reformer effluent 32 provides the heat to the shell side of heat exchanger reformer 35 to yield reformate stream 36 that is routed to the heat recovery section 40,

[0058] F. Cooled syngas from the heat recovery section 40 and optional mixing steam 43 are admitted to the water gas shift reaction section 50,

[0059] G. Effluent 51 from the water gas shift reaction section 50 is admitted to the Hydrogen and Carbon Dioxide Recovery Unit 60, that separates the stream into a CO2 rich stream 61, a Hydrogen-rich stream 62 and a hydrogen and C02 depleted waste gas stream 63,

[0060] H. Part 64 of waste stream 63 from the Hydrogen and carbon dioxide recovery unit 60 is used as fuel for fired heater 90, while the remainder 64 or all of the waste stream 63 are compressed by compressor 70,

[0061] I. The compressed waste stream 71 enters the membrane separation system 80, that separates the stream into a hydrogen-rich permeate stream 82 and a hydrocarbon- rich retentate stream 81,

[0062] J. Part or all of hydrogen-rich permeate stream 82 are used as fuel to fired heater 90

[0063] K. Flue gas 92 from fired heater 90 heats hydrocarbon feed and / or water stream 49 to generate steam 41 in the convection section 100 and the heat recovery section 40

[0064] L. The hydrocarbon-rich retentate stream 81 is recycled to one or more of the following : a) A part 83 mixing with hydrocarbon feed 1 to the purification section (if applicable) b) A part 84 mixing with pre-treated hydrocarbons 11 and process steam 44 to the pre-reformer 20 (if applicable) c) A part 86 mixing with hydrocarbon stream 23 and process steam 42 to the heat exchanger reformer 35 d) A part 85 mixing with reformer inlet stream 37 and process steam 45 to the reformer 30 e) A part 87 mixing with the cooled reformate and process steam 43 to the water gas shift reaction 50

[0065] M. The fired heater receiving oxygen required for combustion via stream 91.

[0066] All variations to the flow scheme applied to Figure 1 can without any modification be applied as well to the flow scheme of Figure 3 including the heat-exchanger reformer 35 installed in series.

[0067] The inclusion of a heat-exchanger reformer 35 to the flow scheme brings particular benefits to the heat transfer taking place in the plant. Heat-exchanger reforming, whether installed in parallel or in series, utilizes heat transfer by conduction from the tube wall to the catalyst and flowing hydrocarbon stream instead of by pure convection when the hydrocarbon stream is heated only by a stream of flue gas or process gas. The heat consumed in the heat exchanger reformer is then no longer contained in the reformate effluent 36 that is at much lowertemperature in comparison to reformer effluent 32 when entering heat recovery section 40, which allows a substantial reduction in excess steam generation. The use of high-level heat for the generation of hydrogen instead of steam also results in a reduction in the firing demand and further reduces the fuel demand.

[0068] More in particular, this advantage has a cumulative effect when combined with the hydrogen-permeating membrane separation system 80 in the recycle stream 65. As a result of the hydrogen-permeating membrane, the flow rate of hydrocarbon-enriched recycle stream 81 to the reformer feed section is reduced and thus requiring less feed pre-heating in the fired heater 90. As hydrogen generated within the plant is applied as fuel, any reduction in fuel demand also reduces the required hydrogen generation by the plant, which further reduces firing demand. Due to the low total fuel demand of fired heater 90, the hydrogen-rich permeate stream 81 can provide from 25% to 70% and in some cases even up to 100 % of the total fuel demand. The combined beneficial effect of the heat exchanger reforming with the membrane separation system thus further decreases the hydrogen flowrate required as fuel. The present invention thus allows for a significant decrease of hydrocarbon consumption.

[0069] In a process according to the invention, hydrogen produced from hydrocarbon feed is used as fuel to the fired heater 90 instead of the hydrocarbon feed. The priority fuel is the hydrogen-enriched permeate 82 from the hydrogen-permeating membrane separation system 80 and the required make-up fuel is supplied as part of the product hydrogen 62. When the further reforming reactant comprises steam, carbon dioxide is formed as a side-product in order to obtain said hydrogen and forms part of the reformate 36 (a process gas), rather than being formed by combustion of hydrocarbon in the fired heater 90. Hereby, the generated CO2 can conveniently be captured from the process gas, together with CO2 formed to produce the part of the hydrogen that becomes (part of) the hydrogen-comprising product, withdrawn from the process, rather than being emitted to the atmosphere. As the hydrogen firing increases the feed consumption as more hydrogen is required, a heat exchange reformer reactor is applied to minimize the firing demand. Thus, the design of the reformer and the use of the produced hydrogen act in combination to reduce greenhouse gas emissions and / or global carbon footprint. By the further inclusion of a pre-reformer to the reformer configuration and the use of hydrogen firing with carbon dioxide capture to remove carbon dioxide from process gas (reformate, shift reactor product), the carbon footprint is further reduced. In preferred embodiments, the present invention allows a reduction of up to 99.7 % of the direct emissions of the hydrogen plant. The preferred embodiment in particular provides additional savings of close to 15 % of the power consumption and its associated indirect CO2 emissions and a reduction of up to 15 % of hydrocarbon consumption compared to a state of the art low carbon emission hydrogen plant already achieving a 95 % CO2 reduction of direct emissions. Considering that removal of CO2 becomes progressively more difficult, this is a significant improvement. As illustrated in the examples, the preferred embodiment (case 3) shows a capture rate of almost 99.7 % of the direct CO2 emissions, which is a further improvement of 95.3 % compared to the conventional operation base case (case 1) without the hydrogen membrane in the recycle stream. Additionally, power consumption is decreased as well resulting in a reduction of global CO2 footprint of 6 % to 69 %, depending on the quantity of renewable power in the energy mix. When available electrical power has low carbon intensity, an overall CO2 footprint reduction of up to 69% is achievable.

[0070] It is further noteworthy that the Examples illustrate that the Invention allows a further reduction in carbon dioxide emissions while decreasing the reformer size compared to a state-of-the-art plant design, as exemplified by case 1. In fact, the complete plant size, including the front-end desulfurization 10 and the back-end shift 50 and Hydrogen and Carbon Dioxide Recovery section 60 can be reduced compared to the state-of-the-art plant.

[0071] In a final embodiment of the current invention, the heat recovery section 40 can be optimized to generate steam 41 at a much higher pressure than actually required for the process, while reducing temperature pinches in the heat recovery section 40. The generated high-pressure steam 41 can then be used to generate electrical power in a back-pressure steam turbine 47 expanding the stream 48 to the pressure level required by the process. This embodiment enables a further reduction of the net electrical power required forthe operation of the plant, as sketched in Figure 4. All variations to the flow scheme applied to Figure 1 can without any modification be applied as well to the flow scheme of Figure 4 including the heat-exchanger reformer 35 installed in parallel or in series.

[0072] The invention is described more fully herein with reference to the accompanying figures, in which embodiments of the invention are shown, including some optional elements, e.g. feed purification unit 10 (Figures 1,2, 3,4), pre-reformer 20 (Figures 1,2, 3,4), heat-exchanger reformer 35 (Figures 2 and 3) and steam turbine 47 (Figure 4). Also, locations of units and process lines may deviate from what is schematically shown. E.g. in some embodiments, inside hydrogen and carbon dioxide recovery unit 60, the Carbon dioxide recovery unit is upstream of the hydrogen recovery unit, whereas in other embodiments the order can be the inverse. In specific embodiments of the Hydrogen and Carbon Dioxide Recovery Unit 60, internal streams may be routed to the rest of the plant, e.g. a crude hydrogen stream 68 that is generated after CO2 removal from the shifted gas stream 51 and that is rich in hydrogen but CO2 depleted, and stream 69 that is a flash gas obtained during the regeneration of the amine-based solvent by reducing the pressure, causing dissolved gases to evaporate to the vapor phase. In the drawings, the absolute and relative sizes of systems, components, layers, and regions may be exaggerated for clarity. Embodiments may be described with reference to schematic and / or cross-section illustrations of possibly idealized embodiments and intermediate structures of the invention. In the description and drawings, like numbers refer to like elements throughout. Relative terms as well as derivatives thereof should be construed to refer to the orientation as then described or as shown in the drawing under discussion. These relative terms are for convenience of description and do not require that the system be constructed or operated in a particular orientation unless stated otherwise.

[0073] A person skilled in the art will be able to design and operate suitable operational units of the hydrogen plant or used in a process according to the invention, using the present disclosure in combination with common general knowledge and optionally one or more of the documents cited herein. E.g. a person skilled in the art will be able to provide suitable process / plant units (e.g. reformer units, shift reactor zone units, carbon dioxide recovery units, hydrogen recovery units, heat exchanger units) and passage ways, e.g. pipes, lines, tubes or other channels for passing gases or liquids from one processing unit to another, directly or indirectly, based on the present disclosure, the cited documents and common general knowledge.

[0074] For the purpose of clarity and a concise description, features are described herein as part of the same or separate embodiments, however, it will be appreciated that the scope of the invention may include embodiments having combinations of all or some of the features described.

[0075] Next, processes and plants according to the invention are described in further detail.

[0076] The hydrocarbon feedstock 1 fed into the reformer system can be any hydrocarbon feedstock suitable for being subjected to reforming by reaction with steam. It can in particular be a feedstock wherein the hydrocarbon is a feedstock at least substantially consisting of methane, such as natural gas, a biogas-based methane stream or e-methane; propane gas (LPG), naphtha or refinery off-gas, or streams with comparable compositions originating from the hydroprocessing of renewable feed streams or corresponding e-Fuels.

