Reactor for low-temperature methane splitting and solid catalyst to use therein and methods thereof

Abstract

The present disclosure relates to the field of catalytic reactors, specifically to a low-temperature catalytic methane splitting reactor operating between 500°C and 650°C. The reactor is designed for the continuous production of high-purity hydrogen and graphitic carbon, with no CO₂ emissions, using a novel combination of supported catalysts, reaction module, and a membrane-based gas separation system.

Description

D E S C R I P T I O NREACTOR FOR LOW-TEMPERATURE METHANE SPLITTING AND SOLID CATALYST TO USE THEREIN AND METHODS THEREOFTechnical field

[0001] The present disclosure relates to the field of the gas to solid catalytic reactors, specifically to a low-temperature catalytic methane splitting operating between 500°C and 650°C. The reactor is for the continuous production of high-purity hydrogen and graphitic carbon, with no CO2emissions, using a novel combination of supported catalysts, macro-support structures, and a membrane-based gas separation system. The present disclosure relates to a low-temperature catalytic methane splitting reactor, operating from 500 °C to 650 °C, for producing high-purity hydrogen (H2) and graphitic carbon.Background

[0002] Currently, most of the hydrogen produced worldwide comes from the reforming of natural gas and other hydrocarbons or coal gasification. These processes are gradually becoming environmentally and economically less appealing, due to extensive CO2 emissions. Among many alternatives, methane splitting, also known as methane decomposition, cracking, or pyrolysis, has particular potential to uproot the current status quo:CH4(g) 2H2(g) + C (s), AH0= 74.8 kJ mol1

[0003] Catalytic methane splitting can seamlessly substitute reforming reactors in hydrogen production facilities, thanks to its similarities with hydrocarbon reforming processes. Methane splitting avoids any production of greenhouse gases, making it environmentally and economically more appealing.

[0004] Methane splitting systems can be operated at a variety of conditions, using a variety of catalytic materials. To work at mild temperatures, <650 °C, however, supported metal nanostructured catalysts are widely accepted as the only effective pathway.

[0005] Low-temperature catalytic methane splitting presents some issues that have slowed down its increase in economic relevance, namely: i) fast catalyst deactivation due to carbon growth over active sites; ii) catalyst destruction due to active metal detaching from the support; iii) reactor clogging from uncontrolled carbon growth; and iv) methane conversion limitations per reactor pass, owing to the establishment of an equilibrium for the reversible methane splitting reaction.

[0006] Deactivation due to carbon deposition over catalyst active sites has been heavily tackled in the past. Various regeneration methods are reported for dealing with such deactivation, by removing the formed carbon allotropes from the catalyst active sites. Removing the carbon from the active sites makes the latter again available to adsorb methane and perform the reaction. This method is known as catalyst regeneration.

[0007] The document US20070111051A1 [1] reports a method in which the carbon is removed by combustion, outside of the reactor, effectively regenerating the catalytic material. Carbon combustion- or gasification-based methods effectively eliminate carbon deposits from the catalyst sites, but they are fundamentally flawed. Carbon dioxide (CO2) is produced in such processes, denying the environmental advantages of catalytic methane splitting. Furthermore, the carbon product is a valuable reaction product, being their oxidation detrimental for its valorization and thus the overall process economics.

[0008] The document WO2020121287A1 [2] reports another method for removing carbon deposits, specifically from carbon grown in metallic catalysts. Here a continuous cyclic regeneration method is described, in which a hydrogen stream is periodically fed into the reactor, causing carbon detachment from the catalyst surface. This detachment is promoted by a selective interfacial hydrogenative gasification of carbon (reverse of methane splitting):C (s) + 2H2(g) CH4(g), AH0= -74.8 kJ mol1

[0009] Hydrogen becomes preferentially activated on the surface sites of the metal, hence driving the gasification of carbon deposits at the interface with the metal particles, provoking the carbon to detach from the catalytic metal surface as particles and restoring the availability of the latter for methane splitting catalysis. Carbon further away from the catalyst remains effectively unreacted. As such, carbon products can be mostly recovered, without any production of gaseous by-products.

[0010] The current background art provides technical solutions for catalyst deactivation by carbon deposition, as the ones outlined in the documents cited above for reference. However, other technical aspects, which hamper a sustained operation of low-temperature catalytic methane splitting for long times, remain as technical hurdles. Notably, an effective catalyst online (cyclic) regeneration of the catalyst can only be effective if the catalyst is not irreversibly physically modified / destroyed by the carbon growth.

[0011] Two different carbon-growth mechanisms are known to those skilled in the technique for the methane splitting reaction proceeding on metal catalysts: i) base-growth mechanism; and ii) tip-growth mechanism.

[0012] The base-growth mechanism is characterized by carbon structures nucleating on the surface of static metal particles, gradually covering its surface, and generally forming tubular structures that grow away from the metal particle. In this case, the end of the carbon structure opposite to the static metal particle, wherefrom it grows, is free from metal.

[0013] The tip-growth mechanism is characterized by carbon nucleation at the interface between metal particles and the support on which they are dispersed, generally forming filamentous allotropes such as multi-walled carbon nanotubes or carbon nanofibers. In this case, the carbon structures entrain metal particles during their growth from the surface of the catalyst support material. Therefore, metal particlesare effectively lifted from the support bringing about a permanent physical disruption of the catalyst structure and disabling the reversion of catalyst deactivation. Considering the technical implication of these two different carbon growth mechanisms for catalyst deactivation and online regenerability, the tip-growth mechanism should be suppressed to enable the quasi-continuous (cyclic) regeneration of the catalyst and, therefore, long-term operation of a low-temperature methane splitting process. The current background art does not report the effective suppression of the tip-growth mechanism under low- temperature methane splitting reaction conditions.

[0014] Reactor clogging due to carbon build-up is another under-addressed technical challenge, despite its eminent significance in ensuring long-term continuous process operation.

[0015] The document W02013004398A8 [3] explores the movement of carbon away from the reaction zone, by considering a mobile amorphous carbon bed, over which the methane-derived carbon grows. In this approach, a carbon stream is continuously fed to the reactor, which is then continuously removed from the reactor, together with the carbon product that has been deposited in the course of the reactor pass.

[0016] The document WO2016 / 154666A1 [4] considers a similar system, in which iron ore is used as cheap catalytic material, being continuously fed and removed from the system, along with the carbon allotropes that form over it. These approaches solve issues associated with clogging of flow paths with carbon. However, this is achieved by complex reactor designs, including moving parts and the continuous discharge, and disposal, of catalytic materials that provide low to no hydrogen productivity in a process of low-temperature catalytic methane splitting at reaction temperatures <650 °C.

[0017] Current background art still lacks an effective low-temperature methane splitting system which provides economically viable hydrogen productivity, while avoiding clogging of flow paths by solid carbon products.General Description

[0018] New and improved reactor designs should be considered to avoid clogging of flow paths by carbon, while attaining high hydrogen productivity, enabling the retention of the catalyst in the reactor and providing means for a cyclic catalyst regeneration, which enables a sustained operation over extended period. Furthermore, equilibrium constraints at low temperatures should also be considered. It is important to consider integrated systems (or reactors), such as membrane reactors, which can increase the apparent conversion of methane splitting systems (or reactors) above the equilibrium conversion.

[0019] The present disclosure relates to a catalytic methane-splitting reactor comprising two compartments:a first compartment containing a catalyst-loaded macro-support where the methane splitting reaction occurs. a second compartment featuring a gas separation membrane for extracting high-purity hydrogen.

[0020] In the present disclosure, it is defined that macro-support is a substrate of the inner surface of the channels.

[0021] In an embodiment, the catalyst is a supported metal catalyst consisting of nickel (Ni) or cobalt (Co) thereof, loaded into the mesopores of ceramic support, preferably mesoporous silica support, with a diameter of ca. 20 nm. The catalyst is cyclically regenerated using hydrogen, which selectively removes the carbon at the catalyst / carbon interface, enabling sustained operation over long periods. The structure of the reactor of the present disclosure ensures stable and continuous production of hydrogen and graphitic carbon without CO2emissions.

[0022] In one embodiment, the catalyst may comprise Ni-based nanoparticles enclosed within the pores of ceramic supports. This design prevents carbon adatom diffusion to the metal-support interface, suppressing destructive tip-growth mechanisms. Optionally, Ga doping is used to enhance stability and performance.

[0023] In another embodiment, the catalyst may comprise Co-based nanoparticles, specifically tuned for low-temperature base-growth of carbon allotropes, are used. These fine Co nanoparticles are also enclosed in ceramic support pores, ensuring robust catalyst performance at the specified operating temperatures.

[0024] In another embodiment, the reactor utilizes innovative macro-support structures such as ceramic monoliths, plate-and-frame configurations, honeycombs, or metal tubes. The supported catalyst is deposited on the inner surfaces of these tubular or honeycomb structures. Novel methods, including solgel or thermosetting polymer films, bind the supported catalyst to the macro-support surface.

[0025] In another embodiment, the gas separation membrane reactor is divided into two compartments with internal recirculation. The product stream from the reaction compartment flows to the membrane module, where hydrogen is partially separated and returned to the reaction chamber.

[0026] The methane splitting reaction displays a 100 % selectivity for hydrogen production under the specified conditions, producing only hydrogen and carbon. Methane is fed into the reactor, purified hydrogen exits the system, and carbon is periodically removed.

[0027] In another embodiment, the use of Ni- or Co-based nanoparticles enclosed in ceramic supports for enhanced catalyst stability.

