Plant and method for producing hydrogen from hydrocarbons with reduced CO2 emissions

Abstract

-At least one reformer (30) for converting a flow containing hydrocarbon feedstock (1) into a reformed gas flow (32) containing hydrogen, carbon monoxide, carbon dioxide and at least one hydrocarbon as an impurity via conversion with oxygen (O2)-rich vapor, wherein the reformer (30) includes an exothermic oxygen-based autothermal reforming (ATR) or partial oxidation (POX) reforming and a heat recovery section (40), and -At least one reformer ( Before entering 30), a heating furnace (90) configured to preheat a flow containing hydrocarbon feedstock (1), at least one water-gas shift (WGS) reactor (50) for converting carbon monoxide in the reformed gas flow (32) into a shifted gas flow (51) containing additional carbon dioxide and hydrogen, and a first product flow (61) located downstream of the WGS reactor (50) and which removes carbon dioxide and hydrogen from the shifted gas flow (51) and enriches carbon dioxide and hydrogen - A hydrogen and carbon dioxide recovery unit (60) configured to generate a second generation flow (62), a waste flow (63) in which both hydrogen and carbon dioxide have been depleted, - a compressor (70) for compressing a portion (65) of the waste gas flow (63) from the hydrogen and carbon dioxide recovery unit (60) into a compressed gas flow (71), and - a membrane separation system selective for hydrogen permeation, to which the compressed gas flow (71) is supplied and which generates a hydrogen-enriched permeate (82) flow and a hydrocarbon-enriched concentrate (81) flow. A hydrogen production plant comprising: a stem (80); a passage for supplying at least a portion of the hydrogen-enriched permeate (82) to the furnace (90) so that it can be used as a low-carbon fuel by the furnace (90); and a passage for recycling the hydrocarbon-enriched concentrate (81) to a hydrocarbon feed (1) via a pipeline (83), and / or to a reformer (30) via a pipeline (85), and / or to the inlet of a water-gas shift reactor (50) via a pipeline (87).

Description

Technical Field

[0001] The present invention relates to a hydrogen plant and a process for producing a hydrogen-containing product gas from a hydrocarbon feedstock with reduced CO2 emissions for implementing this plant.

[0002] In particular, the present invention relates to a plant and a process for producing hydrogen from a hydrocarbon feedstock, which hydrocarbon feedstock is subjected to reforming using a reformer that uses autothermal reforming (ATR) with optional pre-reforming for generating synthesis gas, which is subjected to water gas shift conversion to increase the conversion of the hydrocarbon feedstock to hydrogen and carbon dioxide, which is recovered in a hydrogen and carbon dioxide recovery section, generating two product streams (one of which is hydrogen-rich and one of which is carbon dioxide-rich), as well as a waste stream depleted in hydrogen and carbon dioxide, and all or at least a part of this waste stream is compressed into a membrane separation system selective for hydrogen, and as a result, the hydrogen-rich permeate product stream from the membrane separation system is used as a low-carbon fuel for a heating furnace, while the hydrocarbon-enriched concentrate stream from the membrane separation system is recycled at least partially to the pre-reformer feedstock, and / or partially to the reformer feedstock, and / or partially to the water gas shift section.

Background Art

[0003] Plants for producing hydrogen from hydrocarbons, particularly steam methane reforming (SMR) plants, are widely applied in refining complexes for supplying hydrogen for improving the quality of several products, for example, in hydrocracking, hydrogenation or hydrodesulfurization. Additionally, hydrogen is used as a component of synthesis gas (a mixture containing hydrogen and carbon monoxide). Synthesis gas is an essential component for generating, for example, ammonia, methanol, synthetic fuels, and many different chemicals.

[0004] Automated thermal reforming (ATR) is widely applied to synthesis gas production and, combined with integrated CO2 capture, is being developed as an alternative to SMR-based processes for hydrogen production. Furthermore, there is growing interest in using hydrogen as a substitute for petroleum-based fuels in the energy sector as a means of providing seasonal energy storage, in industry to provide high-quality heat, and primarily in mobility for heavy and long-distance transportation. It is estimated that approximately 95% of the world's hydrogen supply is produced from fossil fuels. As a byproduct of this technology, CO2 is produced and released into the atmosphere. CO2 is not only produced by the combustion of carbon-based fuels to heat the feedstock to the temperature required for reforming, but CO2 is also formed as a byproduct in hydrogen production. The steam reforming reaction produces carbon monoxide (using methane as the starting compound: CH4 + H2O ⇔ CO + 3H2), which is then converted to carbon dioxide via a water-gas shift reaction (CO + H2O ⇔ CO2 + H2). The steam reforming reaction is highly endothermic and requires a considerable additional heat input. Exothermic reforming processes, such as autothermal reforming (ATR) or partial oxidation reforming, are carried out partially by the hydrocarbon feedstock (using methane as the combustion starting component: 2 CH4 + O2 ⇔ 2CO + 4H2). As a result, a small heating furnace is required only to preheat the hydrocarbon feedstock to the reactor inlet temperature, and not to supply the heat of reaction for the reforming reaction. The exothermic reforming process thereby consumes oxygen O2, an additional feedstock to the plant. The formed CO is then converted to CO2 by a water-gas shift reaction.

[0005] In recent years, there has been increasing industrial attention to reducing environmental emissions and carbon footprints, and there is a growing demand for the design of hydrogen production facilities (also known as hydrogen production units: HPUs) that can reduce carbon footprints.

[0006] Traditionally, hydrogen plants were integrated with refining or industrial facilities that could utilize the excess steam generated. This was because hydrogen plants were considered efficient steam generators, and the excess steam was optimized to meet external steam demand in order to recover as much lower heat as possible, making the excess steam a valuable by-product of the facility. As a result, combustion, and therefore CO2 emissions, were not major design parameters, but this is changing with the development of regulations on greenhouse gas emissions. Many different solutions have been proposed to reduce steam production in hydrogen plants by reducing combustion demand. Options include preheating the combustion air with flue gas or another indirect heat source (typically up to 600°C), preheating the fuel and / or tail gas, and applying an adiabatic (pre-reforming) step. All of these solutions reduce the required combustion demand, thereby also reducing the discharge steam flow rate and CO2 emissions. However, the majority of CO2 released through the flue gas comes from methane, as well as tail gas containing further residual hydrogen, carbon monoxide, and carbon dioxide produced as part of the reforming reaction. This tail gas is typically generated as a waste stream of final product purification carried out by adsorption-based processes (e.g., Pressure Swing Adsorption, PSA). In recent years, with a focus on CO2 emissions, a commonly applied solution is to capture CO2 from the synthesis gas downstream of the (final) shift reactor. The hydrogen product thus obtained is generally called blue hydrogen. The overall reduction of CO2 emissions is limited by the slip of CH4 and CO from the process-side reformer system and shift section, as well as the amount of additional fossil fuel supplied as replenishment fuel to meet the thermal duty requirements of the reforming process. Methane slip can be reduced by operating at a high self-heat-exchange reformer outlet temperature that requires more oxygen and combustion.

[0007] Therefore, there is a continuing need to provide alternative processes and equipment to reduce greenhouse gas (carbon dioxide, methane) emissions, for example, to enable the modification of existing and new HPUs. In particular, there is a continuing need to provide an efficient method for producing hydrogen from hydrocarbon feedstocks by a reforming process that further reduces greenhouse gas (carbon dioxide and / or methane) emissions and / or further reduces the global carbon footprint. In particular, the primary objective is to reduce hydrocarbon consumption (i.e., the use of hydrocarbons for purposes other than producing hydrogen product gases from hydrocarbons). Direct carbon dioxide emissions from the process are further reduced by making it possible to reuse unconverted carbon in the process instead of using it as fuel. Oxygen-based reforming processes require the addition of an oxygen-rich flow. The processes used to generate this flow may be pressure swing adsorption, vacuum pressure swing adsorption, membrane separation, or cryogenic fractionation of air, all of which require compressing ambient air to high pressure before the actual separation step. Oxygen can also be produced by the electrolysis of water. Therefore, all processes available for producing high-purity oxygen have high power consumption, typically ranging from 0.38 to 0.75 kWh / kg O2, depending on purity, pressure, capacity, and the selected technology. Thus, a secondary objective of oxygen-based reforming processes in particular is to reduce power consumption and the associated CO2 emissions, or, if low-carbon intensity renewable electricity is already readily available, to utilize electricity to reduce the total hydrocarbon consumption of the reforming process. This invention addresses both objectives.

[0008] Since oxygen-based reforming processes provide most of the reaction heat required for catalytic conversion, the heat input provided by the tail gas from the final product purification is often found to be too high in terms of calorific value. When a CO2 removal system is included, the tail gas stream generally has a high hydrogen content (50–75 mol%) and concentrates unconverted carbon molecules in the form of methane CH4 and carbon monoxide CO as balance impurities, along with inert gases (nitrogen, argon, and / or helium) and water vapor from the feed stream. When CO2 removal is not applied, the tail gas stream contains 20–50 mol% CO2, which in turn reduces the concentration of other components without affecting the calorific value of the stream. To avoid excessive combustion not required by the process when the stream is used as fuel in a heating furnace, an alternative use for this stream needs to be found. Literature WO2022 / 038089 describes a plant in which the tail gas is therefore recycled, at least in part, to the inlet of an ATR reformer, an optional pre-reforming, and / or a water-gas shift section. Reference WO2021 / 073834 describes a plant in which tail gas is subjected to an additional hydrogen recovery step to reduce the calorific value of the tail gas to the furnace. This can be considered a more general version of reference WO2020 / 221642, which describes a plant in which tail gas is compressed and sent to a membrane separation section, the hydrogen-rich product flow from the membrane is recycled to a hydrogen purification unit, and the hydrocarbon-enriched product flow from the membrane is used as fuel to the furnace. Reference EP3988502 describes a plant based on SMR reforming in which a portion of the waste gas is used as fuel in the furnace, a portion of the waste gas is recycled as feed to a reformer, and the excess waste gas is discharged as a fuel flow used outside the battery limits of the hydrogen plant.

[0009] Recycling high concentrations of reactants to a reaction section that produces these reactants as products always negatively impacts reaction equilibrium. Since the reforming and water-gas shift reactions are highly equilibrium-controlled, recycling a waste gas stream with a high H2 content to the reforming and / or water-gas shift reaction section negatively impacts reaction equilibrium, although the recycling of carbon molecules benefits overall feed consumption. Therefore, the efficiency of the processes described in references WO2020 / 221642 and WO2021 / 073834 benefits from an additional hydrogen recovery step on the waste gas stream, as hydrogen reactants are removed from the stream recycled to the reaction section.

[0010] Excess heat from process gases and flue gases is recovered in hydrogen plants by steam generation. Reforming and water-gas shift reactions consume steam as a reactant. Since steam generation requires heat provided by the conversion of hydrocarbon feedstocks, reducing steam consumption, such as that required by reforming and water-gas shift reactions, increases process efficiency. Typically, the steam required for reforming and water-gas shift reactions is added to the hydrocarbon feedstock at the plant's feed inlet, which can result in more steam being delivered through the upstream sections of the plant than is actually needed. Reference WO2020 / 221642 describes a plant in which steam addition to the hydrocarbon feedstock is divided section by section, with steam added to the reforming and water-gas shift sections as needed by each individual section. This reduces steam consumption by the process and also reduces the heat available to generate steam. Therefore, an optimized plant aims to generate only the steam needed for the process, without generating discharge steam that is not consumed by the process, and to inject steam into the process where it is needed. An optimized steam generation system utilizes available heat with maximum efficiency and without temperature pinch. Further optimizations, which may optionally be included in certain embodiments of the present invention, can be achieved by recovering the energy stored in the steam as pressure by increasing the pressure in the steam system to maximize heat uptake from process and flue gases, and expanding the steam on the turbine to the pressure required by the process, particularly when power is high-cost or has associated carbon intensity. This approach reduces energy losses in the heat recovery section and reduces the net power taken up by the process.

[0011] For plants where the end user has no destination for the excess steam produced, balanced steam production can be targeted (i.e., no steam is delivered, and only steam consumed within the HPU is produced). Steam production can be reduced through the process modifications already described (combustion air preheating, fuel preheating, pre-reforming), but these steps typically do not allow for a reduction to zero delivered steam. To achieve zero delivered steam, the heat normally used to generate delivered steam is used for additional reforming instead of igniting an additional duty cycle or burning additional feedstock. The present invention achieves this goal by including a heat exchanger reformer, as described in document WO2018 / 104526. This heat exchanger reformer can be installed in parallel, as described in document WO2018 / 104526, or in series with the main reformer, as described in document WO2011 / 077107 or document WO2012 / 057922. A heat exchanger reformer utilizes the heat of the high-temperature process gas flowing through the shell of a shell-and-tube heat exchanger reactor as a heat source for the lower-temperature hydrocarbon feedstock flowing over the 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 to generate steam until the amount of steam produced by the process matches the process steam demand, thus achieving a zero-discharge steam plant.