[0077] Dependent on the purity of the feedstock, the feedstock may be subjected to a pre-treatment in a pre-treatment section 10, such as hydrodesulphurization. Pretreatments, conditions therefore and suitable pre-treatment units, may be based on known technology. In particular, when using a pre-treatment such as hydrodesulphurization, a make-up stream comprising hydrogen is usually added to the feed to ensure purification of the feed in the hydrodesulphurization section. As discussed below, a stream comprising hydrogen produced in a process according to the invention can be used to that purpose, in particular hydrogen product gas, waste gas from a hydrogen and carbon dioxide recovery unit, or hydrocarbon-rich recycle gas from the membrane separation system. The pre-treatment section 10 may be configured to produce a treated hydrocarbon stream 11.

[0078] The hydrocarbon feedstock is mixed with further reformate reactant, i.e. water (steam), carbon dioxide, hydrocarbon-rich recycle gas or a mixture thereof, before being subjected to the reaction in the reformer reaction unit (30 and optionally 35). The reformer 30 converts a stream that includes the hydrocarbon feedstock into a reformed gas stream 32.

[0079] The reformer 30 is an exothermic reformer. The reformer 30 is an autothermal reformer (ATR) or partial oxidation (POX) reformer, e.g. configured for performing autothermal reforming or partial oxidation reforming. The reformer is configured to convert a stream including the hydrocarbon feedstock 1 into a reformed gas stream 32 by conversion with steam rich in oxygen 31. The steam rich in oxygen 31 and the stream including the feedstock 1 may be mixed within the reformer 30 or upstream of the reformer 30 inlet.

[0080] The steam rich in oxygen 31 is a high purity oxygen stream that may contain at least 90% (by volume) oxygen. In some arrangements, the high purity oxygen stream 31 entering the reformer 30 may contain at least 95% (by volume) or at least 98% (by volume) oxygen. The high purity oxygen stream may contain a low amount of steam, e.g. less than 5% by volume, optionally, less than 1% by volume of steam. The stream including the hydrocarbon feedstock 1 may be mixed with an additional steam stream 45 upstream of the reformer 30 or within the reformer 30. The additional steam stream 45 may contain a greater volume of steam than the stream 31, e.g. most of the steam required for the reforming reaction may be provided in stream 45.

[0081] In some embodiments, the reformer 30 may include a reactor vessel that contains a packed bed of catalyst (e.g. a reformer catalyst). The packed bed of catalyst may be provided over a base of the reactor vessel such that fluids in the vessel can pass from the top of the reactor vessel through the packed bed of catalyst and then flow further downwards to an outlet of the vessel, e.g. to a bottom outlet nozzle of the reactor vessel. The reformer 30 may be referred to as a fixed or packed-bed reactor reformer in some embodiments. In contrast, an SMR reformer does not have a packed bed of catalyst installed within a reactor vessel. Instead, an SMR reformer typically includes numerous (e.g. multiple dozens up to multiple hundreds) tubes extending through the reformer, the tubes filled with a catalyst over which the reformer feed passes and inside which the reforming reaction takes place. Advantageously, loading the reformer 30 with a packed bed of catalyst (e.g. on the base of the ATR or POX reformer vessel) is less labour-intensive than loading the catalyst in an SMR reactor, where multiple dozens up to hundreds of tubes are required to be loaded in a similar fashion with catalyst, each tube requiring individual pressure drop measurements to ensure an even flow distribution of the reformer feed gas over the reformer catalyst.

[0082] The reactor vessel of the reformer 30 may be a refractory lined reactor vessel to protect the vessel shell from high temperatures inside the vessel. In some arrangements, the vessel shell may be water-cooled, e.g. by a water-cooling jacket. In some arrangements, the reactor vessel contains a packed bed of catalyst installed on the base of the vessel. The catalyst may be a nickel-based catalyst. The height of the packed bed (e.g. the catalyst) may be less than half of the height of the reactor vessel. The packed bed may be approximately one third of the total reactor height. In some arrangements, the packed bed of catalyst may have a diameter in the range of 2000 to 5500 mm and a height in the range of 2000 to 7000 mm.

[0083] The reformer 30 may contain a burner located in an upper region of the vessel (e.g. above the packed bed). In some arrangements, a plurality of burners are provided. Up to four burners may be provided in some arrangements, e.g. for very large reformer reactors generating several hundred thousand Nm3h1of reformed gas. This contrasts with convection type reformers that generate only around 15,000 Nm3h1of reformed gas and so may have only a single burner or very low number of burners (e.g. less than four). The one or more burners are configured to at least partially combust the hydrocarbon feedstock mixed with steam in the presence of high purity oxygen.

[0084] The upper region of the vessel may define a substantially conical shape. The upper region of the vessel and the one or more burners are arranged so that flames from the burners do not contact the packed bed of catalyst. In some arrangements, a heat-resistant layer may be provided over the uppermost surface of the packed bed of catalyst. The heat-resistant layer may extend over the entire uppermost surface of the packed bed in some arrangements. The heat-resistant layer is configured to protect the catalyst from the high-temperature burner flames. The heat-resistant layer is configured to permit gases from the upper region to enter through the layer into the packed bed below. The reformer 30 may include an outlet nozzle at or toward the base of the vessel through which the reactor effluent (e.g. reformed gas 32) exits the reformer 30. Configuring the reformer 30 to perform ATR or POX reforming is beneficial over conventional SMR reformers since the steam requirements of an ATR or POX reformer are lower than that of an SMR reformer. Specifically, ATR or POX reforming uses steam enriched in oxygen, while SMR requires steam only. Utilising steam enriched in oxygen facilitates the generation of additional steam during the reforming reaction (e.g. via the combustion of hydrocarbons with the oxygen). The steam requirements of a plant having the main reformer configured for SMR are greater than those of a plant having an ATR or POX reformer. In this way, the energy requirements to generate steam in the plant are lower since ATR or POX reforming are implemented in the reformer 30 over SMR.

[0085] In contrast to a traditional endothermic reformer (e.g. performing SMR), the reformer 30 is configured to perform exothermic, oxygen-based reforming (e.g. ATR or POX reforming) to produce a reformed gas stream 32 that includes hydrogen, carbon monoxide, carbon dioxide and at least one hydrocarbon as impurity. In an SMR reformer, the endothermic reaction requires additional heat supply to drive the reforming reaction. In view of this, an SMR reactor is traditionally a fired tubular reformer having a radiant section and a convection section, to which additional fuel is supplied to provide the heat of reaction. In the reformer 30 (e.g. performing ATR or POX reforming), the heat generated by the exothermic reaction is sufficient to drive the process, so no additional heating is required to maintain the reaction (e.g. supplying the heat of reaction). As such, the reformer 30 is not a fired tubular reformer and no additional heat is introduced to the reactor. The process in the reformer 30 is self- sustaining compared with conventional SMR reformers. The energy requirements to provide heat to an ATR or POX reformer (e.g. in the fired heater) are lower than the energy requirements to provide heat to an SMR reformer.

[0086] The plant having an ATR or POX reformer 30 includes a separate fired heater 90 to provide the required stream pre-heating prior to entry into the reformer (e.g. instead of a radiant and convection section within the reformer itself to provide the heat of reaction, as in SMR reforming). The fired heater 90 can advantageously make use of the hydrogen-enriched permeate 82 generated downstream in the process as a low-carbon fuel. In some embodiments, one or more fired heaters 90 may be provided, e.g. in parallel or in series with one another. Providing a plurality of fired heaters 90 may result in improved redundancy in the event one or more of the fired heaters 90 requires maintenance. The one or more fired heaters 90 are provided as separate components to the reformer 30.

[0087] The fired heater 90 (e.g. a small-sized fired heater) is required only for preheating the hydrocarbon feedstock to the reactor inlet temperature, but not for supplying the heat of reaction for the reforming reaction. The hydrocarbon feedstock is heated by the fired heater 90 prior to introduction into the reformer reactor. The feed containing hydrocarbon feedstock may be heated to at least around 600°C by the fired heater 90 prior to introduction into the reformer 30. In some arrangements, the feed containing hydrocarbon feedstock may be mixed with steam (e.g. stream 45) prior to heating the mixture to at least around 600°C by the fired heater 90 upstream of the reformer 30.

[0088] The fired heater 90 is provided as a separate component to the reformer 30, e.g. external to the reformer 30. In embodiments in which a pre-reformer 20 and / or heatexchanger reformer 35 is provided, the fired heater 90 is provided as a separate component to the reformer 30, the pre-reformer 20 and / or the heat-exchanger reformer 35. It will be understood that references to the fired heater 90 being separate to the reformers 20, 30, 35 refers to the fired heater 90 sharing no common structural elements with the reformers 20, 30, 35. In embodiments, the only connection between the fired heater 90 and the reformer 30 is via a tube that directs the pre-heated feed from the fired heater 90 to the reformer 30.

[0089] As noted above, oxidative (e.g. ATR or POX) exothermic reforming provides several process advantages compared to SMR reforming, e.g. reduction in fuel demand, steam demand and hydrocarbon feed consumption. In addition to these, physically separating the fired heater 90 from the reformer 30 provides further advantages over endothermic reforming technologies in which the reformer includes a radiant and convection section. For example, the process steps of feed pre-heating and the reforming reaction can be separated into individual equipment items of smaller size. The layout of the hydrogen plant can therefore be tailored more easily to the area available since the individual equipment items occupy less space than a combined reformer and heater arrangement. Moreover, the orientation of the fired heater 90 relative to the reformer 30 can be adjusted as permitted by available plant space. The provision of a separate fired heater 90 also results in a less complex piping system into and out of the reformer 30, further contributing to a reduction in plant area requirements.