[0028] In another embodiment, the Ni-based nanoparticles can be underneath coated with a Ga nanolayer to prevent destructive carbon deposition mechanisms.

[0029] In another embodiment, the sol-gel or thermosetting polymer films allow robust catalyst binding to macro-supports.

[0030] A sol-gel composition comprises a mixture of one or more metal alkoxide precursors, such as tetraethyl orthosilicate (TEOS) or titanium isopropoxide, in a solvent medium, typically an alcohol like ethanol or methanol, combined with a hydrolysis agent such as water and an optional catalyst, which may be acidic or basic (e.g., hydrochloric acid or ammonium hydroxide). This composition undergoes hydrolysis and condensation reactions to form a gel-like network, which can be processed into a solid material with tailored properties. The sol-gel composition may further include additives such as surfactants, dispersants, or functional dopants to modify characteristics like porosity, adhesion, flexibility, or thermal resistance. The ratio of components and reaction conditions, such as temperature and pH, are precisely controlled to produce a material optimized for specific applications, including binders, coatings, or structural films.

[0031] The present disclosure relates to a reactor for production of hydrogen and carbon from methane splitting, comprising a gas inlet for methane flow, a reaction chamber for methane splitting, a gas separation membrane for separating and purifying hydrogen, a gas outlet for hydrogen flow, a bottom outlet for carbon removing by gravity from the reaction chamber; wherein the reactor comprises a pipe connecting the reaction chamber to the gas separation membrane to flow product reaction for introducing the reaction products of the reaction module into the plurality of gas separation membrane, wherein the reaction chamber comprises a reaction module; wherein the reaction module comprises a plurality of gas flow channels with a macro-support; wherein the macro-support comprises a mesoporous support loaded with a metal catalyst for contacting with the methane flow for catalyzing the methane splitting reaction; wherein the said metal catalyst is a plurality of Ni or Co alloy particles.

[0032] The reaction module can be a reaction chamber.

[0033] In an embodiment, the operation temperature of the reaction module ranges from about 500 °C to 650 °C; preferably from about 550 °C to 600 °C.

[0034] In an embodiment, the reaction module has an operation pressure and the inlet pressure of the gas separation membrane ranges from about 1 xlO5Pa to 10xl05Pa; preferably from about 1 x 105Pa to 5 x 105Pa.

[0035] In an embodiment, the gas separation membrane has an outlet pressure ranging from about 1 x 105Pa - 5 x 105Pa; preferably from about 1 x 105Pa - 2 x 105Pa.

[0036] In an embodiment, transmembrane hydrogen partial pressure differences in the gas separation membranes range from 0.1 x 105Pa - 0.001 x 105Pa; preferably from 0.05 x 105 Pa - 0.01 x 105Pa.

[0037] In an embodiment, the geometry of the reaction module is selected from a hexagonal shape, a honeycomb shape, a tubular shape, a cubical shape or a parallelepipedal shape.

[0038] In an embodiment, the geometry of the plurality of gas flow channels are selected from: tubular, hexagonal or plate stack.

[0039] In an embodiment, the plurality of gas flow channels comprises at least a heating device for controlling the temperature reaction, preferably an electrical heating resistance or a thermal fluid.

[0040] In an embodiment, the heating device is arranged in a tube or a plate.

[0041] In an embodiment, the reactor comprises a temperature measuring device arranged between the plurality of gas flow channels, preferably a thermocouple.

[0042] In an embodiment, the plurality of gas flow channels has an inner equivalent diameter from 0.25 cm to 2 cm; more preferably from 0.5 cm to 1 cm.

[0043] In an embodiment, each channel of the plurality of gas flow channels traverses the entire module, with a preferential length of the tubular channels from 1 cm to 20 m; more preferably from 30 cm to 3 m.

[0044] In an embodiment, the plurality of nickel or cobalt alloy catalysts particles is alloyed with an element of the list consisting of: gallium, indium, germanium, or combination thereof.

[0045] In an embodiment, the reactor further comprising alumina ceramic, or a silica ceramic, or a thermosetting polymer for bounding the plurality of nickel or cobalt alloy catalysts to the macro-support; preferably the catalyst is embedded in the alumina ceramic or a silica ceramic and / or thermosetting polymer.

[0046] In an embodiment, the catalytic support comprises mesopores ranging from 10 nm to 50 nm; preferably from 10 nm to 30 nm; more preferably from 15 nm to 25 nm. Measurement of the mesopore size can be carried out by transmission electron microscopy (TEM), atomic force microscopy (AFM), X-ray diffraction (XRD), among others.

[0047] In an embodiment, the size of each particle of the plurality of nickel or cobalt alloy catalysts particles ranges from 8 nm to 40 nm; preferably from 10 nm to 20 nm; more preferably from 15 nm to 20 nm. Measurement of the size can be carried out by transmission electron microscopy (TEM), atomic force microscopy (AFM), X-ray diffraction (XRD), among others.

[0048] In an embodiment, the distance between consecutive plates ranges from 1.5 mm to 10 mm, more preferably from 2 mm to 4 mm.

[0049] In an embodiment, the depth of the parallelepipedal shaped support is from 5 cm to 10 m; more preferably from 20 cm to 1 m.

[0050] In an embodiment, the mass amount of gallium, indium or germanium in each particle of the plurality of nickel or cobalt alloy catalysts ranges from 0 % to 50 % (wt.eiement / wt.metai particle); preferably from 0.1 % to 30 % (wt.eiement / wt.metai particle); more preferably from 1 % to 15 % (wt.eiement / wt.metai particle), wherein the element is of gallium, indium or germanium and the metal particle is Ni or Co alloy catayst.

[0051] In an embodiment, the mesoporous support has particles ranging from 0.1 pm to 100 pm; preferably from 0.5 pm to 100 pm; more preferably from 5 pm to 25 pm.

[0052] In an embodiment, the metal loading of the reaction module ranges from 5 % to 60 % (w / w); preferably from 10 % to 50 % (w / w); more preferably from 30 %to 50 % (w / w).

[0053] In an embodiment, parallelepipedal shape of the reaction module contains slots between the rectangularly shaped plurality of gas flow channels, for heating device plate placement.

[0054] In an embodiment, the reactor comprises a second gas inlet for introducing a gas regeneration flow, for removing the carbon deposition in the mesoporous support, allowing the regeneration of the plurality of nickel or cobalt alloy catalyst particles.

[0055] In an embodiment, the gas regeneration stream is a hydrogen stream, preferably a pure hydrogen stream.

[0056] In an embodiment, the bottom outlet is cone or funnel shaped.

[0057] In an embodiment, the reactor comprises a container arranged below a bottom of the reactor, for carbon collection.

[0058] In an embodiment, the reactor comprises a third gas inlet to introduce a gas methane flow to the container, for protecting the produced carbon.

[0059] In an embodiment, the support is a metallic support, a ceramic support, a carbon-based support, or a combination thereof.

[0060] In an embodiment, the ceramic support material is selected from a list consisting of: an oxide of silicon, an oxide of aluminium, an oxide of zirconium, an oxide of titanium, an oxide of magnesium, or combinations thereof; preferably an oxide of silicon, an oxide of aluminium, an oxide of zirconium, or combinations thereof; more preferably oxide of silicon, oxide of aluminium or combinations thereof.

[0061] In an embodiment, the pipe comprises a first blower or pump.

[0062] In an embodiment, the gas outlet comprises a second blower or pump to draws hydrogen from the gas separation membranes module to the gas hydrogen outlet.

[0063] In an embodiment, the amount of catalyst in the interior of the plurality of channels ranges from 0.1 to 2 mgcataiyst cm'2, preferably from 0.5 to 1 mgcataiyst cm'2.

[0064] The present disclosure also relates to a process for continuous methane splitting using the reactor described in any of the previous claims, comprising: feeding the reactor with a methane rich flow, wherein the reactor comprises a gas inlet for methane flow, a reaction chamber for methane splitting, a gas separation membrane for separating and purifying hydrogen, a gas outlet for hydrogen flow, a bottom outlet for carbon removing by gravity from the reaction chamber; wherein the reactor comprises a pipe connecting the reaction chamber to the gas separation membranes module to flow product reaction for introducing the reaction products of the reaction module into the plurality of gas separation membrane, wherein the reaction chamber comprises a reaction module; wherein the reaction module comprises a plurality of gas flow channels with a macro-support; wherein the macro-support comprises a catalytic support loaded with a catalyst for contacting with the methane flow for catalyzing the methane splitting reaction; wherein the said catalyst is a plurality of Ni or Co alloy particles, and wherein the operation temperature ranges from about 500 °C to 650 °C; preferably from about 550 °C to 600 °C; and adding a regeneration flow to react at the catalyst interface with deposited carbon, for the catalyst regeneration.

[0065] In an embodiment, the process comprises a further previous step of bonding the catalyst particles to the surface of the support by an alumina ceramic or a silica ceramic material or a carbonized thermosetting polymer.

[0066] In an embodiment, the alumina ceramic or a silica ceramic material or carbonized thermosetting polymer is stable at the operation temperature ranges.

[0067] In an embodiment, the process comprises a further second regeneration step where the reactor is stopped and an inert gas flow is introduced for catalyst regeneration, preferably every 15 days to 60 days, more preferably every 20 days to 30 days.

[0068] In an embodiment, the duration of the second regeneration step ranges from 1 hour to 20 hours; preferably from 4 hours to 8 hours.