[0012] Steam delivery and feed conversion efficiency are only two of the key performance indicators for modern hydrogen production units. Other key performance indicators include carbon dioxide emissions from the hydrogen production unit's stack (range 1), associated carbon dioxide emissions generated during the generation of electricity consumed by the unit (range 2), and associated carbon dioxide emissions from residual carbon remaining in the hydrogen products or carbon dioxide emissions throughout the lifecycle of the hydrogen production unit's feedstock (range 3). Typically, carbon intensity is used as a parameter representing the total equivalent CO2 emissions per kilogram of hydrogen produced by the plant. Depending on the accuracy and availability of data for calculating range 2 and range 3 emissions from the plant, carbon intensity can be expressed for emissions in range 1, range 1+2, range 1+3, or range 1+2+3. To reduce range 1 emissions, more CO2 should be captured by the plant's CO2 capture unit, and methane and carbon monoxide slip into the fuel for the furnace should be minimized. To reduce range 2 emissions, lower power consumption is preferable. To reduce emissions in range 3, it is preferable that the residual content of carbon molecules (methane, carbon monoxide, and carbon dioxide) in the hydrogen product be low. Using a hydrogen purification step based on pressure swing adsorption (PSA) is required to achieve this goal. The PSA process generates an exhaust gas stream containing residual carbon molecules, as described above. To reduce emissions in range 1, the process should utilize only hydrogen from the waste stream as fuel, while recycling carbon molecules as feed for the reformer. The processes described in references WO2021 / 073834 and WO2020 / 221642 utilize the hydrocarbon-enriched stream, which is depleted of hydrogen obtained from the exhaust gas of a PSA unit, as fuel for a heating furnace, thereby reducing overall feed consumption by utilizing unconverted carbon from the hydrocarbon feed as fuel instead of recycling it as feed. Accordingly, the present invention extracts hydrogen from the exhaust gas of a PSA unit and uses it as a low-carbon fuel, reducing CO2 emissions in range 1 from combustion, while improving feed conversion by returning carbon molecules to the reformer for recycling.

[0013] The oxygen (O2) required for the internal partial combustion reaction in oxygen-based reforming processes must be supplied externally, and typically requires high-purity oxygen to avoid contamination of the resulting synthesis gas with excess nitrogen. The production of high-purity oxygen by air separation is the most commonly used process for this purpose. While various techniques exist for this process (pressure swing adsorption, vacuum swing adsorption, membrane, cryogenic fractionation), they all require compression of ambient air as the oxygen source, which makes the air separation process highly energy-intensive. While the reduction in fuel consumption is beneficial for direct CO2 emissions from the furnace stack (so-called range 1 emissions), the substantial increase in power consumption for oxygen generation is accompanied by a potentially greater increase in CO2 emissions associated with power generation (so-called range 2 emissions), depending on the power source. For this reason, total CO2 emissions from oxygen-based reforming processes are supported by a high contribution from electricity generated from renewable resources.

[0014] In regions where the carbon intensity of the generated electricity is relatively high, any reduction in net intake power consumption benefits the overall carbon intensity of the process. Plant optimization for generating electricity within the plant by generating higher-pressure steam and recovering the power within the steam turbine while expanding the steam to process pressure, as described by the present invention, may be beneficial under these conditions.

[0015] Another recent approach to reducing CO2 emissions, or more specifically, carbon intensity (the total amount of CO2 produced directly or indirectly by the process per unit of hydrogen produced in the reforming process), is to use non-fossil or e-hydrocarbons as feedstock for the process. Non-fossil sources for hydrocarbon feedstock can be obtained from waste hydrocarbon-containing streams generated by the hydrogenation of vegetable oils and / or other long-chain hydrocarbon-containing organic molecular sources. e-hydrocarbons are hydrocarbon molecules produced by processes that combine captured CO2 with H2 generated from renewable electricity and are used as energy carriers that are converted back to H2 when the storage and / or transport of H2 is considered impractical where H2 is generated. Since the carbon in hydrocarbon molecules used as feedstock for the reforming process is of non-fossil or recapture origin, it can be considered neutral in terms of CO2 emissions when converted to CO2 by the reforming process. Capturing the CO2 generated by this process even allows for a negative carbon intensity for the reforming process. [Overview of the project]

[0016] The present invention aims to increase reforming efficiency, reduce CO2 emissions from the stack, reduce fuel consumption, and reduce power consumption by making several improvements to existing techniques for hydrogen production using oxygen-based reforming, particularly autothermal reforming (ATR). The main objective of the present invention is to add a hydrogen selective membrane separation system to the waste flow of a hydrogen and carbon dioxide recovery unit, in particular to use the hydrogen-rich permeate flow as a low-carbon fuel, while recovering the carbon-rich concentrated flow as feed to the reformer. In this text, the reformer refers to a reforming reactor that utilizes exothermic oxygen-based reforming, preferably autothermal reforming (ATR) or partial oxidation (POX). This makes it possible to reduce CO2 stack emissions by removing carbon from the fuel via the membrane separation system which is selective for hydrogen, and to use the resulting hydrogen-rich permeate flow as a low-carbon fuel. Alternatively, the reforming efficiency is improved by recycling the carbon used as fuel in the heating furnace (as feed to the reformer). The reduction in additional fuel consumption and output steam generation is achieved by the use of heat exchanger reforming, in addition to self-thermal reforming with an optional auxiliary reformer. The heat recovery and additional hydrocarbon conversion achieved by the heat exchanger reformer reduce fuel requirements and increase reforming efficiency. This further reduces the oxygen consumption of the ATR, thereby further reducing the power requirements associated with oxygen generation. Optimizing the heat recovery system to reduce the temperature pinch while generating process steam at higher pressures enables power recovery by steam expansion and reduces net imported power consumption. The object of the present invention is to address one or more of these needs. One or more alternative or additional objectives that may be addressed are derived from the following description.

[0017] The object of the present invention is a hydrogen generation plant, and the hydrogen generation plant is - At least one reformer (30) for converting hydrocarbon feedstock 1 into a reformed gas stream 32 containing hydrogen, carbon monoxide, carbon dioxide, and at least one hydrocarbon as an impurity via conversion with oxygen-rich vapor, wherein the reformer 30 includes an oxygen-based reforming and a heat recovery section 40, - At least one water-gas shift (WGS) reactor 50 for converting the carbon monoxide in 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 is located downstream of the WGS reactor 50 and is configured to remove carbon dioxide and hydrogen from the shifted gas flow 51, and to generate a first production flow 61 enriched with carbon dioxide, a second production flow 62 enriched with hydrogen, and a waste flow 63 in which both hydrogen and carbon dioxide are depleted. -A compressor 70 for compressing a second portion 65 of the waste gas flow 63 from the hydrogen and carbon dioxide recovery unit 60 into a compressed gas flow 71, -A membrane separation system 80 selective for hydrogen permeation, configured to supply a compressed gas flow 71 and generate a hydrogen-enriched permeate flow 82 and a hydrocarbon-enriched concentrate flow 81, - A passage for supplying at least a portion of the hydrogen-enriched permeate 82 to the heating furnace 90, -The system includes a passage for recycling the hydrocarbon enriched concentrate 81 to the hydrocarbon feed 1 via pipeline 83, and / or to the reformer 30 via pipeline 85, and / or to the inlet of the water-gas shift reactor 50 via pipeline 87.

[0018] Depending on the embodiment, the hydrogen plant according to the present invention may include one or more of the following features. - The reformer 30 includes oxygen-based thermal reforming (ATR) or partial oxidation (POX) reforming, which generates heat. -A heating furnace 90 is provided, configured to preheat the flow containing hydrocarbon feedstock 1 before it enters at least one reformer 30. - At least one heater 90 is provided that is configured to preheat the stream containing the hydrocarbon feedstock 1 before it enters at least one reformer 30. - The passageway is arranged to supply at least a portion of the hydrogen-enriched permeate 82 to the heater 90 for use as a low-carbon fuel by the heater 90. - The plant comprises a passageway for supplying a first portion 64 of the waste gas stream 63 from the hydrogen and carbon dioxide recovery unit 60 to the heater 90. - The hydrogen and carbon dioxide recovery unit 60 is configured to produce a flash gas stream 69 and / or a stream 68 that is depleted in carbon dioxide and rich in hydrogen. - The heater 90 receives its fuel from a) at least a portion 64 of the waste gas stream 63 from the hydrogen and carbon dioxide recovery unit 60, b) at least a portion of the hydrogen-enriched permeate 82 produced by the hydrogen permeable membrane separation system 80, c) at least a portion of the hydrogen-enriched product 62 from the hydrogen and carbon dioxide recovery unit 60, d) at least a portion of the hydrocarbon feedstock stream 1, e) a make-up fuel stream taken from battery limits, f) at least a portion of the flash gas stream 69, and g) at least a portion of an optional stream 68, from at least one of the streams. - The heat recovery section 40 is configured to generate a steam stream 41, and the plant comprises means for routing at least a portion 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 send-out steam 46 to battery limits. - The hydrogen plant comprises a feed purification section 10 that is upstream of the reformer 30 and is configured to produce a treated hydrocarbon stream 11. - The heat recovery section 40 is configured to generate a vapor flow 41, and the hydrogen plant comprises at least one pre-reforming reactor 20 located upstream of the reformer 30 and configured to produce a pre-reformed synthesis gas flow 21 from at least one or both of a portion of the treated hydrocarbon flow 11 and a hydrocarbon-rich flow 81 via means 84, and means for routing at least a portion 44 of the vapor flow 41 to the inlet of the pre-reforming reactor 20. Preferably, the hydrogen plant comprises means for recycling a portion 44 of the vapor flow 41 generated by the heat recovery section 40, and / or means for recycling a portion 44 of the vapor flow 41 generated by the heat recovery section 40. - The heat recovery section 40 is configured to generate a steam flow 41, and the plant comprises a heat exchanger reformer 35 installed in series with the reformer 30 and receiving at least one of a hydrocarbon feed 11, a pre-reformer feed 23, or a portion of a hydrocarbon-rich flow 81 provided via means 86, which is mixed with steam 42 to the tube-side inlet to generate a reformed flow 37 from the tube-side outlet; means for supplying this reformed flow 37 to the reformer 30 to generate a reformed flow 32, wherein the reformed flow 32 is configured to enter the shell-side inlet of the heat exchanger reformer 35, provide reaction heat to the tube side of the heat exchanger reformer 35, and generate a reformed flow 36 from the shell-side outlet of the heat exchanger reformer 35; means for sending the reformed flow 36 to the heat recovery section 40; and means for mixing at least a portion 42 of the steam flow 41 to the tube inlet of the heat exchanger reformer 35. - The heat recovery section 40 is configured to generate a vapor stream 41. The plant comprises means for splitting the reformer feed stream 21 into a first portion 23 and a second portion 22. The reformer 30 is configured to receive the second portion 22 and generate a reformed gas stream 32. The plant comprises a heat exchanger reformer 35 installed in parallel with the reformer 30 and configured to receive at least one or both of at least a part of the first portion 23 of the reformer feed stream 21 and at least a part of the hydrocarbon-rich stream 81 via the means 86 and the reformed gas stream 32 to generate a reformed stream 36, 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 the heat exchanger reformer 35, and means for mixing at least a part 42 of the vapor stream 41 into the tube inlet of the heat exchanger reformer 35. The heat recovery section 40 is configured to receive the reformed stream 36. - In one embodiment, there are provided means for mixing the reformed feed gas stream 32 to 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 the heat exchanger-reformer 35, and means for mixing at least a part 42 of the vapor stream 41 into the tube inlet of the heat exchanger-reformer 35. The heat recovery section 40 is configured to receive the reformed stream 36. - The hydrogen production plant comprises a plurality of heat exchanger reformers 35. - The heat recovery section 40 is configured to generate a vapor stream 41. The plant comprises a steam turbine 47 configured to receive at least a part of the vapor stream 41 and generate a low-pressure stream 48. Preferably, the hydrogen plant comprises a passage for recycling at least a part of the low-pressure stream 48 to the heat exchange reactor 35, and / or a passage 43 for recycling at least a part of the low-pressure stream 48 to the water gas shift reactor 50, and / or a passage 44 for recycling at least a part of the low-pressure stream 48 to the pre-reformer 20, and / or a passage 45 for recycling at least a part of the low-pressure stream 48 to the reformer 30, and / or a passage 46 for sending at least a part of the low-pressure stream 48 outside the plant. - The membrane separation system 80 comprises membrane elements based on polysulfone, polyimide, polyaramid, cellulose acetate, any combination thereof, or other polymer materials, or palladium sheets, which exhibit selectivity to preferentially permeate hydrogen to lower pressures. - The hydrogen plant is equipped with a passage for introducing off-gas 72 containing hydrogen and hydrocarbons from the battery limit to the membrane separation system 80 as a feed, and recovering hydrogen from the off-gas into a hydrogen-enriched permeate 82 and hydrocarbons into a hydrocarbon-enriched concentrate 81. - The hydrogen plant includes a passage for routing at least a portion of the hydrogen-rich permeable logistics 81 to a heating furnace 90 as fuel, and means for recycling the remaining portion of the hydrogen-rich permeable logistics 81 to the inlet of a hydrogen recovery unit in a hydrogen and carbon dioxide recovery unit 60. - At least a portion 88 of the hydrogen-rich flow 82 from the membrane separation system 80 is sent to the plant battery limit as a product flow. - 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 being configured to remove hydrogen from the reformed gas stream 32 before it enters the water-gas shift reactor 50. - The heating furnace 90 is provided as a component separate from at least one of the reformers 30. - The reformer may include a reaction vessel containing a packed bed of catalyst (e.g., reformer catalyst). - The reaction vessel may include a catalyst-packed bed located at the bottom of the vessel. - The hydrocarbon supply flow is a) Fossil-based hydrocarbon sources such as natural gas, liquefied petroleum gas, naphtha, or any combination thereof, b) Hydrogen-rich off-gases from processes that treat fossil-based hydrocarbon flows, c) Biogenic flows, for example, liquid or vapor product flows generated by the processing of vegetable oils, cooking oils and other similar sources, and whose hydrocarbon content is similar to that of liquefied petroleum gas, naphtha, or any combination thereof, d) The hydrogen-rich off-gas generated by the process intended in c) of this claim, e) Off-gases generated by the fermentation process, containing hydrogen, methane, and / or carbon monoxide, f) Methane-rich gas obtained from biological sources or landfill gas, g) obtained by at least one of the following: e-hydrocarbons obtained through the synthesis of hydrogen obtained from renewable or other sources such as e-methane with captured CO2, and their derivatives, or by-products obtained from the synthesis of their derivatives by Fischer-Tropsch synthesis or similar reactions. [Brief explanation of the drawing]

[0019] [Figure 1] The flow scheme of the present invention is shown. [Figure 2] This includes the element shown in Figure 1, which adds a heat exchanger reformer in parallel. [Figure 3] The system includes the elements shown in Figure 1, which add a heat exchanger reformer in series. [Figure 4] The system includes the elements shown in Figure 1, which incorporates a steam turbine for power generation. [Modes for carrying out the invention]

[0020] Advantageously, the exhaust gas stream 63 contains H2, CO, and usually methane, preferably at least 70% H2, preferably up to 90% H2, and at least 2-4% CO and 2-3% methane.