[0090] It will be understood that the fired heater 90 is configured to only heat the feedstock before or upstream of the reformer 30 and not inside the reformer 30, e.g. in contrast to a radiant and convection section of an SMR reformer where heat of reaction is provided within the reformer itself. No reforming reaction occurs within the fired heater 90. The reforming reaction takes place externally to the fired heater 90 in the respective reformer. In contrast to the described arrangement, heat provided in the radiant and convection section of an SMR reformer provides the duty for the SMR reaction.

[0091] The fired heater 90 may have a chamber and one or more tubes extending through the chamber. The chamber may be referred to as a combustion chamber in some arrangements. The chamber of the fired heater 90 is separate from and external to the reformer reactor 30. The fired heater 90 is configured such that a fuel is combustible in the chamber to provide heat to the one or more tubes. The one or more tubes are configured to transfer the stream including the hydrocarbon feedstock 1 (and the steam stream 45) through the chamber of the fired heater 90 and into the reformer 30 such that the stream including the hydrocarbon feedstock 1 is preheated in the fired heater 90 before entering the reformer 30. In some arrangements, the only connection between the fired heater 90 and the reformer 30 is via the tubes. As will be discussed in more detail, a hydrogen-enriched permeate 82 obtained downstream of the reformer may be directed to the fired heater 90 (e.g. to the combustion chamber) to be used as a low-carbon fuel by the fired heater 90. Other additional streams may be used as fuel to the fired heater 90 in some arrangements (e.g. on start-up, fresh fuel is required to be provided prior to the generation of the hydrogen-enriched permeate 82).

[0092] The chamber of the fired heater 90 may contain one or more burners that are fed with a fuel and an oxygen-containing stream 91. The burners may be installed in a base region of the chamber, with hot flue gas exiting from an upper region of the chamber. Such an arrangement allows the burners to be accessible from the ground level, allowing for simple assembly and maintenance. This contrasts with an SMR reforming system, in which for large capacity plants (e.g. exceeding 15 000 Nm3h1of hydrogen production, such as up to or exceeding 200 000 Nm3h1) burners are required to be installed in walls (e.g. side walls) or an upper wall of the SMR reformer to optimize heat release from the flame for the endothermic reforming reaction taking place in tubes extending through the reformer. The SMR reforming system therefore involves a more complex construction, more difficult access to the burners, as well as a more complex piping arrangement compared with the exothermic reformer 30 and separate fired heater 90 of the embodiment of this invention.

[0093] In some arrangements, the burners of the fired heater 90 operate within the combustion chamber at an elevated pressure (e.g. around 3 bar g). In alternative arrangements, the combustion chamber may be at a lower pressure, e.g. the fuel may be supplied at or near atmospheric pressure and the combustion chamber may operate slightly below atmospheric pressure.

[0094] The tubes extend through the chamber (e.g. through the upper region of the chamber). The tubes may be in the form of radiant coils. In some arrangements, the tubes have a processed outer surface to improve heat transfer characteristics (e.g. finned tubes to increase area available for heat transfer). It will be appreciated that no reforming reaction of the hydrocarbon feedstock 1 or steam takes place in the fired heater 90. For example, there is no catalyst present in the tubes of the fired heater 90 and so no reforming reaction occurs therein. The stream passing through the tubes in the fired heater 90 does not undergo a reforming reaction and is only pre-heated prior to exiting the fired heater 90 and being introduced into the reformer for an ATR or POX reforming reaction.

[0095] In some arrangements, the tubes are installed to extend substantially perpendicularly relative to the direction of flow of the flue gas from the burner (e.g. substantially perpendicular to the direction of flame extension of the burners in the chamber). This promotes a cross-current flow of the flue gases over the tubes which improves heat transfer therebetween. In alternative arrangements, the tubes may be installed to extend substantially parallel relative to the direction of flow of the flue gas from the burner to promote a co-current flow of flue gases over the tubes. The plant may include a convection section 100 that is external to the chamber of the fired heater 90. The plant may be configured such that flue gases exit the chamber of the fired heater 90 and enter the external convection section 100. The convection section 100 may be positioned above (e.g. on top of) the chamber of the fired heater 90. Such an arrangement may reduce the footprint plot area occupied by the fired heater, since the convection section is installed on top of the combustion chamber instead of adjacent thereto. Further heat may be recovered from the flue gases in the convection section 100 (e.g. for use elsewhere in the plant). Cooled flue gases may exit the convection section via a stack. The tubes containing the hydrocarbon feedstock (and steam) may extend into the convection section 100 of the fired heater 90 (e.g. in addition to or as an alternative to extending through the combustion chamber of the fired heater 90).

[0096] The steam rich in oxygen 31 may be pre-heated prior to introduction into the reformer 30. The steam rich in oxygen 31 may be pre-heated separately to the hydrocarbon feed and additional steam stream, e.g. in a dedicated heat exchanger. In some arrangements, the oxygen stream 31 may be preheated to approximately 200- 230°C.

[0097] The reformer 30 includes a heat recovery section 40. The heat recovery section 40 is configured to receive the heat from at least part of the reformed gas 32 generated in the reformer 30. A means for sending at least part of the reformed gas 32 to the heat recovery section 40 is provided, e.g. a conduit or a passageway. The heat recovery section 40 may be configured to generate process steam, e.g. for use in the reformer 30.

[0098] Optionally a pre-reformer reaction unit 20 is provided, upstream of the reformer reaction unit 30, and - if present -upstream of the heat-exchanger reformer 35. The prereformer reaction unit 20 may be located immediately upstream of the reformer reaction unit 30 (e.g. the ATR or POX reaction unit). The inlet of the reformer 30 may include the outlet of the pre-reformer 20. It will be understood that the pre-reformer reaction unit 20 is an optional unit that may not be present in some arrangements (i.e. the ATR of POX reformer may be present in isolation, or with the heat-exchanger reformer 35 only). The use of one or more pre-reformer units 20, usually one or more adiabatic prereformer units, to partially perform the reforming reaction (before preheating the prereformed mixture to the inlet temperature of the main reformer with the fired heater 90), is advantageous to unload the duty of the reforming reaction. In the pre-reforming, generally a minor part of the hydrocarbon is converted, whereby -amongst others - CO is formed. The pre-reformer 20 is configured to produce a pre-reformed syngas stream 21 from the treated hydrocarbon stream 11 and / or the hydrocarbon feedstock 1.

[0099] The pre-reformer reaction unit 20 may be a single vessel reactor the contains a reforming catalyst. The vessel diameter may be in the range of 1000 to 4000 mm. The vessel height may be in the range of 1500 to 10000 mm. In some arrangements, a second pre-reformer may be installed in parallel, for example to facilitate catalyst maintenance to one pre-reformer (e.g. catalyst replacement) without shutting down the entire plant.

[0100] The mixture to be fed into the reformer (or pre-reformer(s) in case a prereformer is used) usually at least substantially consists of hydrocarbon and the further reactant.

[0101] Steam and hydrocarbon feed may be fed into the reformer 30 in ratios known in the art. Usually the ratio steam to carbon fed into the reformer reaction unit is at least 1.0 mol / mol, preferably at least 2.0 mol / mol, in particular at least 3.0 mol / mol. Usually the ratio steam to carbon fed into the reformer reaction unit is 5.0 mol / mol or less, preferably 4.0 mol / mol or less, more preferably about 3.0 mol / mol or less. A ratio of steam to carbon of 1.0 to 2.0 mol / mol or lower is generally preferred as this results in the minimized hydrocarbon consumption and CO2 emissions. The steam to carbon ratio may range from 0.6 to 2.5, preferably between 0.8 and 1.2.

[0102] This invention enables the addition of steam at multiple process locations : stream 44 for addition to the pre-reformer 20, stream 45 for addition to the reformer 30, stream 43 for addition to the water gas shift section 50 and stream 42 for addition to the heat exchanger reformer 35. The overall ratio of steam to hydrocarbon can be 3.0 mol / mol and preferably less, for example 2.5, or 2.0, with individual section ratio's less than 1.0 mol / mol forthe pre-reformer and reformer, maintaining ratio's higher than 1.0 mol / mol and preferably higher than 2.0 mol / mol as required by the catalyst for the water gas shift and heat exchanger reformer 35. In accordance with the invention, the required fired duty for pre-heating the feed streams and driving the reforming reactions in the reformer reaction unit is strongly reduced by applying a heat exchanger reforming reactor 35. As already explained, such reduction in fired duty would not be achieved by extra pre-heating of the feed before entering the reformer reaction unit.