[0069] In an embodiment, for running the methane splitting reaction at low temperatures, it is necessary to have a stable, active, and low-cost catalyst. Although the most active catalyst for this reaction is basedon N i, it deactivates very fast, normally remaining active for less than 100 hours on-stream. The present disclosure relates, in the first aspect, to a reactor to perform a low-temperature methane splitting process with high hydrogen productivity, and long-term stability without catalyst replacement. The reactor enables a cyclic, online regeneration of the catalyst by hydrogenative interfacial carbon gasification, as well as quasi-continuous recovery of carbon product without clogging gas flow paths. The disclosure relates, in a second aspect, to a solid catalyst that is highly active for methane splitting at reaction temperatures <650 °C, does not undergo irreversible physical degradation due to carbon growth, and can be integrated in the form of a thin film on a macro-support, which is loaded and sealed in the reactor.

[0070] In an embodiment, the reactor hosts a solid catalyst bound to a macro-support. The solid catalyst is a supported metal catalyst, which comprehends catalytic metals or metal alloys, which comprise Ni, cobalt (Co), gallium (Ga), and any combination thereof, dispersed on a nanostructured oxide support. The reactor comprises two compartments: a first compartment that hosts the catalyst-loaded macro-support, wherein the methane splitting reaction occurs, and a second compartment that hosts a gas separation membrane module. For a stable operation, sustained over long periods, the catalyst is cyclically regenerated using hydrogen, which provokes the selective carbon methanation at the catalyst / carbon interface. The disclosed catalytic membrane reactor allows the stable and continuous production of hydrogen and carbon, without carbon dioxide (CO2) emissions.The following pertains to the reactor

[0071] According to the disclosure, the reactor, is divided into two main compartments. The methane splitting reaction takes place in a first compartment, the reaction compartment, which houses the solid catalyst. The selective recovery of the hydrogen product from unconverted methane takes place in a second compartment, i.e. in the membrane compartment, which houses a module with membranes for selective gas separation. Between the two compartments, a blower or pump draws the gas outlet stream from the reaction compartment to the membrane compartment. Overall, the membrane reactor inputs a methane stream and outputs a continuous pure hydrogen stream. The solid carbon product is cyclically removed, as described below.

[0072] In an embodiment, the transmembrane pressure in the membrane module should be such that it allows hydrogen to selectively permeate through the membranes and be recovered in high purity in the permeate stream, whereas unconverted methane is retained in the retentate stream and is recycled to the reaction module. This means that the partial pressure of hydrogen at the permeate side should be lower than the partial pressure of hydrogen at the retentate side. This transmembrane hydrogen partial pressure difference is preferably from 0.1 x 105Pa to 0.001 x 105Pa (0.1 to 0.001 bar), and more preferably from 0.05 x 105Pa to 0.01 x 105Pa (0.05 to 0.01 bar). The gas pressure at the inlet of the membrane module is from 1 x 105Pa to 10 x 105Pa (1 to 10 bar), and preferably from 1 x 105Pa (1 bar) to 5 x 105Pa (1 to 5 bar).

[0073] According to the disclosure, the retentate stream is driven back to the reaction compartment that operates at a total pressure (absolute pressures) in the range of 1 x 105Pa to 5 x 105Pa (1-5 bar), preferably at a total pressure in the range of 1 x 105Pa to 2 x 105Pa (1-2 bar) and more preferably at atmospheric pressure, since higher reaction pressures result in lower methane conversion values at the thermodynamic equilibrium for the methane splitting reaction.

[0074] In an embodiment, the reaction compartment is designed such that the solid carbon, detached from the catalyst layers during the cyclic catalyst regeneration step, freely fall down by gravity to the bottom of the reaction compartment. To facilitate the release of the carbon from the catalyst as particles, the reactor may be vibrated, either continuously or specifically during the catalyst regeneration step, with the help of an external vibrating unit, preferably an axial vibration unit. The bottom of the reactor is cone- shaped - cf. Figure 4, designed to collect the solid carbon. Periodically, an evacuated carbon collection container is hermetically docked to the bottom of the reactor - cf. Figure 4; and a valve at the bottom of the reaction compartment is opened to allow the carbon particles to flow into the carbon collection container. After this transfer process, the valve at the reactor's bottom is closed again to continue operation.The following pertains to Gas Separation Membranes

[0075] In an embodiment, there are not many membrane technologies that fulfill the requirements for operating in a low-temperature methane splitting reactor. Among them, the palladium-silver membranes display a very high permeability to hydrogen as well as a very high selectivity. However, this technology uses rare and expensive metals that poison easily in the presence of sulfur-based contaminants, and it is only suitable to operate up to 550 °C. A second technology considers the use of carbon molecular sieve membranes, which also display high separation performances such as 1000 barrer and an H2 / CH4 selectivity of 1000 at 550 °C, are thermally stable up to 600 °C, and do not undergo poisoning with sulfur or acid contaminants. Both technologies are suitable for the reactor of the disclosure, wherein the membranes are shaped like tubes.The following pertains to the supported catalyst

[0076] In an embodiment, through the confinement of nickel nanoparticles within the pores of mesoporous ceramic supports, while retaining access to the dispersed metal close to the pore mouths, the formed carbon shows structures characteristic of the base-growth mechanism. The catalytic activity of nickel is too high, and to improve the morphological stability of the supported catalyst, nickel is alloyed with at least a second metal showing a very low intrinsic carbon solubility, which produces stable alloys with nickel, such as gallium (Ga), indium (In) or germanium (Ge). Namely, gallium displays very low carbon solubility. It is considered that the low carbon solubility exhibited by these additional metals may also reduce the carbon solubility in the nickel-based alloy.

[0077] This effect may hinder the dissolution of dehydrogenated carbon adatoms from dehydrogenation sites on the metal surface into the subsurface positions of the metal lattice and, thereby, the nucleation of carbon allotropes at the interface between the supported metal and the ceramic support material, and the lifting of the metal particles. According to the disclosure, the ceramic support material shows pores with sizes exclusively in the mesopore range. According to the disclosure, the mesopore size is in the range from 10 nm to 50 nm, preferably from 10 nm to 30 nm, and more preferably from 15 nm to 25 nm. The size of the Ni alloy particles should be from 8 nm to 40 nm, preferably from 10 nm to 20 nm, and more preferably to 15 nm to 20 nm. Smaller metal particles display minimum or no methane splitting catalytic activity, while larger particles display lower catalytic activities too. The atomic concentration of Ga, In, Ge, or any combination thereof, in the metal alloy is from 0 % to 50 % (atomic), preferably from 0 % to 30 %(atomic), and more preferably from 0 % to 15 % (atomic). The methane splitting reaction occurs at the surface of the pore-confined metal nanoparticles, particularly those accessible at the pore mouths. The formed hydrogen easily evolves, while carbon forms mostly crystalline tubular structures such as multiwall carbon nanotubes and carbon nanofibers.

[0078] Surprisingly, it was discovered that very small Co particles, with sizes in the range from 5 nm to 10 nm, display intrinsic base-growth behavior. Albeit the methane splitting catalytic activity of cobalt is slower than on nickel and nickel alloy catalysts, the carbon produced are exclusively few wall carbon nanotubes, which exhibit typically from 2 to 10 graphitic walls. During the methane splitting reaction, the supported Co particles keep their integrity and remain static, adhered to the ceramic support regardless of them being or not confined to pores. This makes the Co-based catalysts ideal for producing higher- value carbon nanomaterials, where hydrogen production becomes a side-product of the methane splitting reaction.

[0079] In an embodiment, the ceramic support material is preferably an oxide selected from those of silicon (Si), aluminium (Al), zirconium (Zr), titanium (Ti), magnesium (Mg), and combinations thereof, preferably an oxide selected from those of Si, Al, Zr, and combinations thereof, and more preferably is an oxide selected from silicon dioxide (SiCh), aluminium oxide (AI2O3) and combinations thereof. According to a preferred embodiment, the ceramic support material has particles in the range from 0.1 pm to 100 pm. Particularly if the support is mesoporous, its primary particle size preferably is in the range from 0.5 pm to 100 pm, and, more preferentially, from 5 pm to 25 pm.

[0080] In an embodiment, according to the disclosure, the catalyst shows a metal loading, defined as the mass of total metal, over the mass of the catalyst and expressed as a percentage, which is in the range from 5 % to 60 %, preferentially, from 10 % to 50 % and more preferably from 30 % to 50 %.The following pertains to Methane Splitting Reaction Conditions

[0081] In an embodiment, according to the disclosure, the supported catalyst operates at a temperature from 500 °C to 650 °C, preferentially, from 550 °C to 600 °C. These reaction temperatures are required toretain the functionality of the hydrogen perm-selective membranes placed in the membrane compartment of the reactor. However, methane splitting is an endothermic, reversible reaction, and, as such, the methane conversion per reactor pass is limited by the thermodynamic equilibrium for the reaction at the reaction conditions. For example, at 1 bar total pressure and 550 °C, the equilibrium, i.e. maximum methane conversion is ca. 48 %, and at 1 bar total pressure and 650 °C this maximum methane conversion is 72 %. According to the disclosure, a membrane reactor is applied for removing hydrogen selectively while enabling a recycling of the unconverted methane to the reaction compartment. The membrane reactor allows a quasi-continuous operation of the methane splitting process with a high- purity hydrogen outlet stream, subjected only to cyclic catalyst regeneration and carbon collection at cyclic intervals.The following pertains to the macro-support to the supported catalyst

[0082] In an embodiment, the supported catalyst may be applied onto a macro-support, which is a solid refractory material that confers stability to the catalyst and determines the gas flow patterns inside the reaction compartment of the membrane reactor.