[0021] Advantageously, in certain embodiments of the present invention, the membrane separation system 80 may be divided into several membrane stages. A. In a first embodiment, the membrane separation system 80 comprises a first membrane separation unit that receives its supply gas flow, including a gas flow 71 coming from a compressor 70, and can supply a first permeate and a first concentrate; and a second membrane separation unit that receives the first concentrate and can supply a second permeate and a second concentrate, wherein the second concentrate is a hydrocarbon-enriched concentrate 81, and the mixture of the first permeate and the second permeate is a hydrogen-enriched permeate 82. B. In a second embodiment, the membrane separation system 80 comprises a first membrane separation unit that receives its supply gas flow, including a gas flow 71 coming from a compressor 70, and can supply a first permeate and a first concentrate, and a second membrane separation unit that receives the first concentrate and can supply a second permeate and a second concentrate, wherein the second concentrate is a hydrocarbon-enriched concentrate 81, the first permeate is a hydrogen-enriched permeate 82, and the second permeate is recycled to the suction of the compressor 70. C. In a third embodiment, the membrane separation system 80 comprises a first membrane separation unit that receives its supply gas flow including a gas flow 71 coming from a compressor 70 and can supply a first permeate and a first concentrate, and a second membrane separation unit that receives the first concentrate and can supply a second permeate and a second concentrate, wherein the second concentrate is a hydrocarbon-enriched concentrate 81, the second permeate is a hydrogen-enriched permeate 82, and the first permeate is either sent to the plant battery limit as a byproduct 88 or recycled to a hydrogen and carbon dioxide recovery unit 60, and a compressor may be required to correct the pressure of the first permeate entering the hydrogen and carbon dioxide recovery unit 60. D. Preferably, the first membrane separation unit and the second membrane unit operate at the same temperature and / or permeate pressure. E. Preferably, the first membrane separation unit and the second membrane unit operate at different temperatures and / or permeation pressures. F. Preferably, the first membrane separation unit operates at a lower temperature than the second membrane unit. G. Preferably, the first membrane separation unit operates at a higher permeate pressure than the second membrane unit.

[0022] Depending on the embodiment, the hydrogen plant according to the present invention may include one or more of the following features (also shown in Figure 1): A. The hydrogen plant includes a feed purification section 10 that removes all impurities that may be present in the hydrocarbon feedstock that could be detrimental to the proper operation of the catalyst used in the downstream section, such as (but not limited to) sulfur species (H2S, mercaptans and / or other sulfur-containing species), halides (chlorides and other halogen-containing species), metals (mercury, arsenic and / or others), hydrogenated olefins, and / or other species that degrade or impair the performance of the catalyst. B. Optionally, the hydrogen plant includes a pre-reformer 20 upstream of the reformer 30. The pre-reformer 20 is an adiabatic reforming reactor used to convert all higher hydrocarbons to methane without using an external heat input, which is preferable for the operation of the downstream reformer 30. In the particular embodiment in which this configuration is selected, the addition of steam for the reforming reaction may be carried out only upstream of the pre-reformer 20 (flow 44), only upstream of the reformer 30 (flow 45), or any combination of the upstream pre-reformer 20 (flow 44), the upstream reformer 30 (flow 45), and / or the upstream WGS reactor 50 (flow 43). Similarly, the recycling of the hydrocarbon enriched concentrate 81 may be carried out at one or more points in the flow scheme, namely, a first portion 83 entering the feed purification section 10, and / or a second portion 84 entering the pre-reformer 20, and / or a third portion 85 entering the reformer 30, and / or a fourth portion 87 entering the water-gas shift reaction section 50.

[0023] In some embodiments, the flow rate of steam delivered to the battery limit flow 46 may be zero, meaning that the plant is not delivering steam and all steam generated within the plant is used in the hydrogen generation process.

[0024] Some embodiments of the present invention may further provide associated passages that enable the use of at least or more of the following flows as fuel in the heating furnace 90. a) A portion of the waste stream 63 64 from which hydrogen and CO2 from the hydrogen and carbon dioxide capture unit 60 have been depleted. b) A portion of the high-pressure CO2 depletion flow generated within the hydrogen and carbon dioxide recovery unit 60 after removing CO2 from the gas flow 51 supplied to the hydrogen and carbon dioxide recovery unit. c) A portion of the waste stream available at intermediate pressure from the hydrogen and carbon dioxide capture unit 60, generated by reducing the pressure of the solvent stream used to capture CO2 from the gas stream 51. d) A portion of the pure hydrogen product stream 62 generated by the hydrogen and carbon dioxide recovery unit 60. e) A portion of the hydrogen-rich permeate 82 from the membrane separation system 80. f) Part of the hydrocarbon supply flow 1 to the plant. g) A dedicated fuel stream used in the heating furnace 90.

[0025] In these embodiments, the hydrogen and CO2 recovery unit 60 includes one or more units utilizing amine-based solvent washing, hydrogen selective membrane, CO2 selective membrane, pressure swing adsorption (PSA) or vacuum pressure swing adsorption (VPSA), electrochemical compression, adsorption-enhanced water-gas shift, and / or cryogenic fractionation.

[0026] In these embodiments, the hydrogen and CO2 recovery unit 60 comprises a CO2 capture unit and a subsequent hydrogen purification unit, wherein the CO2 capture unit is preferably an amine-based solvent washing system, and the hydrogen purification unit is preferably a pressure swing adsorption unit.

[0027] In generating the flow 82 from the membrane separation system 80, some embodiments of the present invention can provide the necessary passages within the hydrogen plant to supply at least one of the following flows to the membrane separation system 80. a) Part 65 or all of the waste stream 63 from the hydrogen and carbon dioxide recovery unit 60 after it has been compressed by the compressor 70 into a stream 71 supplied to the membrane separation system 80. b) A portion of the high-pressure CO2 depletion flow generated within the hydrogen and carbon dioxide recovery unit 60 after removing CO2 from the gas flow 51 supplied to the hydrogen and carbon dioxide recovery unit. c) A portion of the waste stream available at intermediate pressure from the hydrogen and carbon dioxide capture unit 60, generated by reducing the pressure of the solvent stream used to capture CO2 from the gas stream 51. d) In these embodiments, the hydrogen and CO2 recovery unit 60 includes one or more units utilizing amine-based solvent washing, hydrogen selective membrane, CO2 selective membrane, pressure swing adsorption (PSA) or vacuum pressure swing adsorption (VPSA), electrochemical compression, adsorption-enhanced water-gas shift, and / or cryogenic fractionation. e) In these embodiments, the hydrogen and CO2 recovery unit 60 comprises a CO2 capture unit and a subsequent hydrogen purification unit, wherein the CO2 capture unit is preferably an amine-based solvent washing system and the hydrogen purification unit is preferably a pressure swing adsorption unit. f) In these embodiments, the membrane separation system 80 comprises membrane elements based on polysulfone, polyimide, polyaramid, cellulose acetate, or other polymers that exhibit selectivity for permeating hydrogen to lower pressures. g) Alternatively, in certain embodiments, the membrane separation system 80 can utilize the selective permeation of hydrogen through a palladium-based membrane operating at high temperatures (300-350°C). The process gas needs to be further heated to the membrane's operating temperature, but the membrane product is supplied at this same temperature. Having a high-temperature permeate fuel flow is beneficial to the fuel consumption of the furnace 90 because less heat is consumed when heating the fuel to the flame temperature. Similarly, the operating temperature of the palladium membrane is very similar to the operating temperature of feed refinement or reduces the amount of heat delivered by the furnace (90) to preheat the low-temperature hydrocarbon feed gas relative to the inlet temperature of the reformer 30. Overall, this means that preheating the waste gas to a high membrane operating temperature does not increase the overall hydrocarbon feed consumption required to produce hydrogen fuel, since the supplied heat can be recovered later in the process.

[0028] A particularly advantageous effect of the present invention is the generation of hydrocarbon-depleted hydrogen-rich permeate flow 82, which is used as a low-carbon fuel for the heating furnace 90. While a portion of the waste gas flow 63 64 typically contains about 50-90% H2, 1-7% CH4, and 1-5% CO, the hydrogen-rich permeate flow 82 can contain 98.0-99.5% H2 and less than 1% CH4 + CO. Typically, more than 90% of the H2 contained in the compressed flow 71 can be recovered into the hydrogen-rich permeate flow 82. This not only reduces CO2 emissions from the heating furnace stack but also reduces the hydrogen concentration and volumetric flow rate of the hydrocarbon-enriched recycled flow 81 by 60-80%, thereby saving electric compression power for the compressor 70 when compressing a portion of the waste gas flow 63 65, and further contributes positively to converting the feedstock into hydrogen and subsequently captured carbon dioxide. This combined effect reduces hydrocarbon feed consumption and CO2 emissions because less H2 is recycled towards the plant's feedstock (requiring preheating), and less H2 needs to be generated as fuel, thus requiring less combustion.

[0029] Furthermore, by optimizing the flow ratio between flows 64 and 65, the plant can be adjusted to target a specific CO2 capture rate, as the flow ratio governs the amount of carbon recycled to the furnace along with the fuel. The hydrogen-rich permeate flow 82 is typically available at a low pressure (0.01–0.1 MPa) sufficient for use as fuel to the furnace 90.

[0030] In another embodiment of the present invention, the reforming section includes a heat exchanger reformer 35. In such an embodiment, the plant may include a reformer 30 (e.g., an ATR or partially oxidized 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. Multiple heat exchanger reformers 35 may be provided. The heat exchanger reformer 35 may not be present in alternative embodiments.

[0031] The heat exchanger reformer 35 can be described as a shell-and-tube exchanger in which the tubes are filled with a reforming catalyst. The reformed gas flow 32 from the reformer 30 can be mixed with the shell side of the heat exchanger reformer. In such a configuration, the plant may include means, e.g., passages or conduits, for mixing the reformed gas flow 32 with the shell side of the heat exchanger reformer 35. The plant may be configured to mix the reformed gas flow 32 with the shell side inside the heat exchanger reformer 35 or with the outside of the heat exchanger reformer 35 (e.g., upstream of the shell side inlet). The heat supplied to the shell side of the vessel by the reformed gas flow (e.g., high-temperature gas effluent) 32 from the reformer 30 is transferred to a mixture of synthesis gas flow 23 and additional steam 42 flowing through the tubes. When the heat exchanger reformer 35 is installed as a parallel reformer, the shell-side effluent and tube-side effluent are mixed at the tube outlet inside the heat exchanger reformer 35 with the synthesis gas flow 36, which is the inlet to the heat recovery section 40 and the water-gas shift reactor 50.

[0032] Figure 2 shows the flow scheme of this embodiment. All changes to the flow scheme applied to Figure 1 are also applied to the flow scheme of Figure 2, including the heat exchanger reformer 35, without any modifications.

[0033] The heat exchanger reformer 35 can also be installed in series, as shown in Figure 3. In this embodiment, A. Hydrocarbon feed 1 and optional hydrocarbon-rich recycled 83 are pretreated in the purification section 10. B. The pre-treated hydrocarbon feed 11 and the optionally selected hydrocarbon-rich recycled 84 are mixed with process steam 44 and supplied to an optional pre-reformer 20. C. The pre-reformer effluent 21 and optional hydrocarbon-rich recycled 86 are mixed with process steam 42 and supplied to the tubular side of the heat exchanger reformer 35 to generate effluent effluent 37. D. The spillage 37 and optional hydrocarbon-rich recycling 85 are mixed with process steam 45 and supplied to the reformer 30, where a flow 31 containing high-purity oxygen O2 is used to produce reformer spillage 32. E. The reformer outflow 32 provides heat to the shell side of the heat exchanger reformer 35 in order to generate reformed logistics 36 which are routed to the heat recovery section 40. F. The cooled synthesis gas and any mixed vapor 43 from the heat recovery section 40 are introduced into the water-gas shift reaction section 50. G. The effluent 51 from the water-gas shift reaction section 50 is introduced into a hydrogen and carbon dioxide recovery unit 60, which separates the flow into a CO2-rich flow 61, a hydrogen-rich flow 62, and a hydrogen and CO2-depleted exhaust gas flow 63. H. A portion 64 of the waste stream 63 from the hydrogen and carbon dioxide recovery unit 60 is used as fuel for the heating furnace 90, while the remainder 64 or all of the waste stream 63 is compressed by the compressor 70. I. The compressed waste stream 71 enters a membrane separation system 80, which separates the stream into a hydrogen-rich permeable stream 82 and a hydrocarbon-rich concentrated stream 81. J. Part or all of the hydrogen-rich permeable logistics 82 is used as fuel for the heating furnace 90. K. Flue gas 92 from the heating furnace 90 heats the hydrocarbon feed and / or water flow 49 to generate steam 41 in the convection section 100 and the heat recovery section 40. L. Hydrocarbon-rich concentrated logistics 81 is recycled into one or more of the following: a) A portion 83 (if applicable) to be mixed with hydrocarbon feed 1 to the purification section. b) A portion 84 (if applicable) of the pre-treated hydrocarbons 11 and process vapors 44 are mixed into the pre-reformer 20. c) A portion 86 of the hydrocarbon flow 23 and process steam 42 are mixed with the heat exchanger reformer 35. d) A portion 85 of the reformer inlet flow 37 and process steam 45 is mixed into the reformer 30. e) A portion 87 of the cooled reformed material and process steam 43 is mixed into the water-gas shift reaction 50. M. The heating furnace receives the oxygen necessary for combustion via flow 91.