[0103] A heat-exchanger reformer unit 35 is optionally provided in a process or plant according to the invention. The heat-exchanger reformer may be based on a heat exchanger known in the art, e.g. from the document WO2018 / 104526, the document W02011 / 077107 or the document WO2012 / 057922. The heat-exchanger reformer is conceived as a shell and tube heat exchanger of which the tubes are filled with reformer catalyst. Fresh or partly converted feed is admitted to the tube side, while higher temperature reformate gas is admitted to the shell side to provide the heat of reaction to the endothermic reforming reaction. The heat-exchanger reformer 35 is configured to produce a reformed stream 37 at the tube side outlet. In case the heat exchanger reformer 35 is installed in series, then the outlets of its tube side and shell side are kept separate and are routed to different destinations in the process flow scheme. In case the heat exchanger reformer 35 is installed in parallel, then the outlets of the tube side and shell side inlet are mixed at the outlet of the heat exchanger reformer, providing a single mixed stream to the downstream process. The outlets of the tube side and shell side inlet may be mixed internally or externally of the heat exchanger reformer. Either configuration (e.g. in parallel or in series) can be applied to the present invention and either configuration takes advantage of the heat exchange taking place between the hot effluent and the colder less converted hydrocarbon feedstock to achieve and increased feedstock conversion with a reduced firing demand and reduced steam production. In some embodiments, the heat-exchanger reformer 35 is provided with the reformer unit 30 and without the pre-reformer 20. In alternative embodiments, the plant may include the pre-reformer 20, the reformer unit 30 and the heat-exchanger reformer 35. It will be appreciated that the heat-exchanger reformer 35 is an optional unit that may not be present in some arrangements (e.g. the autothermal or POX reformer may be present in isolation, or in combination with the pre-reformer 20 only). In some embodiments, the pre-reformer 20 may be provided with the reformer unit 30 and without the heat-exchanger reformer 35. In such an embodiment, all of the pre-reformed syngas stream 21 generated in the pre-reformer 20 may be fed to the reformer reaction unit 30 (e.g. following pre-heating by the fired heater 90).

[0104] In alternative embodiments, the plant may include the pre-reformer 20, the reformer unit 30 and a heat-exchanger reformer 35. In arrangements in which the heatexchanger reformer 35 is installed in series to the reformer 30, the plant is configured so that the pre-reformed syngas stream 21 formed in the pre-reformer 20 is fed to a tube side of the heat exchanger reformer 25. The plant is configured such that the reformed stream 37 generated at the tube side outlet of the heat-exchanger reformer 35 is fed to the reformer 30 (e.g. following pre-heating by the fired heater 90). The plant is configured such that the effluent 32 of the reformer 30 is fed into the shell side of the heat exchanger reformer 35. A reformate 36 exits from the shell side of the heat exchanger reformer 35.

[0105] In arrangements in which the heat-exchanger reformer 35 in installed in parallel to the reformer 30, the plant is configured so that a part 23 of the pre-reformed syngas stream 21 formed in the pre-reformer 20 is fed to the tube side of the heat exchanger reformer 35. The plant is configured such that the other part 22 of the pre-reformed syngas stream 21 is fed to the reformer 30 (e.g. following pre-heating in the fired heater 90). The plant is configured so that the reformer effluent 32 is fed to the shell side of the heat exchanger reformer 35. The reformed stream 37 exiting the tubes is mixed internally or externally to the heat exchanger reformer 35 with the reformer effluent 32 fed to the shell side inlet of the heat exchanger reformer 35. The reformate 36 exits from the shell side outlet of the heat exchanger reformer 35.

[0106] The heat-exchanger reformer is installed in a parallel flow path for the mixture of hydrocarbon feed and steam. In a parallel configuration, a part of the feed to the reformer system is split off 23 and sent to a heat exchanger reforming reactor 35 parallel to the reformer 30. Heat for the reaction in the parallel heat-exchanger reformer 35 is supplied by hot reformer effluent 32, heat available in the flue gas from the radiant section 12 or another high temperature heat source, generally of at least 850 degrees C, in particular of 900 degrees C or higher. An advantage of a parallel heat-exchanger reformer is that, during use, it supplies part of the duty required for hydrogen production, thereby reducing the required duty in the reformer (provided by partial combustion of the hydrocarbon feedstock) and thus the oxygen consumption and its associated electrical power. Generally, the outlet temperature at the catalyst zone (bed) in the heat-exchanger reformer unit is lowerthan in the reformer due to the temperature difference (the driving force) necessary for the heat exchange process to take place. The outlet temperature of the catalyst bed is typically in the range of 850 to about 1000 degrees C, with the proviso that it is lower than the outlet temperature of the reformer 30, usually at least about 30 to 50 degrees C lower. The pressure in the heat exchanger reformer 35 is generally about equal to the pressure in the reformer. Due to the lower outlet temperature, the methane slip is typically higher than in the reformer. To decrease the methane slip, additional steam 42 or carbon dioxide or a mixture thereof is added at the feed of the parallel heat exchanger reactor 35 to drive the reforming reaction towards hydrogen production. Thus, the ratio of further reactant (steam, carbon dioxide, mixture thereof) to hydrocarbon feed, does not have to be the same in different reformer reaction units. The reformer effluent 36 from the heat-exchanger reformer unit 35 is typically combined with the reformer effluent 32 from the reformer unit 30 before further processing, which generally comprises further conversion in a water gas shift section 50, see also below. Because the heat-exchanger reformer is parallel to the reformer 30 and reduces the required duty of the reformer (when comparing at equal hydrogen output), the reformer size can be reduced. When using a parallel configuration, usually about 10 to about 30 wt.% of the hydrocarbon feed, preferably 15 - 25 wt.% is fed to the heat-exchanger reformer reaction unit 35. A higher split ratio will unload the reformer further but reduce the driving force for the heat-exchanger reformer, resulting in an increasingly larger heat exchanger reformer.

[0107] Another configuration would use a process flow scheme in which the heatexchanger reformer 35 would be installed in series with the reformer 30, where all of the feed and steam mixture is sent through a catalyst bed in a multi-tubular heat exchanger reformer 35 to form a partially converted reforming stream. This stream typically exits the heat exchanger reformer below 850°C, more preferably below 750°C before it enters the reforming section 30. The partially converted reforming gas 37 then reacts further in the reforming section 30. The heat for the heat exchanger reformer 35 is supplied by the reforming effluent 32 from the reforming section 30. Similar benefits can be achieved as in the parallel configuration, and a detailed discussion is not included here to avoid repetition.

[0108] The reformate 36 from the heat exchanger reformer 35 is typically cooled in the heat recovery section 40 before further processing. This can be done in a manner known per se. Particularly useful is a waste heat boiler wherein, during use, steam is produced using heat from the reformate stream. Steam 41 is further preferably generated using heat from the flue gases from the fired heater 90. This is generally done in a heat exchanger configured in a heat-exchanging configuration in the convection section 100. The produced steam or part thereof (42,43,44,45) is used as the steam to be reacted with the hydrocarbon feed to be subjected to reaction in the reforming. In principle, produced steam 43 may also be combined with reformate 36 prior to the shift reactor zone 50 or fed to the process gas inside the shift reactor. However, good results are achieved without adding steam to the reformate or inside the shift reactor zone. The process or plant according to the invention can be adapted to be completely self- sufficient in steam production for the process or based with export and / or import steam 15. The lower the required steam production inside the unit battery limits, the more heat integration can be applied and thus lower firing requirement.

[0109] Advantageously, in some embodiments, the steam may be generated at a higher pressure than is required by the reforming process. In such conditions, part or all of the generated steam may be expanded in a steam turbine 47 to recover the energy generated during the expansion as electrical energy and reduce the plant's net electrical power import. In locations where the carbon intensity of electrical power is high, this can be an effective means to lower the direct and indirect CO2 emissions of the plant. In such embodiments, the steam is generated at pressures from 0.1 MPa to 15 MPa above the normal process pressure, preferably more than 8 MPa above the process pressure. The expanded steam 48 at a pressure close to the process pressure is then used in the process. Any export steam 46 remaining after the extraction of the process steam, may be used directly or may be further expanded and condensed, with the condensate returned to the plant for generating fresh steam.

[0110] The reformate, typically after cooling, is usually fed into a water gas shift reactor 50, wherein carbon monoxide reacts with water to form further hydrogen and carbon dioxide. The reformate fed into the water gas shift reactor 50 may be / include the reformed gas stream 32 generated by the reformer 30 (e.g. in embodiments where the heat-exchange reformer 35 is not provided) or may be / include the reformate 36 from the heat-exchanger reformer 35. The reaction in the water gas shift reactor 50 is usually done in the presence of a shift catalyst, which is known per se, e.g. an iron- or copperbased shift catalyst. Thus, the treatment in the shift reactor zone results in a shift reactor product (shift reactor process gas) having an increased hydrogen and carbon dioxide content compared to the reformate. The shift reactor zone design and reaction conditions may in principle be based on known technology, e.g. as described in the prior art cited herein.

[0111] The shift reaction zone is generally operated at a lower temperature than the reformer system. Generally, the temperature during shift reaction is in the range of about 190 to about 500 degrees C. The type of shift applied is generally indicated in three categories based on the outlet temperature of the catalyst. High temperature shift with an inlet of 300-400 degrees C and outlet of 350-500 degrees C; medium temperature shift with an inlet temperature of 190-230 degrees C and an outlet temperature of 280-330 degrees C and low temperature with an inlet temperature of 180-230 degrees C and an outlet temperature of 200-250 degrees C. As the water gas shift reaction is exothermic, the temperature rise is larger for a higher CO concentration at the inlet.

[0112] For a low carbon dioxide emission, it is advantageous to operate the shift reactor zone under conditions wherein the water gas shift reaction (CO + H2O CO2 + H2) is shifted towards the formation of H2 and CO2. As the reaction is exothermic, this is favored by a low temperature. Preferably, a high temperature shift followed by low temperature shift is applied in a process according to the invention to maximize the CO conversion to H2 and CO2 (which CO2 is thereafter captured) and thus minimize the CO2 emissions. Alternatively, only a high temperature shift, a medium temperature shift reaction or an isothermal shift reaction (a cooled shift reactor at a constant temperature) or a combination of any of them can be applied.