[0083] In an embodiment, according to the disclosure, the macro-support material, once loaded with catalyst, is inserted and sealed into the reaction compartment of the reactor. The macro-support exists in the form of one single or a multitude of individual pieces in the reaction compartment of the reactor. The supported catalyst ideally forms a dense overlayer on the macro-support; this prevents the carbon growth in the space between catalyst layers and, thereby, prevents the peeling-off the topmost catalyst layers. Additionally, the particles of the supported catalyst should be bound to the macro-support using a layer that is impermeable to methane. Such a binding underlayer prevents the carbon growth in the space between the macro-support surface and the first layer of supported catalyst particles, and it thereby prevents the detachment of the catalyst layer from the macro-support. Different approaches can be followed to attain this binding of the catalyst to the macro-support; using a sol-gel approach has, however, many advantages. It is a rheological liquid when applied, which allows to capture a monolayer of the supported catalyst particles and can be easily transformed into a solid binding film via calcination or annealing. The calcined underlayer should display a high melting temperature, preferably above 700 °C, to prevent atomic mobility at the operating temperatures and should be impermeable to methane. Materials such as silica or alumina are suitable for this application.

[0084] In an embodiment, the macro-support should comprehend a bundle of channels to host the supported catalyst on their inner surface. The macro-support can be made of metal tubes, such as steel, or a ceramic monolith, or a metallic honeycomb; this bundle of channels is hereafter named as macrosupport module. The preferential implementation uses metal tubes, with an inner diameter from 0.25 cm to 2 cm, and preferably from 0.5 cm to 1 cm.

[0085] In an embodiment, since the methane splitting reaction is endothermic, the macro-support incorporates heating elements. In one preferred embodiment of the disclosure, these elements consist of electrical resistances for direct ohmic heating inserted in the shell side of the tubes - in the case of metal tubes - or of the channels - in the case of a ceramic monolith. In another embodiment, a thermal fluid is circulated along flow channels within the macro-support.The following pertains to catalyst regeneration

[0086] In an embodiment, as the methane splitting reaction progresses, the metal particles on the catalyst may be encapsulated with carbon, and therefore the catalyst activity for methane splitting decreases. A catalyst regeneration treatment is performed periodically, in a cyclic fashion, to recover partially or totally its activity. This regeneration is performed by a selective interfacial hydrogenative gasification of carbon, as described in document WO202121287A1.

[0087] In an embodiment, the reactor operates by cyclically alternative hydrogen production stages and catalyst regeneration stages. During the regeneration stage, pure hydrogen is supplied to the top of the reaction channels. The gas directing valve should be turned to link the hydrogen input to the channels of the reaction compartment, closing the inlet to the membranes compartment. The methane feed stream to the reactor should be ceased during the duration of the catalyst regeneration stage. The membrane compartment continues operating, fed from the methane recycling tube. The total pressure at the reaction compartment increases during the regeneration step because of the hydrogen inlet, which promotes partial carbon hydrogenative gasification. The regeneration stage normally takes froml min to 5 min, while the hydrogen production stage normally takes from 1 hour to 5 hours. The production time interval should be set to maximize hydrogen productivity. However, the size of the carbon particles depends on the production time; hydrogen productivity and carbon particle size should be balanced. It was observed a second catalyst deactivation mechanism, with a much longer time constant. It was discovered that every 15 days to 60 days, and more preferably every 15 days to 30 days, it is necessary to stop the reaction and feed the reactor with an inert gas, such as nitrogen or argon, from 1 hour to 20 hour, and more preferably from 4 hours to 8 hours. The inert gas should contain a very low oxygen concentration, which should be lower than 2 ppm in the case of argon and lower than 1 ppm in the case of nitrogen. It is considered that this second regeneration procedure is needed for spelling carbon atoms dissolved in the metal catalyst.

[0088] In an embodiment, working at low temperatures has many advantages and disadvantages. At low temperatures, there is a smaller thermal energy waste, the materials last longer and are normally cheaper, and the operation of the reactor is safer. However, the reaction equilibrium methane conversion is lower, requiring a more complex membrane reactor. This low-temperature methane splitting reactor allows easier thermal integration than alternative reactors operating at higher temperatures. For example, itallows the production of hydrogen needed for running a hydrogen turbine, where the turbine's exhausting gases can be used for heating the methane splitting reactor.Brief Description of the Drawings

[0089] The following figures provide preferred embodiments for illustrating the description and should not be seen as limiting the scope of the disclosure.

[0090] Figure 1: Simplified drawing illustrating three stages of the confinement of particles of a catalytically active metal to the pores of porous ceramic support, to produce a nanostructured catalyst according to one embodiment of the disclosure: i) empty pores; ii) pores filled with metal precursor solution; and iii) metal nanoparticle confined to the pore, wherein:(1.1) Porous ceramic support;(1.2) Pores of the support;(1.3) Pores filled with metal precursor solution;(1.4) Metal nanoparticle confined to the pores.

[0091] Figure 2: Schematic representation for the supports, for loading particles of the nanostructured catalyst, according to different embodiments of the disclosure, wherein:(2.1) Ceramic monolith macro-support module;(2.2) Metal tubular macro-support;(2.3) Bundle of metal tubes;(2.4) Reaction module;(2.5) Plate-and-frame reaction module;(2.6) Inner wall of macro-support channel, for catalyst loading;(2.7) Inserted tube for passing thermal fluids or electrical resistances;(2.8) Inserted thermocouple for monitoring temperature;(2.9) Inserted tube for passing thermal fluids or electrical resistances;(2.10) Inserted electrical resistances for heating the reaction module.Figure 3: Simplified scheme of the process of coating a solid catalyst on the macro-support. The nanostructured catalyst is depicted embedded in a primary sol-gel coating, being:(3.1) Macro-support;(3.2) Sol-gel layer;(3.3) Catalyst particles;(3.4) Calcinated / hardened sol-gel layer.

[0092] Figure 4: Simplified scheme of the membrane reactor, consisting of a reaction compartment with a nanostructured catalyst loaded on a macro-support; a membrane compartment in which H2 is pumpedthrough a membrane module out of the reactor and methane is recycled to the entrance of the reaction compartment; and a carbon collection container, that remains isolated from the reactor, and it is only periodically brought on line with the reaction compartment through a solid collection valve, for carbon removal, wherein:(4.1) Reaction compartment;(4.2) Separation compartment;(4.3) Reaction module;(4.4) Upper chamber;(4.5) Perforated plate - connecting macro-support channels to the upper chamber;(4.6) Three-way valve;(4.7) Blower / pump driving gas from reaction chamber to separation chamber;(4.8) Regeneration hydrogen inlet;(4.9) Retentate tube;(4.10) Membrane module;(4.11) Low-pressure permeated hydrogen;(4.12) Permeate vacuum pump;(4.13) Methane inlet to the membrane reactor;(4.14) Methane inlet valve;(4.15) Cone-shaped reactor bottom;(4.16) On-off valve controlling carbon exit;(4.17) Carbon receiver container;(4.18) On-off valve of the carbon receiver container.

[0093] Figure 5: Illustration of a cyclic methane splitting profile, using periodic interfacial hydrogenation, over a nickel gallium alloy (NiGa) / silicon dioxide (SiO2) catalyst, being:(5.1) Production stage;(5.2) Regeneration stage;(5.3) Production peak after the regeneration stage.Detailed Description

[0094] The present disclosure relates to the field of catalytic reactors, specifically to a low-temperature catalytic methane splitting reactor operating between 500 °C and 650 °C. The reactor is designed for the continuous production of high-purity hydrogen and graphitic carbon, with no CO2emissions, using a novel combination of supported catalysts, macro-support structures, and a membrane-based gas separation system.

[0095] The disclosure describes a reactor to conduct the methane splitting reaction, also known as methane decomposition or cracking, which transforms methane into hydrogen and solid carbon. Specifically, the present disclosure presents a low-temperature catalytic methane splitting reactor - from 500 °C to 650 °C, for producing high-purity hydrogen and graphitic carbon. Solid catalyst systems to be used in the reactor are also disclosed - supported nanoparticle composite catalyst bound to a macrosupport. The reactor comprises two compartments: a first compartment that hosts the catalyst-loaded macro-support, wherein the methane splitting reaction occurs, and a second compartment that hosts a gas separation membrane module. The catalyst comprehends catalytic metals or metal alloys - nickel, cobalt, gallium, and any combination thereof, dispersed on a nanostructured oxide support. Stable reactor operation is described, by the implementation of cyclical regeneration, using hydrogen as regeneration gas. The disclosed catalytic membrane reactor allows the stable and continuous production of hydrogen and carbon, without carbon dioxide (CO2) emissions.The following pertains to the supported catalyst

[0096] It was observed that, by confining nickel nanoparticles within the pores of mesoporous ceramic supports while retaining access to the dispersed metal close to the pore mouths - Figure 1, the formed carbon shows structures characteristic of the base-growth mechanism. The catalytic activity of nickel is too high, and, to improve the morphological stability of the supported catalyst, nickel is alloyed with at least a second metal showing a very low intrinsic carbon solubility, which produces stable alloys with nickel, such as gallium (Ga), indium (In) or germanium (Ge). Namely, gallium displays very low carbon solubility. It is considered that the low carbon solubility exhibited by these additional metals may reduce the carbon solubility in the Ni-based alloy, too. This effect may hinder the dissolution of dehydrogenated carbon adatoms from dehydrogenation sites on the metal surface into the subsurface positions of the metal lattice and, thereby, the nucleation of carbon allotropes at the interface between the supported metal and the ceramic support material, and the lifting of the metal particles.