[0034] All changes to the flow scheme applied in Figure 1 can be similarly applied to the flow scheme in Figure 3, which includes the heat exchanger reformer 35 installed in series, without any modifications.

[0035] Including the heat exchanger reformer 35 in the flow scheme offers specific benefits to the heat transfer occurring within the plant. Whether installed in parallel or series, the heat exchanger reformer utilizes heat transfer by conduction from the tube walls to the catalyst and the fluid hydrocarbon flow, rather than by pure convection where the hydrocarbon flow is heated solely by the flue gas or process gas flow. The heat consumed in the heat exchanger reformer is no longer present in the reformer effluent 36, which is at a much lower temperature compared to the reformer effluent 32 when it enters the heat recovery section 40, thereby substantially reducing excess steam generation. Using high levels of heat instead of steam to generate hydrogen results in reduced combustion demand and further reduces fuel demand.

[0036] More specifically, this advantage has a cumulative effect when combined with the hydrogen permeate membrane separation system 80 in the recycle flow 65. As a result of the hydrogen permeate membrane, the flow rate of the hydrocarbon-enriched recycle flow 81 to the reformer feed section is reduced, and therefore less preheating of the feed in the furnace 90 is required. Since the hydrogen generated in the plant is used as fuel, the reduction in fuel demand also reduces the amount of hydrogen required by the plant, which further reduces the combustion demand. Due to the low total fuel demand of the furnace 90, the hydrogen-rich permeate flow 81 can provide 25% to 70%, and in some cases up to 100%, of the total fuel demand. The beneficial effect of the combination of heat exchanger reforming and the membrane separation system therefore further reduces the flow rate of hydrogen required as fuel. Thus, the present invention enables a significant reduction in hydrocarbon consumption.

[0037] In the process according to the present invention, hydrogen produced from a hydrocarbon feed is used as fuel to the furnace 90 instead of the hydrocarbon feed. The preferred fuel is hydrogen-enriched permeate 82 from a hydrogen permeate membrane separation system 80, and any necessary supplemental fuel is supplied as part of the product hydrogen 62. If further reforming reactants include steam, carbon dioxide is formed as a byproduct to obtain the hydrogen, rather than being formed by the combustion of hydrocarbons in the furnace 90, and forms part of the reformer 36 (process gas). Thus, the generated CO2 can be conveniently captured from the process gas along with the CO2 formed to produce part of the hydrogen that becomes part of the hydrogen-containing product, and can be extracted from the process rather than released into the atmosphere. Since hydrogen combustion increases feed consumption as more hydrogen is needed, a heat exchange reforming reactor is applied to minimize combustion demand. Therefore, the design of the reformer and the use of the generated hydrogen work together to reduce greenhouse gas emissions and / or the global carbon footprint. The carbon footprint is further reduced by including a pre-reformer in the reformer configuration and by using hydrogen combustion along with carbon dioxide capture to remove carbon dioxide from the process gas (reformer, shift reactor products). In a preferred embodiment, the present invention enables a reduction of up to 99.7% in direct emissions from the hydrogen plant. The preferred embodiment provides, in particular, an additional saving of nearly 15% in power consumption and its associated indirect CO2 emissions, as well as a reduction of up to 15% in hydrocarbon consumption, compared to state-of-the-art low-carbon emission hydrogen plants that have already achieved a 95% reduction in direct CO2 emissions. This is a significant improvement considering that removing CO2 becomes increasingly difficult. As shown in the examples, the preferred embodiment (Case 3) exhibits a capture rate of nearly 99.7% of direct CO2 emissions, which is a further improvement of 95.3% compared to a conventional operating base case (Case 1) that does not include a hydrogen membrane in the recycle flow. In addition, power consumption is also reduced, resulting in a 6% to 69% reduction in the global CO2 footprint, depending on the amount of renewable electricity in the energy mix. If available power has a low carbon intensity, a reduction of up to 69% in the total CO2 footprint can be achieved.

[0038] Furthermore, it should be noted that the examples illustrate that the present invention enables further reductions in carbon dioxide emissions while reducing the size of the reformer compared to state-of-the-art plant designs, as exemplified by Case 1. In fact, the overall plant size, including the front-end desulfurization 10, the back-end shift 50, and the hydrogen and carbon dioxide recovery sections 60, can be reduced compared to state-of-the-art plants.

[0039] In the last embodiment of the present invention, the heat recovery section 40 can be optimized to generate steam 41 at a much higher pressure than actually required by the process, while reducing the temperature pinch within the heat recovery section 40. The generated high-pressure steam 41 is then used to generate power in the back-pressure steam turbine 47, which can expand the flow 48 to the pressure level required by the process. This embodiment allows for a further reduction in the net power required to operate the plant, as shown in Figure 4. All modifications to the flow scheme applied in Figure 1 can also be applied without any modification to the flow scheme in Figure 4, which includes heat exchanger reformers 35 installed in parallel or in series.

[0040] The present invention is fully described herein with reference to the accompanying drawings, which illustrate embodiments of the invention, including several optional elements, e.g., a feed purification unit 10 (Figures 1, 2, 3, and 4), a pre-reformer 20 (Figures 1, 2, 3, and 4), a heat exchanger reformer 35 (Figures 2 and 3), and a steam turbine 47 (Figure 4). The locations of units and process lines may deviate from those schematically shown. For example, in some embodiments, within a hydrogen and carbon dioxide recovery unit 60, the carbon dioxide recovery unit is upstream of the hydrogen recovery unit, while in other embodiments, the order may be reversed. In certain embodiments of the hydrogen and carbon dioxide recovery unit 60, internal flows, e.g., a crude hydrogen flow 68 generated after CO2 removal from a shift gas flow 51 and rich in hydrogen but depleted of CO2, and a flash gas flow 69 obtained during the regeneration of an amine-based solvent by reducing pressure and evaporating dissolved gases into the gas phase, may be routed to the rest of the plant. In the drawings, the absolute and relative sizes of systems, components, layers, and areas may be exaggerated for clarity. Embodiments may be described with reference to schematic and / or cross-sectional views of the most idealized embodiments and intermediate structures of the invention. In the description and drawings, similar numbers refer to similar elements throughout. Relative terms and their derivatives should be interpreted as referring to the orientation described therein or shown in the drawings under consideration. These relative terms are for illustrative purposes only and do not require the system to be constructed or operated in a particular orientation unless otherwise stated.

[0041] A person skilled in the art can use this disclosure in combination with common general knowledge and, optionally, one or more of the documents referenced herein to design and operate units for a hydrogen plant or suitable operating units for use in a process according to the present invention. For example, a person skilled in the art can, based on this disclosure, the referenced documents and common sense, provide suitable process / plant units (e.g., reformer units, shift reactor zone units, carbon dioxide recovery units, hydrogen recovery units, heat exchanger units) and passages, such as pipes, lines, tubes or other channels for passing gas or liquid directly or indirectly from one processing unit to another.

[0042] For the purposes of clarity and concise description, features are described herein as part of the same or distinct embodiments, but it will be understood that the scope of the invention may include embodiments having all or some combinations of the described features.

[0043] Next, the method and plant according to the present invention will be described in more detail.

[0044] The hydrocarbon feedstock 1 supplied to the reformer system can be any hydrocarbon feedstock suitable for reforming by reaction with steam. In particular, it can be a feedstock consisting of at least substantially methane, such as natural gas, biogas-based methane streams, or e-methane, propane gas (LPG), naphtha or refinery off-gas, or a renewable feedstock or a stream with an equivalent composition derived from the hydrogenation of a corresponding e-fuel.

[0045] Depending on the purity of the feedstock, it may be subjected to pretreatment such as hydrodesulfurization in the pretreatment section 10. Pretreatment, conditions therefor, and suitable pretreatment units can be based on known technologies. In particular, when using pretreatment such as hydrodesulfurization, a hydrogen-containing supplemental stream is usually added to the feedstock to ensure the purification of the feedstock in the hydrodesulfurization section. As described below, hydrogen-containing streams produced in the process according to the present invention, in particular hydrogen product gases, waste gases from hydrogen and carbon dioxide recovery units, or hydrocarbon-rich recycled gases from membrane separation systems can be used for this purpose. The pretreatment section 10 may be configured to produce a treated hydrocarbon stream 11.

[0046] The hydrocarbon feedstock is mixed with further reformer reactants, namely water (steam), carbon dioxide, hydrocarbon-rich recycled gas, or a mixture thereof, before being subjected to the reaction in the reformer reaction unit (30 and optionally 35). The reformer 30 converts the flow containing the hydrocarbon feedstock into a reformed gas flow 32.

[0047] The reformer 30 is an exothermic reformer. The reformer 30 is, for example, an autothermal reformer (ATR) or a partial oxidation (POX) reformer configured to carry out autothermal reforming or partial oxidation reforming. The reformer is configured to convert a flow containing hydrocarbon feedstock 1 into a reformed gas flow 32 by conversion using oxygen-rich steam 31. The oxygen-rich steam 31 and the flow containing feedstock 1 can be mixed in the reformer 30 or upstream of the inlet of the reformer 30.

[0048] The oxygen-rich vapor 31 is a high-purity oxygen stream that may contain at least 90% (by volume) of oxygen. In some configurations, the high-purity oxygen stream 31 entering the reformer 30 may contain at least 95% (by volume) or at least 98% (by volume) of oxygen. The high-purity oxygen stream may contain small amounts of vapor, for example, less than 5% by volume, and optionally less than 1% by volume. The stream containing the hydrocarbon feedstock 1 may be mixed with an additional vapor stream 45 upstream of or within the reformer 30. The additional vapor stream 45 may contain a larger amount of vapor than the stream 31, for example, most of the vapor required for the reforming reaction may be supplied in the stream 45.

[0049] In some embodiments, the reformer 30 may include a reaction vessel containing a packed bed of catalyst (e.g., a reformer catalyst). The packed bed of catalyst may be provided above the bottom of the reaction vessel, thereby allowing the fluid in the vessel to flow from the top of the reaction vessel, through the packed bed of catalyst, and then further downward to the outlet of the vessel, e.g., a bottom outlet nozzle of the reaction vessel. In some embodiments, the reformer 30 may be referred to as a fixed-bed or packed-bed reactor reformer. In contrast, an SMR reformer does not have a packed bed of catalyst installed within the reaction vessel. Instead, an SMR reformer typically includes a number of tubes (e.g., tens to hundreds) extending through the reformer, which are filled with catalyst, over which the reformer feed passes, and within which the reforming reaction takes place. Advantageously, loading the catalyst bed into the reformer 30 (for example, at the base of the ATR or POX reformer vessel) requires less effort than loading the catalyst into an SMR reactor, in which dozens to hundreds of tubes need to be loaded with catalyst in a similar manner, and each tube requires individual pressure drop measurements to ensure a uniform flow distribution of the reformer feed gas across the reformer catalyst.

[0050] The reaction vessel of the reformer 30 may be a refractory-lined reaction vessel to protect the vessel shell from high temperatures inside the vessel. In some configurations, the vessel shell may be water-cooled, for example, by a water-cooling jacket. In some configurations, the reaction vessel includes a catalyst-packed bed located at the bottom of the vessel. The catalyst may be a nickel-based catalyst. The height of the packed bed (e.g., catalyst) may be less than half the height of the reaction vessel. The packed bed may be about one-third of the total height of the reactor. In some configurations, the catalyst-packed bed may have a diameter in the range of 2000 to 5500 mm and a height in the range of 2000 to 7000 mm.

[0051] The reformer 30 may include a burner located in the upper region of the container (e.g., above the packed bed). In some configurations, multiple burners are provided. For example, several hundred thousand Nm 3 h -1 For very large reforming reactors that generate reformed gas, some configurations can provide up to four burners. This is approximately 15,000 Nm³ 3 h -1 This is in contrast to convection reformers, which only produce reformed gas and therefore can only have a single burner or a very small number of burners (e.g., fewer than four). One or more burners are configured to at least partially combust the hydrocarbon feedstock mixed with vapor in the presence of high-purity oxygen.

[0052] The upper region of the vessel may be substantially conical in shape. The upper region of the vessel and one or more burners are arranged so that flames from the burners do not come into contact with the catalyst bed. In some configurations, a heat-resistant layer may be provided above the top surface of the catalyst bed. In some configurations, the heat-resistant layer may extend over the entire top surface of the bed. The heat-resistant layer is configured to protect the catalyst from the high-temperature burner flames. The heat-resistant layer is configured to allow gases from the upper region to pass through the layer into the bed below. The reformer 30 may include an outlet nozzle at or toward the base of the vessel, through which reactor effluent (e.g., reformed gas 32) exits the reformer 30.

[0053] Configuring the reformer 30 to perform ATR or POX reforming is more advantageous than a conventional SMR reformer because the steam requirements for ATR or POX reformers are lower than those for SMR reformers. Specifically, ATR or POX reforming uses oxygen-rich steam, while SMR requires only steam. Utilizing oxygen-rich steam facilitates the generation of additional steam during the reforming reaction (e.g., through the combustion of hydrocarbons with oxygen). The steam requirements of a plant with a main reformer configured for SMR are greater than those of a plant with an ATR or POX reformer. Thus, since ATR or POX reforming is implemented in the reformer 30 rather than in SMR, the energy requirements for generating steam in the plant are lower.