[0113] Advantageously, the performance of the water gas shift reaction (CO + H2O CO2 + H2) can be improved by selectively extracting Hydrogen from the reformate (32) or (36) as such extraction would shift the reaction equilibrium to generate more H2 and CO2, thereby resulting in an improved conversion of hydrocarbon feedstock to hydrogen and carbon dioxide. In example arrangements, the plant may include a hydrogen recovery unit (not shown) positioned between the reformer 30 and the water gas shift reactor 50. The hydrogen recovery unit is configured to remove or extract hydrogen from the reformed gas stream 32 before entry into the water gas shift reactor. In embodiments in which a heat-exchanger reformer 35 is present, the hydrogen recovery unit may be positioned between the heat-exchanger reformer 35 and the water gas shift reactor 50 so as to extract or remove hydrogen from the reformate 36 from the heatexchanger reformer 35 before entry into the water gas shift reactor 50. Said extracted hydrogen stream can be used as fuel for fired heater (90). Membrane systems based on selective permeation of hydrogen through palladium sheets at elevated temperature (300-350 °C) would be particularly suitable for this application.

[0114] Advantageously, at least the reformer reaction unit 30 is operated at a relatively high temperature, preferably in the range of 850-1050 degrees C, more preferably in the range of 900-1050 degrees C (at the outlet end of the catalyst) and a deep shift conversion (resulting in a high CO conversion to CO2), typically with an outlet temperature of the shift reactor in the range of 200-250 degrees C is applied in the shift reactor zone to maximize the conversion of feed to hydrogen and thereby reducing the CO2 emissions as well as the hydrocarbon consumption as per invention.

[0115] Usually, in accordance with the invention, the reformate 36 (gas stream coming from the reforming system) or the shift reactor process gas 51 (gas stream coming from the water gas shift reactor) (the latter when - as is usual - a treatment in a shift reactor zone is carried out) is subjected to a carbon dioxide removal treatment in the Hydrogen and Carbon Dioxide Recovery Section 60. Herein, the carbon dioxide content in said reformate 36 or shift reactor process gas 51 is reduced to form a carbon dioxidedepleted process gas (carbon dioxide-depleted product) and a carbon dioxide side- product 61. Capturing the CO2 can be accomplished in a manner known per se, e.g. as described in the cited prior art. Advantageous techniques are temperature swing adsorption (TSA), Vacuum Swing adsorption (VSA), Pressure Swing Adsorption (PSA), Sorption enhanced Water-Gas shift (SEWGS), cryogenic condensation / recovery and amine-based adsorption / stripping process. Typically, amine-based adsorption / stripping process is preferred where all duty for the amine reboiler is supplied by the process gas and thereby minimizing the external heat input, and allocated CO2 emissions of the external heat. Capturing CO2 from the process gas (reformate or shift-reactor product) is preferred as this is the least energy intensive, as the process gas is available at high pressure (typically about 2-3.5 MPa) and therefore has a high driving force for the separation of CO2. The captured CO2 is sent as a high purity stream 61 to the battery limit (outgoing stream from the plant) and could be used for other process, food and beverage industry as well as for storage as in Carbon Capture for Utilization and Storage (CCUS). Further, captured CO2 may be used as a reformer reactant, when applying a dry reforming process. This may be a reformer process different from the present reformer process; however, it is also possible to recycle a gas enriched in carbon dioxide obtained in the carbon dioxide removal treatment, directly or after further purification to higher carbon dioxide content, to the reformer 30 or heat exchanger reformer 35 in a process according to the invention.

[0116] In a specific embodiment in which the CO2 capture technology involves amine- based absorption / stripping, the amine-based solvent regeneration by pressure reduction may generate an additional waste stream (69) at intermediate pressure that can also be integrated in the rest of the plant, either as fuel to the fired heater 90, or as feed to the membrane separation system 80, or as additional feed to the hydrogen purification unit. Similarly, part of the treated gas stream obtained after CO2 removal that is H2 rich and depleted in CO2, may be routed as a stream (68) at full pressure to other sections in the plant, such as the fuel inlet of fired heater (90), or the feed inlet to the membrane separation system (80), to generate additional low-carbon hydrogen fuel.

[0117] Typically, from 70 up to 99% of the CO2 is removed from the process gas, although technically 99.99+% removal of the CO2 is possible; this corresponds to a reduction of the overall CO2 emissions of the hydrogen production plant by about 30- 60%.

[0118] It is noted that CO2 capture on conventional reforming with CO2 removal in the flue gas 101 is also known in the art, but this is very energy intensive and requires complicated technology compared to the present invention, especially when aiming to remove the carbon dioxide in a stream at least substantially. As illustrated in the examples, in accordance with the present invention it is possible to achieve more than 98% reduction in CO2 emissions, by capturing it from a process gas (upstream of a unit wherein hydrogen product gas is obtained or in a tail gas, i.e. the waste gas, from a hydrogen recovery unit) without the requirement of CO2 capture from the flue gas 101. Also from a plant-size perspective it is advantageous not to use / provide a capturing device configured to remove carbon dioxide from flue gas. Still, in an (atypical) embodiment wherein the flue gas in a process according to the invention comprises a significant amount of carbon dioxide, additionally carbon dioxide capture from the flue gas 101 is feasible.

[0119] As will be understood by a person skilled in the art, the reformate, the shift reactor product respectively the carbon dioxide-depleted product still contain substantial amounts of components other than hydrogen. Accordingly, generally hydrogen recovery in the Hydrogen and Carbon Dioxide Recovery Unit 60 is carried out in order to obtain the hydrogen comprising product gas 62 of satisfactory purity, usually at least about 95 mol %, preferably at least 98 % more preferably at least 99 mol. %, in particular at least 99.9 mol %. The purity may be 100 % or less, in particular 99.9999 %or less, 99.999 % or less, 99.99 % or less, 99.9 % or less, 99.5 % or less, 99.0 % or less, dependent on the needs and technology used. Suitable techniques to recover hydrogen from a process gas (such as reformate, shift reactor product respectively the carbondioxide depleted product) can be based on known technology. In a particularly preferred embodiment, a pressure swing adsorption (PSA) unit is provided. PSA allows the production of hydrogen gas 62 that is essentially free of other components. In a further preferred embodiment, a hydrogen-retaining membrane separator is provided. In a further preferred embodiment, an electrochemical compressor is provided. Electrochemical compressors can be used to obtain high purity high pressure hydrogencomprising product gas 62.

[0120] The shift reactor product 51 is subjected to a carbon dioxide removal treatment in hydrogen and carbon dioxide recovery unit 60 to obtain the carbon dioxide-depleted product, of which part is used as fuel in the fired heater 90. The gas enriched in hydrogen 62 is recovered from the carbon dioxide-depleted product using the pressure swing adsorption treatment, the hydrogen-retaining membrane or the electrochemical compressor. In a further preferred embodiment, the carbon dioxide-depleted product is separated into the gas enriched in hydrogen, preferably the hydrogen-comprising product 62, having an increased hydrogen content compared to the carbon dioxide- depleted product and a tail gas 63 (waste gas) having a reduced hydrogen content compared to the carbon dioxide-depleted process stream, said tail gas comprising hydrocarbon and a part 64 of which tail gas product is used as fuel in the fired heater 90, preferably after mixing it with other fuel component or components, such as hydrogen-comprising product. A part 65 of the tail gas (waste gas stream from hydrogen and carbon dioxide recovery unit 60) is compressed to a pressure between 0.5 and 5 MPa, preferably l-3MPa and fed to a membrane separation system 80 separating the compressed tail gas 71 into a hydrogen rich permeate stream 82 at low pressure (i.e. at a pressure comprised between.0.05 and 1 MPa) and a hydrocarbon rich retentate 81, comprising of CH4, CO and some H2 and usually an inert such as nitrogen, at high pressure (i.e. at a pressure comprised between 0.5 and 5MPa). The membrane separation system 80 is selective for the permeation of hydrogen. Hydrogen-rich permeate is used as a fuel to the fired heater 90. In particular for this aspect of the invention it is beneficial to have a first tail gas compression step up to the hydrogenpermeating membrane unit's operating pressure (typically 0.5 to 1.5 MPa), and a second compression step of the hydrocarbon rich retentate stream 81 compressing to the required pressure upstream the reformer. Because the volumetric flow rate of hydrocarbon-rich retentate stream 81 is much reduced in comparison to compressed waste stream 71, this will significantly reduce the total required compression power and total compressor capital cost. In a further preferred embodiment, part of the CO2 depleted product from the CO2 Recovery Unit in the Hydrogen and carbon dioxide recovery unit 60 is fed to the membrane separation system 80, that separates it in a hydrogen rich stream 82 used as fuel in the fired heater and recycles the hydrocarbon rich stream 81. In this embodiment, all of the waste gas 63 is compressed to compressor 70 and membrane separation system 80.

[0121] Recycling the hydrogen-enriched permeate 82 to be used as a fuel in the fired heat 90 advantageously reduces the fuel requirements of the plant. Recycling the hydrogen-enriched permeate 82 is particularly advantageous since the reformer is configured for exothermic ATR or POX reforming, which has less heating (and thus fuel) requirements compared with conventional endothermic SMR reformers. In the present arrangement, the fired heater 90 only requires sufficient fuel to pre-heat the feed into the reformer 30, and so the recycled hydrogen-enriched permeate 82 can make up a greater proportion of the fuel, resulting in lower emissions with the flue gas (compared with using fuels that are rich in hydrocarbons).