[0097] According to the disclosure, the ceramic support material - Figure 1, feature 1.1 - shows pores with sizes exclusively in the mesopore range - Figure 1, feature 1.2. According to the disclosure, the mesopore size is in the range from 10 nm to 50 nm, preferably from 10 nm to 30 nm, and more preferably from 15 nm to 25 nm. The size of the Ni alloy particles should be from 8 nm to 40 nm, preferably from 10 nm to 20 nm, and more preferably from 15 nm to 20 nm. Ni alloy particle size is controlled by filling the mesopores of the support - Figure 1, feature 1.3 - with active metal precursors, which are then turned into metallic alloys - Figure 1, feature 1.4. Smaller metal particles display minimum or no methane splitting catalytic activity, while larger particles display lower catalytic activities too. The atomic concentration of Ga, In, Ge, or any combination thereof, in the metal alloy is from 0 % to 50 % (atomic), preferably from 0 % to 30 %, and more preferably from 0 % to 15 %. The methane splitting reaction occurs at the surface of the pore-confined metal nanoparticles, particularly those accessible at the pore mouths.The formed hydrogen easily evolves, while carbon forms mostly crystalline tubular structures such as multiwall carbon nanotubes and carbon nanofibers. The formed carbon is mechanically prevented from diffusing through the side interfaces metal catalyst / support and reaching the bottom of the inclosing pore where the metal catalyst is inserted; thereby the so-called tip growth is hampered. The nickel-based catalysts are designed to maximize hydrogen production rates, wherein carbon is a market-valued sideproduct of the methane splitting reaction.

[0098] Surprisingly, it was discovered that very small Co particles, with sizes in the range of 5 nm to 10 nm, display intrinsic base growth behavior. Albeit the methane splitting catalytic activity of cobalt is slower than on nickel and nickel alloy catalysts, the carbon particles produced are exclusively few wall carbon nanotubes, which exhibit typically from 2 to 10 graphitic walls. During the methane-splitting reaction, the supported cobalt particles keep their integrity and remain static, adhering to the ceramic support regardless of them being or not confined to pores. This makes the cobalt-based catalysts ideal for producing higher-value carbon nanomaterials, where hydrogen production becomes a side-product of the methane splitting reaction.

[0099] In an embodiment, the ceramic support material is preferably an oxide selected from those of Si, Al, Zr, Ti, Mg, and combinations thereof, preferably an oxide selected from those of Si, Al, Zr, and combinations thereof, and more preferably is an oxide selected from SiC , transitional AI2O3 and combinations thereof.

[0100] According to a preferred embodiment of the disclosure, the size of the ceramic support material ranges from 10 nm to 100 pm. Preferably from 0.5 pm to 100 pm, and more preferentially, from 5 pm to 25 pm.

[0101] According to the disclosure, the catalyst shows a metal loading, defined as the mass of total metal, over the mass of the catalyst and expressed as a percentage, which is in the range from 5 % to 60 %(wt • meta l / wt • meso porous support+metai), preferentially, from 10 % to 50 % (wt.metai / wt • mesoporous support+metai) and, more preferably from 30 % to 50 % (wt.metai / wt • mesoporous support+metai).The following pertains to macro-support to the supported catalyst

[0102] In an embodiment, the macro-support serves to hold the supported catalyst and should allow: i) the free drop of the formed carbon particle; ii) maximize the macro-surface per volume of the reactor; iii) minimize the construction costs; iv) minimize the hydrogen volume needed for the catalyst regeneration; and v) control the temperature since the methane splitting reaction is endothermic. A preferential design of the reaction module, which fulfills these requirements, considers using a ceramic monolith - Figure 2, feature 2.1, a metal honeycomb module - Figure 2, feature 2.2, a bundle of metal tubes - Figure 2, feature 2.3 - assembled as a module - Figure 2, feature 2.4, or a metallic, carbon-based or ceramic plate-and- frame module - Figure 2, feature 2.5. For simplicity, all types of channels with the macro-support will behereafter referred to reaction modules - Figure 2, features 2.1 to 2.5. A preferential implementation of the reaction module, comprehend tubular channels that go through the entire module with an inner diameter from 0.25 cm to 2 cm, and, more preferentially, from 0.5 cm to 1 cm; larger diameters are required when the set size for the carbon particles is also large, or the length of the tubular holes are lengthy. The equivalent diameter is such that it should encompass at least four times the length of the carbon particles, to guarantee that no clogging of each channel, gap or tube occurs during carbon detachment. The preferential length of the tubular holes is from 1 cm to 20 m, and, more preferentially, from 30 cm to 3 m. Longer tubular channels are prone to clogging with the produced carbon particles. The geometry of the tubular holes should facilitate the unimpeded growth of the carbon particles. Two preferential geometries are the cylindrical and the honeycomb - Figure 2, feature 2.2 and feature 2.3, where the honeycomb geometry displays a larger specific area and then a higher power density of the catalytic module. In the case of the plate-and-frame reaction module - Figure 2, feature 2.4 -the distance between two consecutive plates should be from 1.5 mm to 10 mm, with a more preferential distance from 2 mm to 4 mm; the depth of the plate-and-frame macro-support should be from 5 cm to 10 m, and more preferentially from 20 cm to 1 m.

[0103] In an embodiment, the supported catalyst should be applied over the bore surface side of the macro-support channels, as a monolayer - Figure 2, feature 2.6. For holding the supported catalyst to the macro-support and preventing catalytic activity at the interface of the macro-support / supported catalyst, a thin and dense binding underlayer should be applied over the macro-support surface (or the inner surface of the channels) and the supported catalyst applied as a monolayer - cf. Figure 3. This Figure illustrates the macro-support surface (or the inner surface of the channels) -feature 3.1 -and a deposited sol-gel layer - feature 3.2. Figure 3, feature 3.3 illustrates a supported catalyst, applied as a monolayer over the sol-gel film, and Figure 3, feature 3.4, illustrates the sol-gel layer already calcinated and binding a monolayer of supported catalyst particles.

[0104] In an embodiment, treferential implementations of this binding underlayer use a sol-gel approach, such as using tetraethoxisylane (TEOS) and 3-(trimethoxysilyl)propyl methacrylate (MAP) [5], to prepare a silica layer, or a sol-gel approach to produce an alumina layer [6], The sol-gel layer can be applied evenly [7] by suitable methods such as dip-coating, spraying, or slot-die. Another preferential implementation considers using a thermosetting polymer, such as a resin of resorcinol-formaldehyde, applied over the bore side surface of the catalytic module holes, normally as an ink formulation [8], The supported catalyst particles should be spread over the boding underlayer before it is calcinated or carbonized, respectively, for the sol-gel or polymeric formulations; the adhesion displayed by the paste phase of this coating allows for holding the supported catalyst particles as a monolayer. After drying, the sol-gel layer should be calcinated at the maximum operating conditions of the reactor, from 500 °C to 650 °C, and preferentially at 600 °C, and the thermosetting polymer coating should be carbonized at themaximum operating conditions of the reactor, from 500 °C to 650 °C, and preferentially at 600 °C, under an inert atmosphere. The boding underlayer should cover less than 50 % of the supported catalyst particles and more preferentially ca. 25 %. During the drying process, the macro-support can be rotated horizontally to avoid inhomogeneities in thickness of the silica ceramic, or the alumina ceramic or the thermosetting polymer. A preferential method for applying a homogeneous monolayer of the supported catalytic particles over the sol-gel coating is flushing the particles down the tubular supports. The sol-gel coating, or the polymeric ink, will capture the flowing down particles to a monolayer thickness; the excess particles should be reused. The macro-support, after being loaded with the supported catalyst, should be named herein as the reaction module or catalytic module.

[0105] In an embodiment, since the methane splitting reaction is endothermic, reaction modules should allow local heating, either using electric resistances or thermal fluid. The heating power of these heating systems should be at least equal to the reaction enthalpy and preferentially two to three times above for assisting during the start-up of the reactor to heat it more quickly to the steady-state temperature. The reactor should have means to control the internal temperature to a maximum deviation of ± 50 °C, and, more preferentially, to a maximum deviation of ± 15 °C. In a preferential implementation, the reaction modules may comprehend special holes for thermal controlling - e.g. inserting a thermocouple, or thermal heating, e.g., inserting a tube for circulating a thermal fluid or electrical heating resistances - cf. Figure 2, feature 2.7. Figure 2, features 2.8 and 2.9 illustrates a preferential implementation of the thermal control and for thermal heating. A more preferential implementation of the heating resistances for the case of a plate-and-frame catalytic reaction module considers the use of electric conductive fibrous carbon plates (e.g. carbon paper) or folded metal foil plates, as sketched in Figure 2, feature 2.10. On the inner surface of the metal foil surface, it should be applied a thin foil of a thermostable electric insulating material, such as Thermiculite™. When folding the plate to the middle, wire-shaped electrical resistances should be placed inside. These resistances should then be connected to the plate-and-frame holders, where one should be the phase and the other the neutral - in the case of alternate current (AC) is used - or opposite potential signs if direct current (DC) is used. In the case of carbon-based plates, such as carbon paper, taking advantage of the electrical resistivity of these substrates, an electrical potential can be applied directly to the opposite ends of the reaction module, generating heat directly at the reaction module.The following pertains to the reactor