[0054] In contrast to conventional endothermic reformers (e.g., those performing SMR), reformer 30 is configured to perform exothermic oxygen-based reforming (e.g., ATR or POX reforming) to produce a reformed gas stream 32 containing hydrogen, carbon monoxide, carbon dioxide, and at least one hydrocarbon as an impurity. In an SMR reformer, the endothermic reaction requires an additional heat supply to drive the reforming reaction. From this perspective, an SMR reactor is traditionally a combustion-type tubular reformer with radiating and convection sections, to which additional fuel is supplied to provide reaction heat. In reformer 30 (e.g., performing ATR or POX reforming), the heat generated by the exothermic reaction is sufficient to drive the process, and therefore no additional heating (e.g., supplying reaction heat) is required to maintain the reaction. Thus, reformer 30 is not a combustion-type tubular reformer, and no additional heat is introduced into the reactor. The process in reformer 30 is self-sufficient compared to a conventional SMR reformer. The energy requirements for supplying heat to an ATR or POX reformer (e.g., in a heating furnace) are lower than the energy requirements for supplying heat to an SMR reformer.

[0055] A plant having an ATR or POX reformer 30 includes a separate furnace 90 for providing the necessary flow preheat before entering the reformer (instead of the radiative and convection sections within the reformer itself to provide reaction heat, as in SMR reforming). The furnace 90 can advantageously utilize the hydrogen-enriched permeate 82 generated downstream of the process as a low-carbon fuel. In some embodiments, one or more furnaces 90 may be provided, for example, in parallel or in series with each other. Providing multiple furnaces 90 can improve redundancy if one or more of the furnaces 90 require maintenance. One or more furnaces 90 are provided as components separate from the reformer 30.

[0056] The heating furnace 90 (e.g., a small heating furnace) is required only to preheat the hydrocarbon feedstock to the reactor inlet temperature and is not required to supply the heat of reaction for the reforming reaction. The hydrocarbon feedstock is heated by the heating furnace 90 before being introduced into the reforming reactor. The feed containing the hydrocarbon feedstock may be heated to at least about 600°C by the heating furnace 90 before being introduced into the reformer 30. In some configurations, the feed containing the hydrocarbon feedstock may be mixed with steam (e.g., flow 45) before the mixture is heated to at least about 600°C by the heating furnace 90 upstream of the reformer 30.

[0057] The heating furnace 90 is provided as a separate component from the reformer 30, for example, outside the reformer 30. In embodiments where a pre-reformer 20 and / or a heat exchanger reformer 35 are provided, the heating furnace 90 is provided as a separate component from the reformer 30, the pre-reformer 20, and / or the heat exchanger reformer 35. A reference to a heating furnace 90 separate from the reformers 20, 30, and 35 will be understood to refer to a heating furnace 90 that does not share common structural elements with the reformers 20, 30, and 35. In embodiments, the only connection between the heating furnace 90 and the reformer 30 is via a tube that guides the preheated feed from the heating furnace 90 to the reformer 30.

[0058] As described above, oxidative (e.g., ATR or POX) exothermic reforming offers several process advantages compared to SMR reforming, such as reduced fuel demand, steam demand, and hydrocarbon feed consumption. In addition to these, physically separating the furnace 90 from the reformer 30 provides further advantages over endothermic reforming techniques that include radiative and convection sections. For example, the process steps of feed preheating and reforming reactions can be separated into smaller individual appliance items. Thus, since individual appliance items occupy less space than a combined reformer and heater configuration, the layout of the hydrogen plant can be more easily adapted to the available area. Furthermore, the orientation of the furnace 90 relative to the reformer 30 can be adjusted to be permissible by the available plant space. By providing a separate furnace 90, the complexity of the piping systems to and from the reformer 30 is also reduced, further contributing to a reduction in plant area requirements.

[0059] It will be understood that the heating furnace 90 is configured to heat the feedstock only in front of or upstream of the reformer 30, and not inside the reformer 30, in contrast to the radiant and convection sections of the SMR reformer, where the reaction heat is provided inside the reformer itself. No reforming reaction takes place inside the heating furnace 90. The reforming reaction takes place outside the heating furnace 90 in each reformer. In contrast to the described arrangement, the heat provided in the radiant and convection sections of the SMR reformer provides the duty for the SMR reaction.

[0060] The heating furnace 90 may have a chamber and one or more tubes extending through the chamber. In some configurations, the chamber may be called the combustion chamber. The chamber of the heating furnace 90 is separated from the reformer reactor 30 and is located outside the reformer reactor 30. The heating furnace 90 is configured such that the fuel is combustible in the chamber and provides heat to one or more tubes. One or more tubes are configured to transport a flow containing hydrocarbon feedstock 1 (and vapor flow 45) through the chamber of the heating furnace 90 into the reformer 30 so that the flow containing hydrocarbon feedstock 1 is preheated in the heating furnace 90 before it enters the reformer 30. In some configurations, the only connection between the heating furnace 90 and the reformer 30 is through the tubes. As will be described in more detail, the hydrogen-enriched permeate 82 obtained downstream of the reformer may be led into the heating furnace 90 (e.g., the combustion chamber) so that it can be used by the heating furnace 90 as a low-carbon fuel. Other additional flows may be used as fuel for the heating furnace 90 in some configurations (for example, fresh fuel needs to be provided at startup before the hydrogen-enriched permeate 82 is generated).

[0061] The chamber of the heating furnace 90 may include one or more burners to which a fuel and oxygen-containing flow 91 is supplied. The burners may be installed in the base region of the chamber, and the hot flue gas exits from the upper region of the chamber. This configuration allows the burners to be accessed from the ground, enabling simple assembly and maintenance. This is suitable for large-capacity plants (e.g., 15,000 Nm³). 3 h -1 Hydrogen production, for example, 200,000 Nm 3 h -1 In cases up to or exceeding a certain temperature, the burner needs to be installed on the wall (e.g., side wall) or top wall of the SMR reformer to optimize heat dissipation from the flame for the endothermic reforming reaction occurring in the tubes extending through the reformer. Thus, the SMR reforming system involves a more complex structure, more difficult access to the burner, and a more complex piping arrangement compared to the exothermic reformer 30 and separate heating furnace 90 of the embodiments of the present invention.

[0062] In some configurations, the burner of the combustion heater 90 operates at high pressure (e.g., about 3 bar g) within the combustion chamber. In alternative configurations, the combustion chamber may operate at lower pressures; for example, the fuel may be supplied at or near atmospheric pressure, and the combustion chamber may operate at a pressure slightly below atmospheric pressure.

[0063] The tubes extend through the chamber (e.g., through the upper region of the chamber). The tubes may be in the form of radiating coils. In some configurations, the tubes have treated outer surfaces to improve heat transfer properties (e.g., finned tubes to increase the area available for heat transfer). It will be understood that the reforming reaction of hydrocarbon feedstock 1 or vapor does not occur inside the furnace 90. For example, since there is no catalyst in the tubes of the furnace 90, the reforming reaction does not occur there. The flow passing through the tubes inside the furnace 90 does not undergo the reforming reaction and is only preheated before leaving the furnace 90 and being introduced into the reformer for the ATR or POX reforming reaction.

[0064] In some configurations, the tubes are positioned to extend substantially perpendicular to the direction of flue gas flow from the burner (for example, substantially perpendicular to the direction of the burner flame's extension within the chamber). This promotes cross-flow of flue gas on the tubes and improves heat transfer between the tubes. In alternative configurations, the tubes may be positioned to extend substantially parallel to the direction of flue gas flow from the burner to promote parallel flow of flue gas on the tubes.

[0065] The plant may include a convection section 100 located outside the chamber of the furnace 90. The plant may be configured so that flue gas exits the chamber of the furnace 90 and enters the external convection section 100. The convection section 100 may be located above the chamber of the furnace 90 (e.g., at the top of the chamber). Such an arrangement can reduce the footprint plot area occupied by the furnace, as the convection section is located above the combustion chamber instead of adjacent to it. Further heat may be recovered from the flue gas in the convection section 100 (e.g., for use elsewhere in the plant). The cooled flue gas can exit the convection section via a stack. Tubes containing hydrocarbon feedstock (and vapors) may extend within the convection section 100 of the furnace 90 (e.g., in addition to, or as an alternative to, extending through the combustion chamber of the furnace 90).

[0066] The oxygen-rich steam 31 may be preheated before being introduced into the reformer 30. The oxygen-rich steam 31 may be preheated separately from the hydrocarbon feed and additional steam streams, for example, in a dedicated heat exchanger. In some configurations, the oxygen stream 31 may be preheated to about 200-230°C.

[0067] The reformer 30 includes a heat recovery section 40. The heat recovery section 40 is configured to receive heat from at least a portion of the reformed gas 32 generated in the reformer 30. Means, such as conduits or passages, are provided for sending at least a portion of the reformed gas 32 to the heat recovery section 40. The heat recovery section 40 may be configured, for example, to generate process vapor for use in the reformer 30.

[0068] 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 pre-reformer reaction unit 20 may be located immediately upstream of the reformer reaction unit 30 (e.g., an 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 configurations (i.e., the ATR of the POX reformer may exist separately or together with only the heat exchanger reformer 35).

[0069] Using one or more pre-reformer units 20, usually one or more insulated pre-reformer units, to partially carry out the reforming reaction (before preheating the pre-reformed mixture to the inlet temperature of the main reformer in the heating furnace 90) is advantageous in reducing the load on the reforming reaction. In pre-reforming, generally, a small portion of the hydrocarbons is converted, thereby forming CO, among other things. The pre-reformer 20 is configured to produce a pre-reformed synthesis gas flow 21 from the treated hydrocarbon flow 11 and / or hydrocarbon feedstock 1.

[0070] The pre-reformer reaction unit 20 may be a single-vessel reactor containing a reforming catalyst. The diameter of the vessel may be in the range of 1000 to 4000 mm. The height of the vessel may be in the range of 1500 to 10000 mm. In some configurations, a second pre-reformer may be installed in parallel to facilitate catalyst maintenance (e.g., catalyst replacement) for one pre-reformer without shutting down the entire plant.

[0071] The mixture supplied to the reformer (or to the pre-reformer, if one is used) typically consists of at least substantially hydrocarbons and further reactants.

[0072] Steam and hydrocarbon feed can be supplied to the reformer 30 in ratios known in the art. Typically, the ratio of steam to carbon supplied to the reformer reaction unit is at least 1.0 mol / mol, preferably at least 2.0 mol / mol, and particularly at least 3.0 mol / mol. Typically, the ratio of steam to carbon supplied to the reformer reaction unit is 5.0 mol / mol or less, preferably 4.0 mol / mol or less, and more preferably about 3.0 mol / mol or less. A steam-to-carbon ratio of 1.0 to 2.0 mol / mol or less is generally preferred because it minimizes hydrocarbon consumption and CO2 emissions. The steam-to-carbon ratio may be in the range of 0.6 to 2.5, preferably 0.8 to 1.2.

[0073] The present invention enables the addition of steam at multiple process locations, such as a flow 44 for addition to the pre-reformer 20, a flow 45 for addition to the reformer 30, a flow 43 for addition to the water-gas shift section 50, and a flow 42 for addition to the heat exchanger reformer 35. The overall ratio of steam to hydrocarbons is 3.0 mol / mol, preferably less than that, for example, 2.5 or 2.0, and the ratio for the individual sections for the pre-reformer and reformer can be less than 1.0 mol / mol, while the ratio can be maintained higher than 1.0 mol / mol, preferably higher than 2.0 mol / mol, as required by the catalyst for the water-gas shift and heat exchanger reformer 35.

[0074] According to the present invention, the combustion duty required to preheat the feedstream and drive the reforming reaction within the reformer reaction unit is significantly reduced by applying the heat exchanger reformer reactor 35. As already described, such a reduction in the combustion duty cannot be achieved by additional preheating of the feed before it enters the reformer reaction unit.

[0075] A heat exchanger reformer unit 35 is optionally provided for a process or plant according to the present invention. The heat exchanger reformer may be based on heat exchangers known in the art, for example, from references WO2018 / 104526, WO2011 / 077107, or WO2012 / 057922. The heat exchanger reformer can be conceived as a shell-and-tube heat exchanger in which the tubes are filled with a reformer catalyst. A fresh or partially converted feed is introduced into the tube side, while a hotter reformed gas is introduced into the shell side to provide heat of reaction for the endothermic reforming reaction. The heat exchanger reformer 35 is configured to generate a reformed flow 37 at the tube side outlet. When the heat exchanger reformers 35 are installed in series, their tube side and shell side outlets are kept separate and routed to different destinations in the process flow scheme. When the heat exchanger reformers 35 are installed in parallel, the tube-side inlet outlet and the shell-side inlet are mixed at the outlet of the heat exchanger reformer, providing a single mixed flow to the downstream process. The tube-side outlet and the shell-side inlet may be mixed inside or outside the heat exchanger reformer. Either configuration (e.g., parallel or series) can be applied to the present invention, and either configuration utilizes the heat exchange that takes place between the high-temperature effluent and the lower-temperature, less converted hydrocarbon feedstock to achieve increased feedstock conversion with reduced combustion demand and reduced steam generation. In some embodiments, the heat exchanger reformer 35 comprises a reformer unit 30 and does not include a pre-reformer 20. In alternative embodiments, the plant may include a pre-reformer 20, a reformer unit 30, and a heat exchanger reformer 35. It will be understood that the heat exchanger reformer 35 is an optional unit that may not be present in some arrangements (e.g., a self-heating or POX reformer may exist separately or in combination with only the pre-reformer 20).

[0076] In some embodiments, the pre-reformer 20 may include a reformer unit 30 and may not include a heat exchanger reformer 35. In such embodiments, all of the pre-reformed synthesis gas flow 21 generated in the pre-reformer 20 can be supplied to the reformer reaction unit 30 (for example, after preheating by a heating furnace 90).

[0077] In an alternative embodiment, the plant may include a pre-reformer 20, a reformer unit 30, and a heat exchanger reformer 35. In an arrangement where the heat exchanger reformer 35 is installed in series with the reformer 30, the plant is configured such that the pre-reformed synthesis gas flow 21 formed in the pre-reformer 20 is supplied to the tubular side of the heat exchanger reformer 25. The plant is configured such that the reformed flow 37 generated at the tubular side outlet of the heat exchanger reformer 35 is supplied to the reformer 30 (for example, after preheating by a heating furnace 90). The plant is configured such that the effluent 32 from the reformer 30 is supplied to the shell side of the heat exchanger reformer 35. The reformed material 36 exits from the shell side of the heat exchanger reformer 35.