[0122] Since the ATR or POX reformer 30 has lower fuel requirements compared with conventional SMR techniques, in some embodiments, not all of the hydrogen-enriched permeate 82 may be required as fuel in the fired heater 90. Part of the hydrogen- enriched permeate 82 may be used as fuel to the fired heater 90. The remaining part of the hydrogen-enriched permeate 82 may be recycled to the inlet of the hydrogen recovery unit of the hydrogen and carbon dioxide recovery unit 60 so as to extract additional, valuable hydrogen from the plant.

[0123] In another preferred embodiment, part of the Hydrogen product 62 obtained in the hydrogen and carbon dioxide recovery unit 60 can also be used to provide fuel for the fired heater 90 in case ultra low CO2 emissions with the flue gas 101 are desired. In this embodiment, all of the waste gas 63 is compressed to compressor 70 and membrane separation system 80, while the hydrogen rich permeate 82 will then be recycled to inlet of the PSA unit in the hydrogen and carbon dioxide recovery unit 60.

[0124] Further, the hydrogen product 62 can be used to provide hydrogen to a hydrocarbon feedstock purification 3, typically hydrodesulphurization. For such recycle, typically 0-5 % of the produced hydrogen product is used, in particular 0.5-3 %, dependent on the quality of the feed and whether another source of hydrogen is used or not.

[0125] Providing a membrane separation system 80 selective for the permeation of hydrogen allows recovery of unconverted carbon monoxide and methane, e.g. in the hydrocarbon-rich retentate stream 81. The hydrocarbon-rich retentate from the hydrogen-permeating membrane separation system may still contain hydrogen gas. It may further contain residual non-converted hydrocarbon from the feedstock, in particular methane, and carbon monoxide formed in the reformer system or shift reaction zone respectively. Accordingly, all or part of the hydrocarbon rich retentate 81 may be recycled from the hydrogen-permeating membrane separation system 80 and combined with one or more of: the hydrocarbon feed 1 upstream of the feed purification 10 (e.g. via a passageway or stream 83), the feed to the pre-reformer 20 (e.g. via passageway or stream 84), the feed to the reformer 30 (e.g. via passageway or stream 85), the feed to the heat exchanger reformer 35 (e.g. via passageway or stream 86) or with reformate 32 / 36 before introduction in the shift reactor zone 50 (e.g. via passageway or stream stream 87). This is schematically shown in Figure 1, 2, 4 but can equally be applied to Figure 3.

[0126] The recycle of the hydrocarbon-enriched retentate at reaction conditions in a reformer unit (20,30, 35) reduces greenhouse gas emissions. When feeding tail gas comprising CO and / or hydrocarbon (in particular methane) to the radiant section for combustion it generates CO2 that is typically emitted into the atmosphere, whereas the CO2 generated from the recycled tail gas can be captured in the carbon dioxide recovery unit and stored or be put to further use. The tail gas is compressed in the compressor 70, since the tail gas pressure is usually considerably lower than the pressure of the stream with which it is to be combined, e.g. more than 10 times lower. E.g. tail gas pressure from a PSA may be about atmospheric or slightly higher, e.g. about 0.13 MPa, compared to a feed pressure in the range of about 2.5 to about 4.0 MPa.

[0127] A minor part of the waste gas stream 63 from hydrogen and carbon dioxide recovery unit 60 (tail gas) may need to be purged, if inert gas (in particular nitrogen) builds up due to recycling. This is accomplished by feeding a part of the tail gas to fired heater 90, where combustible components serve as fuel and the inert gas is purged. Usually less than 10 % of the tail gas needs to be purged (fed to the fired heater 90, via line 64) in this embodiment. The maximum that is advantageously recycled is generally based on when an unacceptable build-up of gases occurs. In the absence of such buildup, essentially all of the tail gas 63 can be advantageously recycled, as the fuel demands of the fired heater can be provided by part of the hydrogen product gas 62 obtained in the hydrogen and carbon dioxide recovery unit 60 as already described.

[0128] The invention of applying a hydrogen-permeating membrane separation system 80 on the tail gas 63 is in particular an advantageous effect for further reduction of carbon dioxide emissions / footprint, but will also reduce the reformer feed requirements, due to the recycled hydrocarbon molecules to the reformer reaction units. Thus, in a process according to the invention, the fraction of the tail gas 63 from the hydrogen and carbon dioxide recovery unit 60 that is recycled can be any fraction from 0 - 100 % ; and an optimum can be chosen dependent on assessing the needed / desired CO2 emission reductions, the needed / desired global CO2 footprint reduction and what is a desired or acceptable size of the reformer system (or plant as a whole). The benefit of combining heat exchanger reforming with tail gas recycle with respect to lowering CO2 emissions becomes bigger with larger recycle flowrate. Thus, for an additional beneficial effect on CO2 emissions, a tail gas recycle of at least about 10 %, preferably at least 25 %, more preferably at least 40 %, in particular at least 50 % can be used. Depending on acceptable requirements of the reformer system, the tail gas recycle can be up to 100%, up to 80 %, up to 60 %, up to 40 %, or up to 20 %.

[0129] The inclusion of the membrane separation system as introduced by this invention avoids the wind-up effect of hydrogen firing on the total fuel demand : additional firing with hydrogen product requires more hydrocarbon feed to generate the hydrogen fuel, which requires more firing and results in an increased tail gas recycle, but the increased firing again requires more feed, etcetera, resulting in size increase of the plant. Instead, the present invention extracts the hydrogen used as fuel from the tail gas recycle and recycles the unconverted carbon molecules back as hydrocarbon feed to the reforming section, which allows a significant net hydrocarbon conversion to hydrogen-comprising product 62 that can be taken from the process by further. The corresponding reduction in net hydrocarbon feed reduces the firing requirements which reduces the amount of fuel hydrogen to be generated by the plant, further reducing the plant size and firing demands. The hydrogen-rich permeate stream 82 reduces the tail gas recycle stream through the reformer unit and thus reduces the required heat input in the reformer section. In comparison to the present art flow scheme with tail gas recycle 63 without membrane separation system, the extraction of hydrogen from the tail gas recycle stream 63 by the membrane separation system 80 reduces the hydrogen partial pressure in the reformer section which increases the conversion of hydrocarbons to hydrogen. The hydrogen rich permeate 82 further provides up to 95%, more typically 80-90%, of the fuel for fired heater 90. This reduces the required crude or pure hydrogen make-up fuel requirements.

[0130] The advantages of the heat exchanger reformer on the feed conversion efficiency and steam generation have already been described above: a significant reduction in firing achieved by the internal heat recovery which further reduces the need for generation of additional hydrogen fuel and its associated waste gas recycle when using hydrogen product as fuel. This further reduces the absolute amount of hydrocarbon feed and tail gas to be compressed in comparison to a technology not utilizing heat exchanger reforming. The heat exchanger reforming thereby substantially reduces the direct and indirect CO2 emissions of the plant.

[0131] Combining the benefits of heat-exchanger reforming and hydrogen fuel extracted from the tail gas recycle by the membrane separation system, a hydrogen plant extracting the generated CO2 from the process gas (reformate and / or shift reactor product) will become a highly efficient low carbon footprint process as their effects are more than cumulative.

[0132] The retentate recycle 81 results still further in lower firing requirement. This combined effect results in an unexpected significant reduction in hydrocarbon feed consumption compared to the same process respectively installation without heat exchanger reformers.

[0133] Thus, in a particularly preferred embodiment of the invention, the combination of the reformer design arrangement with a heat exchanger reformer reaction unit), CO2 removal and tail gas 63 recycle from the in Hydrogen and Carbon Dioxide Recovery Unit 60, with a hydrogen-permeating membrane separation system 80 in the tail gas to extract a major part of hydrogen from the tail gas as hydrogen-rich permeate 82 used as fuel for the fired heater 90, where the hydrocarbon-rich retentate stream 81 is recycled allows in a reduction of CO2 emissions from a reformer plant, in particular from an ATR-based reforming plant, of about 99 % or more, in particular while minimizing the hydrocarbon feed consumption (for generation of fuel); at least in certain embodiments, without the requirement of a purge stream from the tail gas a reduction of more than 99.9% is feasible, making it almost a net zero CO2 emissions plant.

[0134] As follows from the above, hydrogen used as fuel or recycled to be combined with the hydrocarbon feed can be part of the hydrogen-comprising product produced in a process according to the invention, advantageously obtained from the hydrogen and carbon dioxide recovery unit 60. It is however also possible to provide an (additional) hydrogen recovery unit at an upstream location in the hydrogen production plant. Advantageously an (additional) hydrogen recovery unit, is configured to recover a hydrogen enriched gas from the reformate 36 and positioned between the reformate outlet of the reformer system and the feed inlet of the shift reactor zone 50. This has the additional effect of reducing the reformate's hydrogen concentration, which can help to draw the shift reactor towards the production of hydrogen and thus further reduce the hydrocarbon consumption and maximize the CO2 capture ultimately resulting in lower CO2 emissions.

[0135] Such effect may also be achieved by positioning a hydrogen recovery unit in between two shift reactor sections of a shift reactor zone comprising two or more shift reactor units. A preferred embodiment of such hydrogen recovery unit could be a hydrogen-selective membrane utilizing palladium.

[0136] Another option is to provide an (additional) recovery downstream of the shift reactor zone 50, which may be positioned upstream or downstream of a carbon dioxide recovery unit.