[0106] In an embodiment, a preferential implementation of the methane-splitting reactor is sketched in Figure 4. Since it works from 500 °C to 650 °C, preferentially from 550 °C to 600 °C, the reactor should be well thermally insulated and electrically or thermally heated. According to the disclosure, the reactor - cf. Figure 4 - is divided into two main compartments. The methane splitting reaction takes place in the first compartment, the reaction compartment, which houses the solid catalyst - Figure 4, feature 4.1. Theselective recovery of the hydrogen product from unconverted methane takes place in a second compartment, the membrane compartment, which houses a module with membranes for selective gas separation - Figure 4, feature 4.2. The reaction compartment contains the reaction modules - cf. Figure 4, feature 4.3. The reaction modules should fit into the reactor, leaving a chamber between the upper part of the reaction and the top of the reaction modules - cf. Figure 4, feature 4.4; this chamber is named as the top reactor chamber, for the distribution of the methane. Moreover, the reaction modules should fit into the reactor, such as if a gas feed enters through the top of the reactor; they can only reach its bottom if goes through the tubular channels of the reaction modules - Figure 4, feature 4.5. This top reactor chamber should connect to the separation compartment through a 3-way valve - cf. Figure 4, feature 4.6. During the production stage, this valve allows the reaction product stream to flow to the membrane module; a blower or a pump drives this flow- cf. Figure 4, feature 4.7. During the regeneration stage, the 3-way valve closes the flow to the membrane module and opens the regeneration hydrogen flow - Figure 4, feature 4.8.

[0107] After crossing the membrane module, the reaction product stream - retentate - should have a hydrogen low concentration, which permeates the membrane driven by a low permeation pressure. The retentate stream is connected to the bottom part of the reactor - cf. Figure 4, feature 4.9. At the membrane module - cf. Figure 4, feature 4.10, the feed pressure is set by the blower or by the pump and should be set from 1X105Pa to 10xl05Pa (1 bar to 10 bar), and preferably from 1 x 105Pa to 5 x 105Pa (1 bar to 5 bar). The permeate pressure should be preferentially set from 1 x 105Pa to 0.001 x 105Pa (1 bar to 0.001 bar), and more preferentially from 0.05 x 105Pa to 0.01 x 105Pa (0.05 bar to 0.01 bar) - cf. Figure 4, feature 4.11, which should be driven by a vacuum pump - cf. Figure 4, feature 4.12. According to the disclosure, the retentate stream is driven back to the reaction compartment, which operates at a total pressure in the range of 15 x 105Pa (15 bar), preferably at a total pressure in the range of l x 105Pa to2 x 105Pa (1-2 bar), and more preferably at atmospheric pressure - cf. Figure 4, feature 4.9. The inlet methane should be fed to the bottom of the reactor - cf. Figure 4, feature 4.13; the inlet methane flowrate is controlled by the removal of hydrogen at the membrane module; generally, the molar feed flowrate of methane should be approximately equal to half of the hydrogen molar flowrate leaving the reactor. The total pressure at the reaction compartment should be kept at approximately 1 x 105Pa (1 bar), during the production stage. During the regeneration stage, the methane inlet is closed - Figure 4, feature 4.14, hydrogen is fed to the top of the reaction modules, filling up the reaction tubular holes and promoting the carbon particles peeling-off and the regeneration of the catalyst. During this stage, the pressure at the reaction compartment would increase, feeding backward the membrane module through the retentate tube 4.9. However, since the membrane module still permeates hydrogen, the total pressure stays mostly constant. The regeneration stage normally takes from 1 min to 5 min, while the hydrogen production stage normally takes from 1 hour to 5 hours.

[0108] The production time interval should be set to maximize hydrogen productivity. However, the size of the carbon particles depends on the production time; hydrogen productivity and carbon particle size should be well-balanced. It was observed a second catalyst deactivation mechanism, with a much longer time constant. Every 15 days to 60 days, and more preferably every 15 days to 30 days, it is necessary to stop the reaction and feed the reactor with an inert gas, such as nitrogen or argon, from 1 hour to 20 hours, and more preferably from 4 hours to 8 hours. The inert gas should contain a very low oxygen concentration, which should be lower than 2 ppm in the case of argon and lower than 1 ppm in the case of nitrogen.

[0109] In an embodiment, the bottom of the reactor is cone-shaped, designed to collect the solid carbon, especially during the regeneration stage - cf. Figure 4, feature 4.15. To facilitate the release of the carbon particles from the catalyst, the reactor may be vibrated, either continuously or specifically during the catalyst regeneration step, with the help of an external vibrating unit. Periodically, an evacuated carbon collection container is hermetically docked to the bottom of the reactor; and a valve at the bottom of the reaction compartment is opened to allow the carbon particles to flow into the carbon collection container - cf. Figure 4, feature 4.16. After this transfer process, the valve at the reactor's bottom is closed again to continue operation. The container should be evacuated to prevent oxygen or an inert gas from entering the reactor, or preferentially, it should be filled with methane - cf. Figure 4, feature 4.17. In a preferential implementation, the carbon receiver container has also an on-off valve for venting the air and filling it with methane or just leaving it under vacuum - cf. Figure 4, feature 4.18. The dead volumes, that may exist when coupling the reactor to the container, should be purged as well.The following pertains to gas separation membranes

[0110] In an embodiment, the membranes are critical to keep the hydrogen partial pressure inside the reactor well below the partial pressure corresponding to the conversion equilibrium. For example, at 550 °C and lxlO5Pa (1 bar) of total pressure, the equilibrium conversion is 48 %, and then, the maximum hydrogen partial pressure is 0.65xl05Pa (0.65 bar). The steady-state methane conversion should be ca. 25 % below the equilibrium constant so the apparent reaction kinetics is not impaired by the backward reaction kinetics. This means that the optimum hydrogen partial pressure is ca. 0.5 x 105Pa (0.5 bar). The hydrogen only permeates a membrane if the permeate pressure is well below the retentate pressure, and then the permeate pressure should be below 0.5 x 105Pa (0.5 bar), and preferably from 0.1 x 105Pa to 0.05 x 105Pa (0.1 bar to 0.05 bar). Preferentially, it should be used palladium / silver membranes or carbon molecular sieve membranes with high performance for the separation of hydrogen from methane; more preferentially, carbon molecular sieve membranes should be used since they display higher stability and low price, though they have a slightly lower separation performance.The following pertains to general methods

[0111] The specific surface area, pore volume, and mesopore mean diameter of the ceramic support materials, as described in the disclosure, were determined by N2-physisorption at 77 K in a Micromeritics ASAP 2000 apparatus. In a standard experiment, the ceramic support material (ca. 100 mg, 0.6-0.8 mm particle size) was dried at 523 K under a dynamic vacuum for 10 hours before recording the N2 physisorption isotherm. The total pore volume was determined from the total nitrogen (N2) uptake at a relative pressure (P / Po) of 0.95, where Po , at atmospheric pressure, is the vapor pressure of nitrogen at 77 K. The specific surface area was determined by applying the Brunauer-Emmett-Teller (BET) method to the adsorption branch of the isotherm in the relative pressure (P / Po) range from 0.05 to 0.3. The average mesopore diameter was determined by applying the BJH formalism to the desorption branch of the isotherm.

[0112] The metal loading on the catalysts was determined by Inductively-Coupled-Plasma Optical Emission Spectrometry (ICP-OES) in an iCAP™ PRO analyzer (Thermo Fisher Scientific) after disaggregating 30 mg of the solid sample in a HNO3:HF (1:3 vol. ratio) solution at room temperature for 20 hours. Quantification was based on calibration lines created by the analysis of certified standards (CertiPUR®) in the expected concentration ranges for each metal. Total metal loadings are expressed as the mass of metal over the mass of the catalyst and expressed as a percentage.

[0113] The metal nanoparticle size on the catalysts was determined by Scanning-transmission electron microscopy ((S)TEM). The catalysts were first activated by a metal oxide reduction treatment by heating the sample from room temperature to 600 °C, using a heating ramp of 2 °C min1, and maintaining the temperature at 600 °C for 5 hours, under flow of 20 % Fh / He (vol.). Next, catalysts consisting of pore- confined Ni or Ni alloy nanoparticles were embedded in a low-viscosity epoxy resin (Spurr, Merck) and cured at 343 K overnight. The embedded specimens were then cut into nm-thick sections (nominal thickness of ca. 150 nm) using a Diatome 35° knife mounted on a Reichert-Jung Ultracut microtome, and the sections were collected on Cu TEM grids coated with a Formvar film and a carbon (300 mesh) overlay. For Co-based catalysts supported on non-porous ceramic supports the powder catalyst was directly dry cast onto Cu TEM grids coated with a lacey carbon film. Microscopy experiments were performed in a JEOL JEM2100F microscope operating at 200 kV. Number-averaged particle diameters were determined by measuring at least 100 nanoparticles on different micrographs.