[0078] In a configuration where the heat exchanger reformer 35 is installed in parallel with the reformer 30, the plant is configured so that a portion 23 of the pre-reformed synthesis gas flow 21 formed in the pre-reformer 20 is supplied to the tubular side of the heat exchanger reformer 35. The plant is configured so that the remaining portion 22 of the pre-reformed synthesis gas flow 21 is supplied to the reformer 30 (for example, following preheating in the furnace 90). The plant is configured so that the reformer effluent 32 is supplied to the shell side of the heat exchanger reformer 35. The reformed flow 37 exiting the tubular side is mixed with the reformer effluent 32 supplied to the shell side inlet of the heat exchanger reformer 35, either inside or outside the heat exchanger reformer 35. The reformed material 36 exits from the shell side outlet of the heat exchanger reformer 35.

[0079] The heat exchanger reformer is installed in a parallel flow path for a mixture of hydrocarbon feedstock and steam. In the parallel configuration, a portion of the feedstock to the reformer system is branched 23 and sent to a heat exchanger reformer reactor 35 in parallel with the reformer 30. The heat for the reaction in the parallel heat exchanger reformer 35 is supplied by the high-temperature reformer effluent 32, the heat available in the flue gas from the radiating section 12, or by another high-temperature heat source generally at least 850°C, particularly above 900°C.

[0080] The advantage of a parallel heat exchanger reformer is that, during use, it supplies a portion of the duty cycle required for hydrogen production, thereby reducing the duty cycle required in the reformer (provided by the partial combustion of hydrocarbon feedstock), and therefore the oxygen consumption and associated power. Generally, the outlet temperature of the catalyst zone (bed) in the heat exchanger reformer unit is lower than that in the reformer due to the temperature difference (driving force) required for the heat exchange process to occur. The outlet temperature of the catalyst bed is typically in the range of 850 to about 1000°C, but is lower than the outlet temperature of the reformer 30, usually at least about 30 to 50°C lower. The pressure inside the heat exchanger reformer 35 is generally about equal to the pressure inside the reformer. Due to the lower outlet temperature, methane slip is typically higher than in the reformer. To reduce methane slip, additional steam 42 or carbon dioxide or a mixture thereof is added at the feed port of the parallel heat exchanger reactor 35 to direct the reforming reaction toward hydrogen production. Therefore, the ratio of further reactants (steam, carbon dioxide, or mixtures thereof) to hydrocarbon feed does not need 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, and then undergoes further processing, generally including further conversion in the water-gas shift section 50 (see also below). The heat exchanger reformer is in parallel with the reformer 30, and since it reduces the required duty cycle of the reformer (when compared at equal hydrogen output), the size of the reformer can be reduced. When using a parallel configuration, typically about 10 to about 30 wt%, preferably 15 to 25 wt%, of hydrocarbon feed is supplied to the heat exchanger reformer reaction unit 35. A higher split ratio further reduces the load on the reformer but reduces the driving force of the heat exchanger reformer, resulting in an increasingly larger heat exchanger reformer.

[0081] Another configuration uses a process flow scheme in which a heat exchanger reformer 35 is installed in series with a reformer 30, where all feed and steam mixture is delivered through a catalyst bed in the multi-tube heat exchanger reformer 35 to form a partially converted reformed flow. This flow typically exits the heat exchanger reformer at below 850°C, more preferably below 750°C, before entering the reforming section 30. The partially converted reformed gas 37 then reacts further in the reforming section 30. Heat for the heat exchanger reformer 35 is supplied by the reformed effluent 32 from the reforming section 30. Similar advantages to the parallel configuration can be achieved, and to avoid repetition, a detailed explanation is not included here.

[0082] The reformed material 36 from the heat exchanger reformer 35 is typically cooled in a heat recovery section 40 before further processing. This can be done by methods known to the extent of the invention. Particularly useful is a waste heat boiler, which generates steam using the heat from the reformed material during use. The steam 41 is more preferably generated using the heat from the flue gas from the heating furnace 90. This generally takes place in a heat exchanger configured in a heat exchange configuration within a convection section 100. The generated steam or a portion thereof (42, 43, 44, 45) is used as steam to react with the hydrocarbon feed to be used in the reaction in the reforming. In principle, the generated steam 43 can also be mixed with the reformed material 36 before the shift reactor zone 50 or supplied to the process gas inside the shift reactor. However, good results can be obtained without adding steam to the reformed material or into the shift reactor zone. The method or plant according to the present invention can be adapted to be completely self-sufficient in steam generation for the process, or can be based on outgoing and / or intake steam 15. The lower the required steam generation within the unit battery limits, the more thermal integration can be applied, and therefore the lower the combustion requirements.

[0083] Advantageously, in some embodiments, the steam may be generated at a higher pressure than required by the reforming process. Under such conditions, some or all of the generated steam can be expanded in the steam turbine 47 to recover the energy generated during expansion as electrical energy, thereby reducing the plant's net power intake. Where the carbon intensity of the power is high, this can be an effective means of reducing the plant's direct and indirect CO2 emissions. In such embodiments, the steam is generated at a pressure 0.1 MPa to 15 MPa higher than the normal process pressure, preferably more than 8 MPa higher than the process pressure. The expanded steam 48 at a pressure close to the process pressure is then used in the process. Any remaining discharged steam 46 after the extraction of process steam can be used directly or further expanded and condensed, and the condensate is returned to the plant to generate fresh steam.

[0084] The reformed material is typically cooled and then supplied to a water-gas shift reactor 50, where carbon monoxide reacts with water to form further hydrogen and carbon dioxide. The reformed material supplied to the water-gas shift reactor 50 may be / include the reformed gas stream 32 generated by the reformer 30 (for example, in embodiments where a heat exchange reformer 35 is not provided), or may be / include the reformed material 36 from the heat exchange reformer 35. The reaction in the water-gas shift reactor 50 is typically carried out in the presence of a shift catalyst, such as an iron or copper-based shift catalyst, which is known in itself. Thus, the processing in the shift reactor zone results in a shift reactor product (shift reactor process gas) having increased hydrogen and carbon dioxide content compared to the reformed material. The design and reaction conditions of the shift reactor zone may, in principle, be based on known techniques, such as those described in the prior art cited herein.

[0085] The shift reaction zone generally operates at lower temperatures than the reformer system. Generally, the temperature during the shift reaction is in the range of approximately 190 to 500°C. The types of shifts applied are generally categorized into three types based on the catalyst outlet temperature: high-temperature shifts with an inlet temperature of 300–400°C and an outlet temperature of 350–500°C, medium-temperature shifts with an inlet temperature of 190–230°C and an outlet temperature of 280–330°C, and low-temperature shifts with an inlet temperature of 180–230°C and an outlet temperature of 200–250°C. Since the water-gas shift reaction is exothermic, the temperature rise is greater with higher CO concentrations at the inlet.

[0086] To minimize carbon dioxide emissions, it is advantageous to operate the shift reactor zone under conditions where the water-gas shift reaction (CO + H2O ⇔ CO2 + H2) is shifted toward the formation of H2 and CO2. Since the reaction is exothermic, this is promoted by low temperatures. Preferably, in the process according to the present invention, a low-temperature shift is applied following a high-temperature shift to maximize the conversion of CO to H2 and CO2 (this CO2 is subsequently captured), and thus minimize CO2 emissions. Alternatively, only a high-temperature shift, a medium-temperature shift reaction, or an isothermal shift reaction (a shift reactor cooled to a constant temperature), or any combination thereof, can be applied.

[0087] Advantageously, the performance of the water-gas shift reaction (CO + H2O ⇔ CO2 + H2) can be improved by selectively extracting hydrogen from the reformed material (32) or (36), thereby shifting the reaction equilibrium and generating more H2 and CO2, which in turn results in improved conversion of the hydrocarbon feedstock to hydrogen and carbon dioxide. In an exemplary configuration, 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 it enters the water-gas shift reactor. In embodiments where 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 to extract or remove hydrogen from the reformed material 36 from the heat exchanger reformer 35 before it enters the water-gas shift reactor 50. The extracted hydrogen stream can be used as fuel for a heating furnace (90). A membrane system based on the selective permeation of hydrogen through a palladium sheet at high temperatures (300-350°C) is particularly well-suited for this application.

[0088] Advantageously, at least the reformer reaction unit 30 operates at a relatively high temperature, preferably in the range of 850-1050°C, more preferably in the range of 900-1050°C (at the catalyst outlet end), and a deep shift conversion (resulting in high CO conversion to CO2) is applied in the shift reactor zone at a shift reactor outlet temperature typically in the range of 200-250°C to maximize the conversion of feed to hydrogen, thereby reducing CO2 emissions and hydrocarbon consumption in accordance with the present invention.

[0089] Typically, according to the present invention, the reformer 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, if processing is carried out in the shift reactor zone as usual) is subjected to carbon dioxide removal treatment in the hydrogen and carbon dioxide recovery section 60. Here, the carbon dioxide content in the reformer 36 or the shift reactor process gas 51 is reduced to form carbon dioxide depletion process gas (carbon dioxide depletion product) and carbon dioxide by-product 61. CO2 capture can be achieved by methods known in themselves, for example, as described in the cited prior art. Advantageous techniques include 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 processes. Typically, an amine-based adsorption / stripping process is preferred, where the entire duty cycle of the amine reboiler is supplied by the process gas, thereby minimizing external heat input and allocating external heat to CO2 emissions. Capturing CO2 from the process gas (reformer or shift reactor product) is preferred because it is the least energy-intensive, as the process gas is available at high pressure (typically about 2-3.5 MPa) and therefore has high driving force for CO2 separation. The captured CO2 is sent as a high-purity stream 61 to the battery limit (outflow stream from the plant) and can be used for other processes, the food and beverage industry, and storage, such as in Carbon Capture for Utilization and Storage (CCUS). Furthermore, the captured CO2 can be used as a reformer reactant when a dry reforming process is applied. This may be a different reformer process from the present reformer process.However, it is also possible to recycle the carbon dioxide-rich gas obtained from the carbon dioxide removal treatment directly or, after further purification to a higher carbon dioxide content, into the reformer 30 or heat exchanger reformer 35 in the process according to the present invention.

[0090] In certain embodiments of the CO2 capture technology, which includes amine-based absorption / stripping, amine-based solvent regeneration under reduced pressure can generate an additional waste flow (69) at intermediate pressure, which can also be integrated into the rest of the plant as fuel to the furnace 90, or as feed to the membrane separation system 80, or as an additional feed to the hydrogen purification unit. Similarly, a portion of the treated gas flow obtained after CO2 removal, which is H2-rich and CO2-depleted, can be routed as a flow (68) at full pressure to other sections of the plant, such as the fuel inlet to the furnace (90) or the feed inlet to the membrane separation system (80), to generate additional low-carbon hydrogen fuel.

[0091] Typically, 70-99% of CO2 is removed from process gases, but technically, 99.99% or more of CO2 can be removed. This represents a reduction of approximately 30-60% of the total CO2 emissions from a hydrogen production plant.

[0092] While CO2 capture in conventional reforming, involving CO2 removal from flue gas 101, is also known in the art, it should be noted that this is highly energy-intensive and requires more complex technology compared to the present invention, especially when the objective is to remove at least substantially all carbon dioxide from the flow. As shown in the examples, according to the present invention, it is possible to achieve a reduction of more than 98% in CO2 emissions by capturing CO2 from the process gas (upstream of the unit from which hydrogen product gas is obtained, or in the tail gas, i.e., waste gas, from the hydrogen recovery unit) without requiring CO2 capture from flue gas 101. Furthermore, from a plant size viewpoint, it is advantageous not to use / provide a capture device configured to remove carbon dioxide from flue gas. Moreover, in (atypical) embodiments in which the flue gas in the process according to the present invention contains a considerable amount of carbon dioxide, capture of carbon dioxide from flue gas 101 is further feasible.

[0093] As those skilled in the art will understand, the reformer, shift reactor products, and carbon dioxide depletion products each still contain substantial amounts of components other than hydrogen. Therefore, generally, hydrogen recovery in the hydrogen and carbon dioxide recovery unit 60 is carried out to obtain a hydrogen-containing product gas 62 of satisfactory purity, typically at least about 95 mol%, preferably at least 98%, more preferably at least 99 mol%, and especially at least 99.9 mol%. The purity may be 100% or less, particularly 99.9999% or less, 99.999% or less, 99.99% or less, 99.99% or less, 99.99% or less, 99.5% or less, or 99.0% or less, depending on the need and technology used. Suitable techniques for recovering hydrogen from process gases (e.g., reformer, shift reactor products, and carbon dioxide depletion products, respectively) can be based on known technologies. In a particularly preferred embodiment, a pressure swing adsorption (PSA) unit is provided. The PSA enables 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. Using the electrochemical compressor, a high-purity, high-pressure hydrogen-containing product gas 62 can be obtained.