[0137] If hydrogen-enriched gas is used as fuel or recycled, it can generally have a lower hydrogen concentration than is desired for the final hydrogen product. Accordingly, less stringent recovery conditions resulting in a lower hydrogen purity can suffice. Particularly suitable is a hydrogen-selective membrane separator. A hydrogen product stream from a methanation unit can also be used as low-carbon fuel. Next, the invention is illustrated by the following Examples.

[0138] Table 1 lists some key performance parameters of the different flow scheme options evaluation for a hydrogen plant capacity based on ATR reforming technology of approximately 9000 kg / h of high purity hydrogen (99.9 % purity). The table utilizes 5 different designs to illustrate the potential benefits of the invention when applied to the ATR-based hydrogen production scheme.

[0139] Numbers in the table expressed in % are always expressed relative to the present state of the art, represented by Case 1. Only the rows for feed efficiency (HHV) and CO2 Capture Rate are an exception to this interpretation, with the feed efficiency defined as the heating value of the hydrogen product relative to the heating value of the hydrocarbon feed (excluding power intake), and the CO2 capture rate defined as the fraction of carbon in the CO2 product exported from the plant of the total carbon entering the plant with the hydrocarbon feed and / or fuel streams. The compression of the CO2 product to 3 MPa is a substantial contribution to the total power consumption of the plant and is listed as a separate contributor to the total for each case as the contribution could be higher when compressing at higher outlet pressure. It can also be used as a relative measure for the total CO2 capture and production of each scheme. Applying the concepts introduced by this invention reduces the electrical power required for the CO2 compression, even though more CO2 is captured from the feed, indicating that the total CO2 production from the plant is also reduced. The reference defining 100 % for all columns is case 1. In case of power consumption, they all refer to 100% power consumption of the state-of-the-art plant represented by case 1, e.g., CO2 compression for Case 4 consumes 20.6 % of the total power consumption of Case 1.

[0140] Carbon Intensity for Direct Emissions (scope 1) and Indirect Emissions from Electrical Power Generation (Scope 2) are expressed in kg CO2 per kg H2 produced. A typical value for a standard hydrogen plant without CO2 capture is in the order of 10 kg C02 / kg H2, and target values for state of the art plants range between 0.5 and 1.5 kg CO2 / kg H2, depending on the carbon intensity of the electrical power used. The table shows the scope 1 + scope 2 values for 3 different locations : location A has a high contribution of renewable energy (50 g CO2eq / kWh, e.g. Norway, Sweden), location B has only medium contribution of renewable energy (250 g CO2eq / kWh, e.g Australia,

[0141] Poland, United Kingdom) and location C has low contribution of renewable energy (380 g CO2eq / KWh, Germany, United States). A higher oxygen consumption typically results in a higher power consumption, while the remainder of the power consumers remain of similar order of magnitude. Renewable energy should be interpreted here as a power source with low associated CO2 emissions, such as solar PV, wind turbines, hydro and / or nuclear power generation.

[0142] Table 1 - Performance evaluation

[0143] Case 1 and Case 2 describe the present state of the art units optimized for 95 % Direct Emissions CO2 Capture, while Case 3, Case 4 and Case 5 describe flow schemes according to the present invention achieving 99.7 %, 98.2 % and 99.7% direct CO2 emissions capture, respectively. To achieve 99.7% capture rate with a present state of the art plant significantly increases utility consumption and capital expenditure. On the contrary, the application of the present invention reduces the hydrocarbon feed consumption, oxygen consumption and overall plant electrical power consumption providing significant benefits to the operation.

[0144] The Carbon Intensity values for location C demonstrate that the application of the present invention still results in a reduction of the total Direct and Indirect CO2 Emissions, but the impact is much less in comparison to location A and B where the carbon intensity of electrical power is much lower. The oxygen generation required for the ATR process is very energy intensive and is hardly compensated by the reduction of the CO2 which is not favourable in location C as known to those skilled in the art. In locations where electrical power is cheap and / or has low carbon intensity, the contribution of the power to the total carbon intensity is (much) lower and then the hydrocarbon conversion efficiency and CO2 capture rate from the process will be the governing parameters to obtain an overall low carbon intensity. This is visible in the results of Case 3, Case 4 and Case 5. Case 1 is a hydrogen plant according to an ATR-based reforming flow scheme, containing a feed purification section 10, pre-reformer 20 and ATR reformer 30, heat recovery section 40 with convection section 100, water gas shift section 50, Hydrogen and Carbon Dioxide Recovery system 60, and a fired heater 90, in which all the waste gas 63 is used as fuel 64 to the fired heater 90 and waste gas recycle 65 nor membrane separation system 80 are applied. Because there is excess heat available in the process, excess steam needs to be generated that is exported for power generation in a steam turbine 47. This reduces the net power import required for the plant but also increases the feed consumption and thereby quantity of CO2 generated during the conversion. Because the unconverted carbon from the waste gas is nor captured nor recycled, the CO2 stack emissions in the flue gas 101 are high, resulting in just over 95 % CO2 capture from the feed hydrocarbons. The captured CO2 is compressed to 30 bar g and exported to the plant battery limits for further treatment.

[0145] Case 2 describes the same plant, with the addition of a heat-exchanger reformer 35 arranged in parallel (Figure 2). Waste gas recycle 65 nor membrane separation system 80 are included. The excess steam is exported for power generation, but as the flow rate is reduced, the power generated inside the turbine is less and the net imported power therefore increases accordingly. The total oxygen consumption significantly reduces which reduces the total power consumed by the plant.

[0146] Case 3 (according to this invention) applies the hydrocarbon-rich recycle 81 from the membrane separation system 80 with the hydrogen-rich permeate stream 82 used as fuel to the fired heater 90 as described in this invention. Also, the heat exchanger reformer 35 is applied. Also, crude hydrogen 68 is fed to membrane separation system 80 to further decarbonize the fuel used for the fired heater 80. A steam turbine 47 is still included for power recovery from generated process steam. In this case, the generation of steam is at higher pressure than required for the process and the power recovery happens in the expansion till process pressure. The CO2 capture rate from feed increases to 99.7 % while the hydrocarbon feed 1 and oxygen 31 consumption further reduce, as does the total power consumption of the plant. The application of the hydrogen-rich stream 81 from the membrane separation system 80 as low-carbon fuel to the fired heater significantly reduces the CO2 emissions from the plant stack. The recycle of the carbon molecules contained in the hydrocarbon-rich stream 81 extracted from the membrane separation system 80 recycled as feed hydrocarbons to the reforming section via streams (84, 85, 86) further reduces the quantity of hydrocarbon feed 1 required for the same total hydrogen production, thereby improving the plant energy efficiency and further reducing the CO2 generated from the hydrocarbon feed. The corresponding reduction in power consumption further reduces the carbon intensity of the plant. In locations with high fossil fuel contribution to the electrical power production (380 g CO2 eq / kWh), this flow scheme still has the lowest carbon intensity (direct and indirect emissions from power generation) of all schemes studied with a value of just above 1 kg CO2 emitted / kg H2 produced. In regions where the renewable energy contribution to the power mix is high (50 g CO2eq / kWh), this scheme enables reaching a carbon intensity of 0.17 kg CO2 emitted per kg of H2 produced. Case 5 is identical to this scheme, with the exception of the power generation block, which may be a cost-driven decision. It can be seen that in Case 5, with respect to Case 3, there is a slight reduction on hydrocarbon consumption, but the carbon intensity is higher in the cases in which the carbon footprint associated to electricity is high (B and C).

[0147] Case 4 (according to this invention) is a variation of Case 3 in which some hydrocarbon feed 1 is used directly in the fired heater 90, replacing the heating value provided by sending some of the crude hydrogen stream 68. This increases the carbon content of the stream 92 as demonstrated by the corresponding reduction in CO2 capture rate from feed to 98.2 % and the related increase in carbon intensity, but also brings a small benefit in hydrocarbon feed consumption that further reduces by 1.3 % and a 2.8 % further reduction in power consumption. This variation to the flow scheme demonstrates that the person skilled in the art has the flexibility of designing the plant for a specific CO2 capture rate or carbon intensity requirement without compromising the plant's energy efficiency through overdesign.

[0148] Case 5 (according to this invention) is another variation of Case 3 in which the steam turbine (47) is not present. Because there is no internal power generation in this flow scheme, the net imported electrical power consumption is higher than for Case 3. In locations with high carbon intensity of electrical power, this will negatively impact the scope 1 + scope 2 carbon intensity of the plant, but in locations where electrical power has low carbon intensity, the same overall carbon intensities as for a plant with internal power generation can be achieved. The comparison of Case 3 and Case 5 performance figures further illustrates that steam generation consumes additional hydrocarbons and oxygen to provide the hydrogen and heat required for the steam generation, but still reducing the overall CO2 emissions and carbon intensity in regions with high carbon intensity for electrical power, indicating that the reduction of carbon intensity achievable by the plants designed according to this invention can be further improved by the addition of own power generation to reduce net power import in locations where electrical power has low renewable contribution.