[0114] Methane conversion and hydrogen production rates were determined online, during methane splitting reaction tests utilizing a mass spectrometer Pfeiffer Vacuum ThermoStar GSD 301.Example 1

[0115] Particulate SiO2 (SiliaSphere silica Gel, 5 pm, Silicycle, CAS No. 112926-00-8) with spherical primary particles of 5 pm average diameter and 20 nm average pore size (narrow normal distribution) was dried at 150 °C under dynamic vacuum for 3 h. Next, a solution of NifNOsh-SF O, and Ga NOsh-xh O (Merck, CAS Nrs 13478-00-7 and 69365-72-6 respectively), with a Ni:Ga molar ratio of 95:5, in 0.1 M HNO3was infiltrated into the porosity of the support under static vacuum. The overall metal concentration of the impregnating solution was adjusted to ca. 4 M. The volume of the impregnating solution applied was equivalent to 90 % of the total pore volume of the SiC support material. Next, the impregnate was dried under dynamic vacuum lxlO3Pa (10 mbar) overnight and the dried solid (in powder form) was loaded into a packed-bed reactor, mounted in an vertically-oriented tubular oven and treated at 350 °C for 3 h under flow of 2 % (vol.) nitric oxide(NO) in He (Linde), using a heating rate of 2 °C min1from room temperature to decompose the pore-confined metal nitrate precursors into pore-confined metal oxide nanocrystals. The impregnation and calcination steps were repeated four times to attain a high cumulative degree of pore filling by the metal oxide nanocrystals and, thus, a high density of pore-confined metal at the pore mouths, i.e. close to the outer surface of the spherical primary SiC particles, which were not completely filled after the first calcination, due to the removed volume of water and nitrates. The overall metal loading was 43 wt.%.

[0116] The supported catalyst material was loaded onto an a-ALOs planar macroporous support through wet-spray deposition of a circular monolayer, corresponding to ca. 0.5 mgcataiyst cm'2. The support was placed into the reaction compartment of the reactor (the inner surface of the channels) and the latter was sealed. To activate the catalyst, it was subjected to a thermal reduction treatment in the reactor, under flow of 50 % H2 in N2, at lxlO5Pa (1 bar) and heating to 550 °C at a heating rate of 1 °C min1to render the pore-confined metals into their active zerovalent alloy state. The catalyst was then exposed to a flow of pure methane, at lxlO5Pa (1 bar) and 550 °C, promoting the methane splitting reaction. As it would happen to any low-temperature catalytic methane splitting system, the catalyst suffers from gradual deactivation, requiring the implementation of a regeneration technique. A catalyst regeneration treatment was performed periodically, in a cyclic fashion, by interfacial hydrogenative gasification of carbon, as described in disclosure WO2020121287A1.

[0117] Cyclically, stages of hydrogen production by methane splitting of 1 hour duration were followed by purges of the gas inside the reaction compartment of the reactor with N2 flow for 10 min, and then catalyst regeneration stages of 5 min duration under flow of pure H2. Through all different cyclically alternating operation stages, the total pressure and temperature were maintained constant at lxlO5Pa (1 bar) and 550 °C. Figure 5 depicts the history of metal-normalized hydrogen production rate, expressed as grams of H2 produced per unit of metal mass in the catalyst and per hour of operation. The peaks refer to the hydrogen production stages - Figure 5, feature 5.1, while the interval refers to the catalyst regeneration stages - Figure 5, feature 5.2; the sharp peaks - Figure 5, feature 5.3, are related to the very fast methane splitting reaction kinetics right after the catalyst regeneration stage; the catalytic activity decreases substantially as the reaction progresses justifying the regeneration stages. It is seen that the catalyst is successfully regenerated from cycle to cycle, generating the shown profile with repeating cyclesthat display initial activity >0.4 gH2 Nica1h1, which then gradually drops to 0.2 gH2 NiGa1h1. The catalyst remains regenerable throughout the duration of the methane conversion text described in this example.Example 2

[0118] A square 7-channel ceramic monolith with a channel side of 0.5 cm, 20 cm long, was used, with a similar structure to the monolith depicted in Figure 2, feature 2.1. 6 inner walls of the monolith were dip- coated with a TEOS layer (the middle channel was not coated), according to [5], and then dip-coated with a monolayer of catalyst. The supported catalyst, prepared as described in Example 1, was applied using an ethylene glycol emulsion with 1 % catalyst mass content, to perform the dip-coating process. The monolith was inserted into an open-to-air oven, at 600 °C, to convert the coating into a silica-bonded catalyst monolayer. The monolith was turned slowly, horizontally, inside the oven. At the end, the channels were inspected to assess the quality of the catalyst layer.

[0119] The ceramic monolith was sealed with glass paste on both ends, to perforated steel plates; the interface between the plates and the tubes was made hermetic by the applied glass, which was hardened in an open-to-air oven, at 600 °C. The perforated steel plates were used to feed gases into the monolith channels. The uncoated central channel of the monolith and the smaller monolith side channels were used to feed thermal fluids, at 600 °C, whose flow was controlled by thermocouples (in the uncoated channels) and flowmeters, which maintained the reaction zones at 550 °C. The coated channels were fed with methane, to promote methane splitting, and periodically by hydrogen, to promote catalyst regeneration, as reported in document WO2020121287A1.

[0120] This reactor was run for 500 h at 550 °C, with a hydrogen production stage duration of 60 min, during which methane was fed at lxlO5Pa (1 bar) with a flowrate of 10 cm3min1, and a catalyst regeneration stage duration of 3 min, during which the hydrogen inlet flowrate was set to 10 cm3min1.Example 3

[0121] A set of 5 steel tubes with an internal diameter of 0.5 cm, 20 cm long, were taken. The bore-side of these tubes was coated, using a cylindrical paintbrush, with a layer of TEOS ca. 25 pm thick. The supported catalyst, prepared as described in Example 1, was applied using a turning brush initially loaded with the catalyst. The excess catalyst was then removed, making air circulate through the tube. Finally, the 5 tubes were inserted into an open-to-air oven, to convert the TEOS coating layer into silica, at 600 °C. The tubes were turned slowly, horizontally, inside the oven. At the end, the tubes were inspected to assess the quality of the catalyst layer.

[0122] On the shell side of the tubes, a heat flow of 40 W was applied from an electric resistance; a thermocouple attached to the external surface of a tube allowed to control the temperature. They were then placed in a perforated steel plate; a top perforated plate was applied. The set was held together using three threaded bars, see Figure 4; the interface between the plates and the tubes was madehermetic by the applied glue. The upper plate was fixed tight to the container; this allowed feeding of the bore side of all tubes through the top of the casing - Figure 4. A set of two tubular palladium / silver (23 wt.% of silver), with 1 cm of external diameter and 30 cm long. The tubular membranes have threaded steel fittings in both extremes, which were used to hermetically screw in the upper and bottom plates - Figure 4. The two plates were tightened hermetically to the casing. A tube was welded to the top side of both reaction and membrane casings. These tubes were used to connect a blower - actually, a small membrane pump - outside the oven; the tube connected to the reactor was equipped with a 3-way automatic valve, which allowed it to connect either to the membrane module or to the regeneration hydrogen inlet. The methane inlet was made through the side of the reactor casing at the bottom. Both casings were also connected using a tube welded on the side bottom; this tube was used to connect the retentate to the reactor.

[0123] A tube welded to the side of the membrane casing, at an intermediate position, was used to connect the membrane permeate side to the vacuum pump - a membrane vacuum pump delivering O.OlxlO5Pa (0.01 bar) of pressure. At the bottom of the reactor casing it was placed a globe valve, is used for collecting the carbon particles.

[0124] This reactor was run for 500 hours at 550 °C, with hydrogen production stages of a duration of 60 min and catalyst regeneration stages of a duration of 3 min. Methane was fed at l.lxlO5Pa (1.1 bar) with the flowrate imposed by the permeated hydrogen; the feed flowrate was approximately 10 cm3min1, while the hydrogen flowrate was approximately 20 cm3min1. The hydrogen flowrate during catalyst regeneration stages was set to 10 cm3min1.

[0125] This project has received funding from the European Commission through the European Union's Horizon 2020 programme - FET Proactive research and innovation action - under grant agreement No. 952219 (112CO2).

[0126] The term "comprising" whenever used in this document is intended to indicate the presence of stated features, integers, steps, components, but not to preclude the presence or addition of one or more other features, integers, steps, components or groups thereof.

[0127] As used herein, the term "about" a number refers to that number plus or minus 10 % of that number. The term "about" a range refers to that range minus 10 % of its lowest value and plus 10 % of its greatest value.

[0128] Where ranges are given, endpoints are included. Furthermore, it is to be understood that unless otherwise indicated or otherwise evident from the context and / or the understanding of one of ordinary skill in the art, values that are expressed as ranges can assume any specific value within the stated ranges in different embodiments of the invention, to the tenth of the unit of the lower limit of the range, unless the context clearly dictates otherwise. It is also to be understood that unless otherwise indicated orotherwise evident from the context and / or the understanding of one of ordinary skill in the art, values expressed as ranges can assume any subrange within the given range, wherein the endpoints of the subrange are expressed to the same degree of accuracy as the tenth of the unit of the lower limit of the range.

[0129] The disclosure should not be seen in any way restricted to the embodiments described and a person with ordinary skill in the art will foresee many possibilities to modifications thereof.

[0130] The above described embodiments are combinable.