[0094] The shift reactor product 51 is subjected to a carbon dioxide removal treatment in the hydrogen and carbon dioxide recovery unit 60 to obtain a carbon dioxide depletion product, a portion of which is used as fuel in the heating furnace 90. The hydrogen-rich gas 62 is recovered from the carbon dioxide depletion product using pressure swing adsorption, a hydrogen retention membrane, or an electrochemical compressor. In a more preferred embodiment, the carbon dioxide depletion product is separated into a hydrogen-enriched gas, preferably a hydrogen-containing product 62 having an increased hydrogen content compared to the carbon dioxide depletion product, and a tail gas 63 (waste gas) having a decreased hydrogen content compared to the carbon dioxide depletion process flow. The tail gas contains hydrocarbons, and a portion of the tail gas product 64 is preferably mixed with other fuel components such as the hydrogen-containing product before being used as fuel in the heating furnace 90. A portion of the tail gas 65 (waste gas flow from the hydrogen and carbon dioxide recovery unit 60) is compressed to a pressure of 0.5 to 5 MPa, preferably 1 to 3 MPa, and supplied to a membrane separation system 80, which separates the compressed tail gas 71 into a low-pressure (i.e., 0.05 to 1 MPa) hydrogen-rich permeate flow 82 and a high-pressure (i.e., 0.5 to 5 MPa) hydrocarbon-rich concentrate 81 containing CH4, CO, some H2, and typically inert substances such as nitrogen. The membrane separation system 80 is selective for hydrogen permeation. The hydrogen-rich permeate is used as fuel for the heating furnace 90. In particular, in this embodiment of the present invention, it is beneficial to have a first tail gas compression step up to the operating pressure of the hydrogen permeate membrane unit (typically 0.5 to 1.5 MPa) and a second compression step of the hydrocarbon-rich concentrate flow 81 to the required pressure upstream of the reformer. Since the volumetric flow rate of the hydrocarbon-rich concentrated logistics 81 is significantly reduced compared to the compressed waste logistics 71, this will result in a substantial reduction in the total compression power and total compressor capital costs required.

[0095] In a more preferred embodiment, a portion of the CO2 depletion products from the CO2 recovery unit within the hydrogen and carbon dioxide recovery unit 60 is supplied to a membrane separation system 80, which separates it into a hydrogen-rich flow 82 used as fuel in a heating furnace and recycles the hydrocarbon-rich flow 81. In this embodiment, all of the waste gas 63 is compressed and sent to the compressor 70 and the membrane separation system 80.

[0096] Recycling the hydrogen-enriched permeate 82 used as fuel in the combustion heat 90 favorably reduces the plant's fuel requirements. Recycling the hydrogen-enriched permeate 82 is particularly advantageous because the reformer is configured for exothermic ATR or POX reforming, which has fewer heating (and therefore fuel) requirements compared to conventional endothermic SMR reformers. In this configuration, the heating furnace 90 requires only enough fuel to preheat the feed to the reformer 30, and therefore recycled hydrogen-enriched permeate 82 can constitute a larger proportion of the fuel, resulting in less flue gas emissions (compared to using hydrocarbon-rich fuel).

[0097] Since the ATR or POX reformer 30 has lower fuel requirements compared to conventional SMR technology, in some embodiments, not all of the hydrogen-enriched permeate 82 may be needed as fuel in the furnace 90. A portion of the hydrogen-enriched permeate 82 can be used as fuel for the furnace 90. The remaining portion of the hydrogen-enriched permeate 82 can be recycled to the inlet of the hydrogen recovery unit of the hydrogen and carbon dioxide recovery unit 60 to extract additional valuable hydrogen from the plant.

[0098] In another preferred embodiment, a portion of the hydrogen product 62 obtained in the hydrogen and carbon dioxide recovery unit 60 can also be used together with the flue gas 101 to provide fuel for the furnace 90 when very low CO2 emissions are desired. In this embodiment, all of the waste gas 63 is compressed and sent to the compressor 70 and membrane separation system 80, while the hydrogen-rich permeate 82 is then recycled to the inlet of the PSA unit in the hydrogen and carbon dioxide recovery unit 60.

[0099] Furthermore, the hydrogen product 62 can be used to provide hydrogen for hydrocarbon feedstock purification 3, typically for hydrodesulfurization. For such recycling, typically 0–5%, and especially 0.5–3%, of the generated hydrogen product is used, depending on the quality of the feedstock and whether another hydrogen source is used.

[0100] By providing a membrane separation system 80 that is selective for hydrogen permeation, for example, it becomes possible to recover unconverted carbon monoxide and methane in the hydrocarbon-rich concentrate logistics 81. The hydrocarbon-rich concentrate from the hydrogen permeation membrane separation system may still contain hydrogen gas. It may further contain residual unconverted hydrocarbons from the feedstock, particularly methane, and carbon monoxide formed in each of the reformer systems or shift reaction zones. Therefore, all or part of the hydrocarbon-rich concentrate 81 may be recycled from the hydrogen permeation membrane separation system 80 and combined with one or more of the hydrocarbon feedstock 1 upstream of the feedstock refined product 10 (e.g., via passage or flow 83), feedstock to the pre-reformer 20 (e.g., via passage or flow 84), feedstock to the reformer 30 (e.g., via passage or flow 85), feedstock to the heat exchanger reformer 35 (e.g., via passage or flow 86), or reformed product 32 / 36 before introduction into the shift reactor zone 50 (e.g., via passage or flow 87). This is schematically shown in Figures 1, 2, and 4, but can be similarly applied to Figure 3.

[0101] Recycling of hydrocarbon-enriched concentrates under reaction conditions in reformer units (20, 30, 35) reduces greenhouse gas emissions. When exhaust gas containing CO and / or hydrocarbons (particularly methane) is supplied to the radiating section for combustion, the exhaust gas produces CO2 that is typically released into the atmosphere, whereas CO2 generated from recycled exhaust gas can be captured and stored in a carbon dioxide capture unit or used for further purposes. The tail gas is compressed in compressor 70 because the pressure of the tail gas is usually considerably lower than the pressure of the flow it is combined with, for example, more than 10 times lower. For example, the tail gas pressure from the PSA may be approximately atmospheric pressure or slightly higher, for example, about 0.13 MPa, compared to the supply pressure, which ranges from about 2.5 to about 4.0 MPa.

[0102] A small portion of the waste gas stream 63 (tail gas) from the hydrogen and carbon dioxide recovery unit 60 may need to be purged if inert gases (particularly nitrogen) accumulate for recycling. This is achieved by supplying a portion of the tail gas to the furnace 90, where the combustible components function as fuel and the inert gases are purged. Typically, in this embodiment, less than 10% of the tail gas needs to be purged (supplied to the furnace 90 via line 64). The maximum amount that can be favorably recycled generally depends on when unacceptable gas accumulation occurs. If no such accumulation is present, as already mentioned, the fuel demand of the furnace can be met by a portion of the hydrogen product gas 62 obtained in the hydrogen and carbon dioxide recovery unit 60, so essentially all of the tail gas 63 can be favorably recycled.

[0103] The present invention, which applies a hydrogen permeable membrane separation system 80 to tail gas 63, has a particularly favorable effect in further reducing carbon dioxide emissions / footprint, but also reduces reformer supply requirements due to recycled hydrocarbon molecules to the reformer reaction unit. Therefore, in the process according to the present invention, the proportion of hydrogen and tail gas 63 from the carbon dioxide recovery unit 60 to be recycled can be any proportion from 0 to 100%, and the optimal value can be selected depending on an evaluation of the required / desired CO2 emission reduction, the required / desired overall CO2 footprint reduction, and the desired or acceptable size of the reformer system (or the entire plant). The benefit of combining heat exchanger reforming with tail gas recycling in terms of CO2 emission reduction increases as the recycling flow rate increases. Therefore, for further beneficial effects on CO2 emissions, tail gas recycling of at least about 10%, preferably at least 25%, more preferably at least 40%, and especially at least 50% can be used. Depending on the acceptable requirements of the reformer system, tail gas recycling may be up to 100%, up to 80%, up to 60%, up to 40%, or up to 20%.

[0104] By including a membrane separation system as introduced by the present invention, the wind-up effect of hydrogen combustion on total fuel demand is avoided. Additional combustion using hydrogen products requires more hydrocarbon feed to generate hydrogen fuel, which requires more combustion and leads to increased tail gas recycling, but increased combustion again requires more feed, leading to an increase in plant size. Instead, the present invention extracts hydrogen to be used as fuel from tail gas recycling and recycles unconverted carbon molecules back into the reformer section as hydrocarbon feed, thereby enabling a significant net hydrocarbon conversion to hydrogen-containing products 62 that can be further extracted from the process. The corresponding reduction in net hydrocarbon feed reduces combustion requirements, which reduces the amount of fuel hydrogen generated by the plant, and further reduces plant size and combustion demand. The hydrogen-rich permeate flow 82 reduces the tail gas recycling flow through the reformer unit and therefore reduces the required heat input in the reformer section. Compared to the flow scheme of this technology using tail gas recycling 63 without a membrane separation system, the extraction of hydrogen from the tail gas recycling flow 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 the furnace 90. This reduces the amount of crude or pure hydrogen replenishment required.

[0105] The advantages of heat exchanger reformers for feed conversion efficiency and steam generation have already been described above, and the significant reduction in combustion achieved by internal heat recovery further reduces the need for further hydrogen fuel generation and its associated waste gas recycling when hydrogen products are used as fuel. This further reduces the absolute amount of hydrocarbon feed and tail gas compressed compared to technologies that do not utilize heat exchanger reforming. Heat exchanger reforming thereby substantially reduces the direct and indirect CO2 emissions of the plant.

[0106] Combining the advantages of heat exchanger reforming with the advantages of hydrogen fuel extracted from tail gas recycling by membrane separation systems, a hydrogen plant that extracts CO2 generated from process gas (reformers and / or shift reactor products) becomes a highly efficient, low-carbon footprint process, as their combined effects outweigh the benefits.

[0107] The concentrate recycling 81 results in even lower combustion requirements. This combined effect leads to an unexpectedly significant reduction in hydrocarbon feedstock consumption compared to the same process or equipment without a heat exchanger reformer.

[0108] Therefore, in a particularly preferred embodiment of the present invention, the combination of the reformer design and arrangement, the recycling of tail gas 63 from the heat exchanger reformer reaction unit, CO2 removal, and hydrogen and carbon dioxide recovery unit 60, and the hydrogen permeate membrane separation system 80 in the tail gas, allows for the extraction of most of the hydrogen from the tail gas as hydrogen-rich permeate 82 used as fuel for the heating furnace 90, and the recycling of hydrocarbon-rich concentrated logistics 81, thereby reducing CO2 emissions from the reformer plant, particularly ATR-based reformer plants, by approximately 99% or more, and in particular, minimizing the consumption of hydrocarbon feedstock (for fuel generation). In at least a specific embodiment, a reduction of more than 99.9% is achievable without requiring a purge flow from the tail gas, resulting in a plant with nearly net-zero CO2 emissions.

[0109] As can be seen from the above, hydrogen that is used as fuel or recycled to be combined with hydrocarbon feedstock is part of the hydrogen-containing products generated in the process according to the present invention, and can advantageously be obtained from the hydrogen and carbon dioxide recovery unit 60. However, it is also possible to install an (additional) hydrogen recovery unit upstream of the hydrogen production plant. Advantageously, the (additional) hydrogen recovery unit is configured to recover hydrogen-enriched gas from the reformer 36 and is located between the reformer outlet of the reformer system and the feed inlet of the shift reactor zone 50. This has the additional effect of reducing the hydrogen concentration in the reformer, which helps to draw the shift reactor toward hydrogen production, and thus can further reduce hydrocarbon consumption, maximize CO2 capture, and ultimately result in lower CO2 emissions.

[0110] Such effects can also be achieved by positioning a hydrogen recovery unit between two shift reactor sections in a shift reactor zone that includes two or more shift reactor units. A preferred embodiment of such a hydrogen recovery unit may be a hydrogen selective membrane utilizing palladium.

[0111] Another option is to provide (additional) recovery downstream of the shift reactor zone 50, which can be located upstream or downstream of the carbon dioxide recovery unit.

[0112] When hydrogen-enriched gas is used as fuel or recycled, it can generally have a lower hydrogen concentration in the final hydrogen product than desired. Therefore, lower strictness recovery conditions that result in lower hydrogen purity may suffice. Hydrogen-selective membrane separators are particularly suitable. The hydrogen product stream from the methane unit can also be used as a low-carbon fuel.

[0113] Next, the present invention will be described with reference to the following examples. [Examples]

[0114] Table 1 lists several key performance parameters for evaluating different flow scheme options for hydrogen plant capacity based on ATR reforming technology of high-purity hydrogen (99.9% purity) at approximately 9000 kg / h. The table utilizes five different designs to illustrate the potential advantages of the present invention when applied to ATR-based hydrogen production schemes.

[0115] The figures in the table, expressed as percentages, are always relative to the current level of technology represented by Case 1. Only the rows for Supply Efficiency (HHV) and CO2 Capture Rate are exceptions to this interpretation. Supply efficiency is defined as the calorific value of hydrogen products relative to the calorific value of hydrocarbon feedstock (excluding power intake), and CO2 capture rate is defined as the proportion of carbon in CO2 products discharged from the plant to the total carbon entering the plant with the hydrocarbon feedstock and / or fuel flow. The compression of CO2 products to 3 MPa is a substantial contribution to the plant's total power consumption and is listed as a separate contribution to the total for each case, as the contribution may be higher when compressed at higher outlet pressures. It can also be used as a relative measure of total CO2 capture and generation for each scheme. Applying the concepts introduced by this invention demonstrates that, despite more CO2 being captured from the feedstock, the power required for CO2 compression is reduced, and total CO2 generation from the plant is also reduced. The benchmark defining 100% for all columns is Case 1. In the case of electricity consumption, these all refer to 100% of the electricity consumption of a state-of-the-art plant represented by Case 1, while CO2 compression in Case 4 consumes 20.6% of the total electricity consumption of Case 1.

[0116] The carbon intensity of direct emissions (range 1) and indirect emissions (range 2) from power generation is expressed in kg of CO2 per kg of H2 produced. A typical value for a standard hydrogen plant without CO2 capture is around 10 kg CO2 / kg H2, while the target value for state-of-the-art plants ranges from 0.5 to 1.5 kg CO2 / kg H2, depending on the carbon intensity of the electricity used. The table shows range 1 + range 2 values ​​for three different locations: Location A has a high contribution from renewable energy (50 g CO2eq / kWh, e.g., Norway, Sweden); Location B has only a moderate contribution from renewable energy (250 g CO2eq / kWh, e.g., Australia, Poland, UK); and Location C has a low contribution from renewable energy (380 g CO2eq / kWh, Germany, USA). Higher oxygen consumption usually leads to higher electricity consumption, while the remaining electricity consumers remain at a similar size. Renewable energy should be interpreted here as power sources with low associated CO2 emissions, such as solar photovoltaic (PV), wind turbines, hydroelectric power, and / or nuclear power.