[0149] Figure Legend

[0150] I Hydrocarbon feed supply

[0151] 10 Feed purification section

[0152] II Purified hydrocarbon feed

[0153] 20 Pre-reformer

[0154] 21 Pre-reformed feed

[0155] 22 Pre-reformed feed to reformer

[0156] 23 Pre-reformer feed to heat exchanger reformer

[0157] 30 Reformer

[0158] 31 Oxygen-rich gas to reformer

[0159] 32 Reformer effluent or reformed gas

[0160] 35 Heat exchanger reformer

[0161] 36 Heat exchanger reformer effluent or reformate stream

[0162] 37 Reformer feed from heat exchanger reformer

[0163] 40 Heat Recovery Section

[0164] 41 Steam

[0165] 42 Process steam to heat exchanger reformer

[0166] 43 Process steam to water gas shift

[0167] 44 Process steam to pre-reformer Process steam to reformer

[0168] Export steam

[0169] Steam turbine

[0170] Expanded steam

[0171] Water

[0172] Water Gas Shift Reactor section

[0173] Hydrogen and Carbon Dioxide Recovery Unit

[0174] CO2 rich product

[0175] H2 rich product

[0176] H2 and CO2 depleted waste stream

[0177] Waste stream to fuel

[0178] Waste stream to compressor

[0179] Crude Hydrogen from CO2 Capture

[0180] Flash gas from CO2 capture solvent regeneration

[0181] Waste gas compressor

[0182] Compressed waste gas stream

[0183] Hydrogen Rich Off-Gas from Battery Limit

[0184] Membrane separation system

[0185] Hydrocarbon-rich retentate

[0186] Hydrogen-rich permeate

[0187] Hydrocarbon-rich retentate recycle to feed

[0188] Hydrocarbon-rich retentate recycle to pre-reformer

[0189] Hydrocarbon-rich retentate recycle to reformer

[0190] Hydrocarbon-rich retentate recycle to heat exchanger reformer

[0191] Hydrocarbon-rich retentate recycle to water gas shift reactor

[0192] Hydrogen-rich permeate export to battery limit

[0193] Fired heater

[0194] Combustion oxygen

[0195] Flue gas

[0196] Convection section

[0197] Flue gas to stack

Claims

CLAIMS1. A Hydrogen production plant comprising :At least one reformer (30) for converting a stream comprising a hydrocarbon feedstock (1) through conversion with steam rich in oxygen 02 into a reformed gas stream (32) comprising hydrogen, carbon monoxide, carbon dioxide and at least one hydrocarbon as impurity, said reformer (30) comprising exothermic, oxygen-based autothermal (ATR) or partial oxidation (POX) reforming and a heat recovery section (40),A fired heater (90) configured to preheat the stream comprising the hydrocarbon feedstock (1) before entry into the at least one reformer (30),At least one water gas shift (WGS) reactor (50) for converting the carbon monoxide of the reformed gas stream (32) into a shifted gas stream (51) containing additional carbon dioxide and hydrogen,A Hydrogen and Carbon Dioxide Recovery Unit (60) located downstream of the WGS reactor (50) and configured to remove carbon dioxide and hydrogen from the shifted gas stream (51), and to produce a first product stream (61) enriched in carbon dioxide and a second product stream (62) enriched in hydrogen, a waste stream (63) depleted in both hydrogen and carbon dioxide,A compressor (70) for compressing a part (65) of the waste gas stream (63) from hydrogen and carbon dioxide recovery unit (60) into a compressed gas stream (71), A membrane separation system (80) selective for the permeation of hydrogen configured to be fed with the compressed gas stream (71) and to produce a hydrogen-enriched permeate (82) stream and a hydrocarbon-enriched retentate (81) stream,A passageway for feeding at least part of the hydrogen-enriched permeate (82) to the fired heater (90) to be used as a low-carbon fuel by the fired heater (90), and A passageway for recycling the hydrocarbon-enriched retentate (81) to the hydrocarbon feed (1) via a pipeline (83) and / or to the reformer (30) via the pipeline (85) and / or to the inlet to the water gas shift reactor (50) via the pipeline (87) .

2. Hydrogen production plant according to claim 1, wherein the plant comprises a passageway for feeding the fired heater (90) with a part (64) of the waste gas stream (63) from hydrogen and carbon dioxide recovery unit (60).

3. Hydrogen production plant according to claim 1 or claim 2, wherein the Hydrogen and Carbon Dioxide Recovery Unit (60) is configured to produce a flash gas stream (69), and / or a stream (68) depleted in carbon dioxide and rich in hydrogen.

4. Hydrogen production plant according to claim 3, wherein the fired heater (90) receives its fuel from at least one of the following streams: a) At least a part (64) of the waste gas stream (63) from the hydrogen and carbon dioxide recovery unit (60); b) At least a part of the hydrogen-enriched permeate (82) produced by the hydrogen-permeating membrane separation system (80); c) At least a part of the hydrogen-enriched product (62) from the hydrogen and carbon dioxide recovery unit (60); d) At least a part of the hydrocarbon feedstock stream (1); e) A make-up fuel stream imported from a battery limit; f) At least a part of the flash gas stream (69) and g) At least a part of the stream depleted in carbon dioxide and rich in hydrogen (68).

5. Hydrogen production plant according to one of claims 1 to 4, wherein the heat recovery section (40) is configured to generate a steam stream (41) and the plant comprises means for routing at least a part of this steam stream (41): a) As process steam (45) to the inlet of the reformer (30) and / or b) As process steam (43) to the inlet of the water gas shift reactor (50) c) As export steam (46) to a battery limit.

6. Hydrogen production plant according to one of claims 1 to 5, wherein the hydrogen plant comprises a feed purification section (10) upstream of the reformer (30), configured to produce a treated hydrocarbon stream (11).

7. Hydrogen production plant according to claim 6, wherein the heat recovery section (40) is configured to generate a steam stream (41) and the hydrogen plant comprises:At least one pre-reformer reactor (20) upstream of the reformer (30), configured to produce a pre-reformed syngas stream (21) from at least one or both of the treated hydrocarbon stream (11) and a part of the hydrocarbon-rich stream (81) via means (84)Means for routing at least a part (44) of steam stream (41) to the inlet of the prereformer (20).

8. Hydrogen plant according to any of the claims 6 or 7, wherein the heat recovery section (40) is configured to generate a steam stream (41) and the plant comprises: a heat exchanger reformer (35) installed in series to the reformer (30) to receive at least one of the hydrocarbon feed (11), the pre-reformer feed (23) or part of the hydrocarbon-rich stream (81) provided via means (86), mixed with steam (42) to the tube side inlet in order to produce a reformed stream (37) from the tube side outlet; means for feeding this reformed stream (37) to the reformer (30) to produce a reformed stream (32), with the reformed stream (32) configured to enter the shell side inlet of heat exchanger reformer (35) and provide the heat of reaction to the tube side of the heat exchanger reformer (35) and to produce a reformate stream (36) from the shell side outlet of the heat exchanger reformer (35), means for sending the reformate stream (36) to the heat recovery section (40). means for mixing at least a part (42) of steam stream (41) to the tube inlet of the heat exchanger reformer (35).

9. Hydrogen production plant according to claim 6 or 7 , wherein: the heat recovery section (40) is configured to generate a steam stream (41)The plant comprises means for dividing the reformer feed stream (21) into a first part (23) and a second part (22),The reformer (30) is configured to receive the second part (22) and to produce the reformed gas stream (32),The plant comprises a heat-exchanger reformer (35) installed in parallel to the reformer (30) and configured to receive at least one or both of the part (23) of the reformer feed stream (21) and at least part of the hydrocarbon-rich stream (81) through means (86), as well as the reformed gas stream (32), to produce a reformate stream (36), and means for mixing the reformed feed gas stream (32) into the shell side of the heat exchanger-reformer (35) with the outlet of the heat exchanger reformed gas from the tube side of the heat exchanger-reformer (35), inside or outside of the heat exchanger-reformer (35), and means for mixing at least a part (42) of steam stream (41) to the tube inlet of the heat exchanger reformer (35), and the heat recovery section (40) is configured to receive reformate stream (36).

10. Hydrogen production plant according to any of the claims 8 or 9, comprising multiple heat exchanger reformers (35).

11. Hydrogen production plant according to any of the claims 1 to 10, wherein the heat recovery section (40) is configured to produce a steam stream (41) and the plant comprises a steam turbine (47) configured to receive at least a part of the steam stream (41) and to produce a lower pressure stream (48).

12. Hydrogen production plant according to any of the claims 1 to 11, wherein the membrane separation system (80) comprises membrane elements based polysulfone, poly-imid, poly-aramid, cellulose acetate, any combination thereof, or other polymeric material, or Palladium sheets, exhibiting a selectivity to preferentially permeate hydrogen to a lower pressure.

13. Hydrogen production plant according to any of the claims 1 to 12, wherein the plant comprises a passageway for importing hydrogen and hydrocarbon containing off-gas (72) from a battery limit as feed to the membrane separation system (80), recovering hydrogen from said off-gas to the hydrogen-enriched permeate (82) and hydrocarbons to the hydrocarbon-enriched retentate (81).

14. Hydrogen production plant according to any of the claims 1 to 13, wherein the plant comprises a passageway to route at least part of the hydrogen-rich permeate stream (82) as fuel to the fired heater (90) and means for recycling the remaining part of the hydrogen-rich permeate stream (82) to the inlet of the hydrogen recovery unit inside the Hydrogen and Carbon Dioxide Recovery Unit (60).

15. Hydrogen production plant according to any of the claims 1 to 14, in which at least a part (88) of the hydrogen-rich permeate stream (82) from the membrane separation system (80) is sent as a product stream to the plant battery limits.

16. Hydrogen production plant according to any preceding claim, further comprising a hydrogen recovery unit positioned between the reformer (30) and the water gas shift reactor (50), the hydrogen recovery unit configured to remove hydrogen from the reformed gas stream (32) before entry into the water gas shift reactor (50).

17. Hydrogen production plant according to any preceding claim, wherein the fired heater (90) is provided as a separate component to the at least one reformer (30).

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