[0131] The following claims further set out particular embodiments of the disclosure.References[1] Muradov NZ. Thermocatalytic process for CO2-free production of hydrogen and carbon from hydrocarbons. US 2007 / 0111051 Al, 2007.[2] Magalhaes Mendes AM, Mateos Pedrero C, Dias Catarino M. Catalytic methane decomposition and catalyst regeneration, methods and uses thereof. WO 2020 / 12287 Al, 2020.[3] Method for parallel production of hydrogen and carbon-containing products. WO 2013 / 004398 A8, 2013.[4] Pocock G, Cornejo A, Chua HT. A process for producing hydrogen and graphitic carbon from hydrocarbons. WO 2016 / 154666 Al, 2016.[5] Longhi M, Casagrande RB, Kunst SR, Santos V dos, Ferreira JZ. Obtainment and Characterization of a Silicon alkoxides-based Coating Applied to a Substrate of Stainless Steel 316L for Use in Biomaterials. Mater Res 2019;22. https: / / doi.org / 10.1590 / 1980-5373-mr-2018-0514.[6] Hubert T, Schwarz J, Oertel B. Sol-gel alumina coatings on stainless steel for wear protection. J Sol- Gel Sci Technol 2006;38:179-84. https: / / doi.org / 10.1007 / sl0971-006-6470-7.[7] Innocenzi P, Abdirashid MO, Guglielmi M. Structure and properties of sol-gel coatings from methyltriethoxysilane and tetraethoxysilane. J Sol-Gel Sci Technol 1994;3:47-55. https: / / doi.org / 10.1007 / BF00490148.[8] Rodrigues SC, Whitley R, Mendes A. Preparation and characterization of carbon molecular sieve membranes based on resorcinol-formaldehyde resin. J Memb Sci 2014;459:207-16. https: / / doi.Org / 10.1016 / j.memsci.2014.02.013.

Claims

C L A I M S1. Reactor for production of hydrogen and carbon from methane splitting, comprising a gas inlet for methane flow, a reaction chamber for methane splitting, a gas separation membrane for separating and purifying hydrogen, a gas outlet for hydrogen flow, a bottom outlet for carbon removing by gravity from the reaction chamber; wherein the reactor comprises a pipe connecting the reaction chamber to the gas separation membranes to flow product reaction for introducing the reaction products of the reaction module into the plurality of gas separation membrane; wherein the reaction chamber comprises a reaction module; wherein the reaction module comprises a plurality of gas flow channels with a macro-support; wherein the macro-support comprises a catalytic support loaded with a metal catalyst for contacting with the methane flow for catalyzing the methane splitting reaction; wherein the said catalyst is a plurality of Ni or Co alloy particles.

2. Reactor according to the previous claim, wherein the operation temperature of the reaction module ranges from about 500 °C to 650 °C; preferably from about 550 °C to 600 °C.

3. Reactor according to any of the previous claims, wherein the reaction module has an operation pressure and the inlet pressure of the gas separation membrane ranges from about l x 105Pa to 10 x 105Pa; preferably from about 1 x 105Pa to 5 x 105Pa.

4. Reactor according to any of the previous claims, wherein the gas separation membrane has an outlet pressure ranging from about 1 x 105Pa to 5 x 105Pa; preferably from about 1 x 105Pa - 2 x 105Pa.

5. Reactor according to any of the previous claims, wherein a transmembrane hydrogen partial pressure differences in the gas separation membranes range from 0.1 xlO5Pa - O.OOlxlO5Pa; preferably from 0.05 xlO5Pa - 0.01xl05Pa.

6. Reactor according to any of the previous claims, wherein the geometry of the reaction module is selected from a hexagonal shape, a honeycomb shape, a tubular shape, a cubical shape or a parallelepipedal shape.

7. Reactor according to any of previous claims, wherein the geometry of the plurality of gas flow channels are selected from: tubular, hexagonal or plate stack.

8. Reactor according to the previous claim, wherein the plurality of gas flow channels comprises at least a heating device for controlling the temperature reaction, preferably an electrical heating resistance or a thermal fluid.

9. Reactor according to the previous claim, wherein the heating device is arranged in a tube or a plate.

10. Reactor according to any of the previous claims, comprising a temperature measuring device arranged between the plurality of gas flow channels, preferably a thermocouple.

11. Reactor according to any of the previous claims, wherein the plurality of gas flow channels has an inner equivalent diameter from 0.25 cm to 2 cm; more preferably from 0.5 cm to 1 cm.

12. Reactor according to the previous claims, wherein each channel of the plurality of gas flow channels traverses the entire module, with a preferential length of the tubular channels from 1 cm to 20 m; more preferably from 30 cm to 3 m.

13. Reactor according to any of the previous claims, wherein the plurality of nickel or cobalt alloy catalysts is alloyed with an element of the list consisting of: gallium, indium, germanium, or combination thereof.

14. Reactor according to any of the previous claims, further comprising alumina ceramic, or a silica ceramic, or a thermosetting polymer for bounding the plurality of nickel or cobalt alloy catalysts to the macro-support; preferably the catalyst is embedded in the alumina ceramic or a silica ceramic and / or thermosetting polymer.

15. Reactor according to the previous claims, wherein the catalytic support comprises mesopores ranging from 10 nm to 50 nm; preferably from 10 nm to 30 nm; more preferably from 15 nm to 25 nm.

16. Reactor according to the previous claim, wherein the size of each particle of the plurality of nickel or cobalt alloy catalysts particles ranges from 8 nm to 40 nm; preferably from 10 nm to 20 nm; more preferably from 15 nm to 20 nm.

17. Reactor according to the previous claims, wherein the distance between consecutive plates ranges from 1.5 mm to 10 mm, more preferably from 2 mm to 4 mm.

18. Reactor according to the previous claims, wherein the depth of the parallelepipedal shaped support is from 5 cm to 10 m; more preferably from 20 cm to 1 m.

19. Reactor according to any of the previous claims, wherein the mass amount of gallium, indium or germanium in each particle of the plurality of nickel or cobalt alloy catalysts ranges from 0 % to 50 % (wt.eiement / wt.metai particle); preferably from 0.1 % to 30 % (wt.eiement / wt. metai particle); more preferably from 1 % tO 15 % (wt.element / wt • metal particle).

20. Reactor according to the previous claims, wherein the mesoporous support has particles ranging from 0.1 pm to 100 pm; preferably from 0.5 pm to 100 pm; more preferably from 5 pm to 25 pm.

21. Reactor according to any of the previous claims, wherein the metal loading of the reaction module ranges from 5 % to 60 % (w / w); preferably from 10 % to 50 % (w / w); more preferably from 30 %to 50 % (w / w).

22. Reactor according to the previous claims, wherein parallelepipedal shape of the reaction module contains slots between the rectangularly shaped plurality of gas flow channels, for heating device plate placement.

23. Reactor according to any of previous claims, further comprising a second gas inlet for introducing a gas regeneration flow, for removing the carbon deposition in the mesoporous support, allowing the regeneration of the plurality of nickel or cobalt alloy catalyst particles.

24. Reactor according to the previous claim, wherein gas regeneration stream is a hydrogen stream, preferably a pure hydrogen stream.

25. Reactor according to any of previous claims, wherein the bottom outlet is cone or funnel shaped.

26. Reactor according to any of the previous claims, comprises a container arranged below a bottom of the reactor, for carbon collection.

27. Reactor according to any of the previous claims, comprising a third gas inlet to introduce a gas methane flow to the container, for protecting the produced carbon.

28. Reactor according to any of the previous claims, wherein the support is a metallic support, a ceramic support, a carbon-based support, or a combination thereof.

29. Reactor according to any of the previous claims, wherein the ceramic support material is selected from a list consisting of: an oxide of silicon, an oxide of aluminium, an oxide of zirconium, an oxide of titanium, an oxide of magnesium, or combinations thereof; preferably an oxide of silicon, an oxide of aluminium, an oxide of zirconium, or combinations thereof; more preferably oxide of silicon, oxide of aluminium or combinations thereof.

30. Reactor according to any of the previous claims, wherein the pipe comprises a first blower or pump.

31. Reactor according to any of the previous claims, wherein the gas outlet comprises a second blower or pump to draws hydrogen from the gas separation membranes module to the gas hydrogen outlet.

32. Reactor according to any of previous claims, wherein the amount of catalyst in the interior of the plurality of channels ranges from 0.1-2 mgcataiyst cm'2, preferably 0.5 -1 mgcataiyst cm'2.

33. Process for continuous methane splitting using the reactor described in any of the previous claims, comprising feeding the reactor with a methane rich flow, wherein the reactor comprises: comprising a gas inlet for methane flow, a reaction chamber for methane splitting, a gas separation membrane for separating and purifying hydrogen, a gas outlet for hydrogen flow, a bottom outlet for carbon removing by gravity from the reaction chamber; wherein the reactor comprises a pipe connecting the reaction chamber to the gas separation membranes module to flow product reaction for introducing the reaction products of the reaction module into the plurality of gas separation membrane, wherein the reaction chamber comprises a reaction module; wherein the reaction module comprises a plurality of gas flow channels with a macro-support; wherein the macro-support comprises a catalytic support loaded with a catalyst for contacting with the methane flow for catalyzing the methane splitting reaction; wherein the said catalyst is a plurality of Ni or Co alloy particles, wherein the operation temperature ranges from about 500 °C to 650 °C; preferably from about 550 °C to 600 °C; adding a shaving-off flow to react at the catalyst interface with deposited carbon, for the catalyst regeneration.

34. Process according to the previous claim, comprising a further previous step of bonding the catalyst particles to the surface of the support by an alumina ceramic or a silica ceramic material or a carbonized thermosetting polymer.

35. Process according to any of the previous claims 33-34, wherein the alumina ceramic or a silica ceramic material or carbonized thermosetting polymer is stable at the operation temperature ranges.

36. Process according to any of the previous claims 33-35, comprising a further second regeneration step where the reactor is stopped and an inert gas flow is introduced for catalyst regeneration, preferably every 15 days to 60 days, more preferably every 20 days to 30 days.

37. Process according to the previous claim, wherein the duration of the second regeneration step ranges from 1 hour to 20 hours; preferably from 4 hours to 8 hours.

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