[0117] [Table 1]

[0118] Table 1 - Performance Evaluation Cases 1 and 2 describe current state-of-the-art units optimized for 95% direct CO2 emission capture, while Cases 3, 4, and 5 describe flow schemes according to the present invention that achieve 99.7%, 98.2%, and 99.7% direct CO2 emission capture, respectively. Achieving a 99.7% capture rate in current state-of-the-art plants significantly increases utility consumption and capital expenditures. Conversely, the application of the present invention reduces hydrocarbon feed consumption, oxygen consumption, and overall plant power consumption, resulting in significant operational benefits.

[0119] The carbon intensity values ​​at location C demonstrate that while the application of the present invention still results in a reduction of total direct and indirect CO2 emissions, the impact is far less compared to locations A and B, where the carbon intensity of electricity is much lower. The oxygen generation required for the ATR process is highly energy-intensive and, as is known to those skilled in the art, is hardly compensated for by the undesirable reduction of CO2 at location C. At locations where electricity is inexpensive and / or has a low carbon intensity, the contribution of electricity to the total carbon intensity is (much) lower, and therefore, the hydrocarbon conversion efficiency and CO2 capture rate from the process are the dominant parameters for obtaining a low carbon intensity overall. This can be seen in the results for cases 3, 4, and 5.

[0120] Case 1 is a hydrogen plant following an ATR-based reforming flow scheme, comprising a feed purification section 10, a pre-reformer 20 and an ATR reformer 30, a heat recovery section 40 with a convection section 100, a water-gas shift section 50, a hydrogen and carbon dioxide recovery system 60, and a furnace 90, where all waste gas 63 is used as fuel 64 to the furnace 90, and neither waste gas recycling 65 nor a membrane separation system 80 is applied. Since there is excess heat available in the process, it is necessary to generate excess steam which is sent out to generate electricity in a steam turbine 47. This reduces the net power intake required by the plant but increases the consumption of feed, thereby increasing the amount of CO2 generated during conversion. Since unconverted carbon from the waste gas is neither captured nor recycled, CO2 stack emissions in the flue gas 101 are high, resulting in CO2 capture of slightly more than 95% of the CO2 from the feed hydrocarbons. The captured CO2 is compressed to 30 bar g and sent out to the plant battery limit for further processing.

[0121] Case 2 describes the same plant with the addition of a heat exchanger reformer 35 (Figure 2) arranged in parallel. Neither the waste gas recycling 65 nor the membrane separation system 80 are included. Excess steam is sent out for power generation, but as the flow rate decreases, less power is generated inside the turbine, and therefore the net intake power increases accordingly. Total oxygen consumption decreases significantly, which reduces the total power consumed by the plant.

[0122] Case 3 (according to the present invention) applies hydrocarbon-rich recycled material 81 from the membrane separation system 80 to the furnace 90 together with hydrogen-rich permeate flow 82 used as fuel, as described in the present invention. A heat exchanger reformer 35 is also applied. Crude hydrogen 68 is also supplied to the membrane separation system 80 to further decarburize the fuel used in the furnace 80. A steam turbine 47 is still included for power recovery from the generated process steam. In this case, the steam is generated at a pressure higher than required for the process, and power recovery occurs in expansion up to the process pressure. The CO2 capture rate from the feed is increased to 99.7%, while the consumption of hydrocarbon feed 1 and oxygen 31 is further reduced, and the total power consumption of the plant is also reduced. By applying the hydrogen-rich flow 81 from the membrane separation system 80 to the furnace as a low-carbon fuel, CO2 emissions from the plant stack are significantly reduced. By recycling carbon molecules contained in the hydrocarbon-rich flow 81 extracted from the membrane separation system 80 and recycling them to the reforming section via the flows (84, 85, 86) as feed hydrocarbons, the amount of hydrocarbon feed 1 required for the same total hydrogen production is further reduced, thereby improving plant energy efficiency and further reducing CO2 emissions from the hydrocarbon feed. The corresponding reduction in power consumption further reduces the carbon intensity of the plant. In areas where fossil fuels contribute a high amount to power generation (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 slightly above 1 kg CO2 emitted / kg H2 produced. In areas where renewable energy contributes a high amount to the power mix (50 g CO2 eq / kWh), this scheme allows for a carbon intensity of 0.17 kg CO2 per kg of H2 produced. Case 5 is identical to this scheme except for the power generation block, which can be a cost-driven decision. In Case 5, there is a slight decrease in hydrocarbon consumption compared to Case 3, but the carbon intensity is higher in cases with a high carbon footprint associated with electricity (B and C).

[0123] Case 4 (according to the present invention) is a variation of Case 3, in which a portion of the hydrocarbon feed 1 is used directly in the heating furnace 90, replacing the amount of heat provided by feeding a portion of the crude hydrogen flow 68. This increases the carbon content of the flow 92, as demonstrated by the corresponding decrease in the CO2 capture rate from the feed to 98.2% and the associated increase in carbon intensity, but also yields a slight benefit in hydrocarbon feed consumption, which further reduces power consumption by 1.3% and 2.8%, respectively. This variation of the flow scheme demonstrates that those skilled in the art have the flexibility to design the plant for specific CO2 capture rate or carbon intensity requirements without compromising the plant's energy efficiency through over-design.

[0124] Case 5 (according to the present invention) is another variation of Case 3 in which the steam turbine (47) is absent. Since there is no internal power generation in this flow scheme, the net taken-in power consumption is higher than in Case 3. Where the carbon intensity of the power is high, this negatively impacts the carbon intensity of the plant in range 1 + range 2, but where the carbon intensity of the power is low, the same overall carbon intensity as a plant with internal power generation can be achieved. A comparison of the performance diagrams of Case 3 and Case 5 further demonstrates that steam generation consumes additional hydrocarbons and oxygen to provide the hydrogen and heat required for steam generation, but still reduces overall CO2 emissions and carbon intensity in regions with high carbon intensity for power, and that the reduction in carbon intensity achievable by a plant designed according to the present invention can be further improved by adding its own power generation to reduce net power import in regions where power has a low renewable contribution.

[0125] [Table 2]

Claims

1. It is a hydrogen generation plant, - At least one reformer (30) for converting a flow containing hydrocarbon feedstock (1) into a reformed gas flow (32) containing hydrogen, carbon monoxide, carbon dioxide, and at least one hydrocarbon as an impurity via conversion with oxygen-rich vapor, wherein the reformer (30) includes an exothermic oxygen-based autothermal (ATR) or partial oxidation (POX) reforming and a heat recovery section (40), - A heating furnace (90) configured to preheat the flow containing the hydrocarbon feed material (1) before it enters at least one reformer (30), - At least one water-gas shift (WGS) reactor (50) for converting the carbon monoxide in 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 flow (51) and generate a first production flow (61) enriched with carbon dioxide, a second production flow (62) enriched with hydrogen, and a waste flow (63) in which both hydrogen and carbon dioxide are depleted, - A compressor (70) for compressing a portion (65) of the waste gas flow (63) from the hydrogen and carbon dioxide recovery unit (60) into a compressed gas flow (71), - A membrane separation system (80) selective for hydrogen permeation, configured to supply the compressed gas flow (71) and generate a hydrogen-enriched permeate (82) flow and a hydrocarbon-enriched concentrate (81) flow, - A passage for supplying at least a portion of the hydrogen-enriched permeate (82) to the heating furnace (90) so that it can be used as a low-carbon fuel by the heating furnace (90), A hydrogen production plant comprising: a passage for recycling the hydrocarbon enriched concentrate (81) to a hydrocarbon feed (1) via a pipeline (83), and / or to the reformer (30) via the pipeline (85), and / or to the inlet of the water-gas shift reactor (50) via the pipeline (87).

2. The hydrogen production plant according to claim 1, wherein the plant includes a passage for supplying a portion (64) of the exhaust gas flow (63) from the hydrogen and carbon dioxide recovery unit (60) to the heating furnace (90).

3. The hydrogen production plant according to claim 1 or 2, wherein the hydrogen and carbon dioxide recovery unit (60) is configured to generate a flash gas flow (69) and / or a carbon dioxide-depleted and hydrogen-rich flow (68).

4. The heating furnace (90) uses its fuel, a) At least a portion (64) of the exhaust gas flow (63) from the hydrogen and carbon dioxide recovery unit (60), b) At least a portion of the hydrogen-enriched permeate (82) produced by the hydrogen permeate membrane separation system (80), c) At least a portion of the hydrogen enrichment product (62) from the hydrogen and carbon dioxide recovery unit (60), d) At least a portion of the hydrocarbon supply raw material flow (1), e) Resupply fuel flow taken in from the battery limit, f) At least a portion of the flash gas flow (69), and g) A hydrogen production plant according to claim 3, which receives hydrogen from at least one of at least a portion of a flow that is depleted of carbon dioxide and rich in hydrogen (68).

5. The heat recovery section (40) is configured to generate a steam flow (41), and the plant uses at least a portion of this steam flow (41) a) Process steam (45) to the inlet of the reformer (30), and / or b) Process steam (43) is supplied to the inlet of the water-gas shift reactor (50), c) A hydrogen production plant according to any one of claims 1 to 4, comprising means for routing the discharged steam (46) to the battery limit.

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

7. The heat recovery section (40) is configured to generate a steam flow (41), and the hydrogen plant - At least one pre-reforming reactor (20) located upstream of the reformer (30) and configured to generate a pre-reformed synthesis gas flow (21) via means (84) from at least one or both of the treated hydrocarbon flow (11) and the hydrocarbon-rich flow (81), - A means for routing at least a portion (44) of the steam flow (41) to the inlet of the pre-reformer (20), as described in claim 6.

8. The heat recovery section (40) is configured to generate a steam flow (41), and the plant is configured - A heat exchanger reformer (35) installed in series with the reformer (30) and receiving at least one of the hydrocarbon feed (11), pre-reformer feed (23), or a portion of the hydrocarbon-rich flow (81) provided via means (86), which is mixed with steam (42) to the tube-side inlet, to generate a reformed flow (37) from the tube-side outlet, - A means for supplying this reformed flow (37) to the reformer (30) to generate a reformed flow (32), wherein the reformed flow (32) enters the shell-side inlet of the heat exchanger reformer (35), provides reaction heat to the tube side of the heat exchanger reformer (35), and generates a reformed flow (36) from the shell-side outlet of the heat exchanger reformer (35), - Means for sending the reformed flow (36) to the heat recovery section (40), - A hydrogen production plant according to claim 6 or 7, comprising means for mixing at least a portion (42) of the steam flow (41) into the tube inlet of the heat exchanger reformer (35).

9. - The heat recovery section (40) is configured to generate a steam flow (41), - The plant is equipped with means for dividing the reformer supply flow (21) into a first part (23) and a second part (22), - The reformer (30) is configured to receive the second portion (22) and generate the reformed gas flow (32), - The plant comprises a heat exchanger reformer (35) installed in parallel with the reformer (30) and configured to receive at least one or both of a portion (23) of the reformer supply flow (21) and at least a portion of the hydrocarbon-rich flow (81) via means (86), and the reformed gas flow (32) to generate a reformed flow (36); means for mixing the reformed supply gas flow (32) to 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) either inside or outside the heat exchanger reformer (35); and means for mixing at least a portion (42) of the vapor flow (41) to the tube inlet of the heat exchanger reformer (35), - The hydrogen production plant according to claim 6 or 7, wherein the heat recovery section (40) is configured to receive reformed logistics (36).

10. A hydrogen production plant according to claim 8 or 9, comprising a plurality of heat exchanger reformers (35).

11. The hydrogen production plant according to any one of claims 1 to 10, wherein the heat recovery section (40) is configured to generate a steam flow (41), and the plant comprises a steam turbine (47) configured to receive at least a portion of the steam flow (41) and generate a low-pressure flow (48).

12. The hydrogen production plant according to any one of claims 1 to 11, wherein the membrane separation system (80) comprises membrane elements based on polysulfone, polyimide, polyaramid, cellulose acetate, any combination thereof, or other polymer materials, or palladium sheets, which exhibit selectivity to preferentially permeate hydrogen to lower pressures.

13. The hydrogen production plant according to any one of claims 1 to 12, wherein the plant provides a passage for introducing an off-gas (72) containing hydrogen and hydrocarbons from the battery limit to the membrane separation system (80) as a feed, and recovering hydrogen from the off-gas into the hydrogen-enriched permeate (82) and hydrocarbons into the hydrocarbon-enriched concentrate (81).

14. A hydrogen production plant according to any one of claims 1 to 13, wherein the plant comprises a passage for routing at least a portion of the hydrogen-rich permeable logistics (82) to the heating furnace (90) as fuel, and means for recycling the remaining portion of the hydrogen-rich permeable logistics (82) to the inlet of the hydrogen recovery unit in the hydrogen and carbon dioxide recovery unit (60).

15. A hydrogen production plant according to any one of claims 1 to 14, wherein at least a portion (88) of the hydrogen-rich permeate flow (82) from the membrane separation system (80) is sent as a product flow to the plant battery limit.

16. The hydrogen production plant according to any one of claims 1 to 15, further comprising a hydrogen recovery unit positioned between the reformer (30) and the water-gas shift reactor (50), wherein the hydrogen recovery unit is configured to remove hydrogen from the reformed gas flow (32) before it enters the water-gas shift reactor (50).

17. The hydrogen production plant according to any one of claims 1 to 16, wherein the heating furnace (90) is provided as a component separate from the at least one reformer (30).

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