Process for producing hydrogen

Abstract

A process of producing hydrogen comprising introducing an external feed stream containing hydrocarbons into a reaction chamber and reacting said hydrocarbons in said reaction chamber in a fixed, fluidized or moving bed of solid materials having a particle size of 0.1 mm to 10 mm at a pressure of 1 to 30 bar, a temperature of 500 to 2000°C, a residence time of the hydrocarbon in the range of 0.1 to 50 s, a conversion rate of 20 to 90% and a carbon deposition rate of 0.1 to 15 wt.-% giving a hydrogen-containing product stream containing unreacted hydrocarbons and solid carbon depositing on said bed material, wherein 10 to 95 vol.-% of said hydrogen-containing product stream is directly recycled internally to said reaction chamber, and wherein the total reactor feed stream containing said external feed stream and said internally recycled feed stream has a volume ratio of H2 / C1+ of 0.5 to 25.

Description

[0001] Process for producing hydrogen Description The present invention comprises a process of producing hydrogen comprising introducing an external feed stream containing hydrocarbons into a reaction chamber and reacting said hydrocarbons in said reaction chamber in a fixed, fluidized or moving bed of solid materials having a particle size of 0.1 mm to 10 mm at a pressure of 1 to 30 bar, a temperature of 500 to 2000°C, a residence time of the hydrocarbon in the range of 0.1 to 50 s, a conversion rate of 20 to 90% and a carbon deposition rate of 0.1 to 15 wt.-% giving a hydrogen-containing product stream containing unreacted hydrocarbons and solid carbon depositing on said bed material, wherein 10 to 95 vol.-% of said hydrogen- containing product stream is directly recycled internally to said reaction chamber, and wherein the total reactor feed stream containing said external feed stream and said internally recycled feed stream has a volume ratio of H2 / C1+ of 0.5 to 25. State of the art Hydrogen is considered potentially the cleanest energy in this century against the background of carbon neutrality beside being a fundamental feedstock for a (petro-)chemical industry. Currently in industry, hydrogen is produced by reforming hydrocarbons, such as coal and natural gas, and only to a small extent by electrolysis of water and me- thane pyrolysis. While reforming of hydrocarbons emits huge amounts of CO2, hydrogen production by water elec- trolysis is relatively expensive and requires high amounts of renewable power. The technology of methane pyrolysis requiring much less renewable power than water electrolysis; however, the methane pyrolysis is still under develop- ment. The generic term methane pyrolysis covers a wide range of different process technologies. The best known and most advanced of these are: Plasma pyrolysis, Metal melting / Metal salt melting, Moving or fluidized bed process, Catalytic process, Partial combustion. These technologies differ in the form of the energy used (thermal, electrical, etc.), the process conditions (temperature, pressure, etc.), the catalysts and / or auxiliary materials used, the process flow and the technical readiness level (TRL). In the pyrolysis of hydrocarbons solid carbon is formed. In some concepts the solid carbon is deposited on a solid carrier: CH 409890 and US 2982622 disclose a process of converting hydrocarbon in a high temperature conversion at 980 to 1650° C into hydrogen and high-grade coke by contact with electrically heated solid particles. The use of a fluid- ized or a moving bed is described. For the fluidized bed mode, it is disclosed that a portion of the hot hydrogen may be recycled to serve to heat the incoming feed. This is particularly desirable when the feed is a liquid and thus, the hydrogen recycle serves as a vaporizing medium. In Fig.1 of US 2982622 the conduct of the hot hydrogen, 20, is located above the conduct of the hydrocarbon feed, 19. US 2002 / 7594 discloses a process for production of hydrogen and carbon by thermocatalytic decomposition of hy- drocarbon fuels over carbon-based catalysts in the absence of air and / or water at a temperature of 800 to 1000°C and 1 to 25 atm. Preferably the process is conducted continuously by using a moving or fluidized bed of carbon parti- cles. Off-gas of the hydrogen separation unit, a hydrogen-depleted gas (HDG) consisting of CH4 and C2+ hydrocar- bons, is recycled to the catalytic reactor. WO 2021 / 78614 also discloses a process for production of hydrogen and carbon producing a hydrogen-containing product mixture that is separated in an electrochemical separation process. The cooled residual anode offgas con- taining less than 10% of hydrogen is mixed outside the reactor with the feedgas. WO 2017 / 035269 discloses a process for production of hydrogen and carbon, wherein the rate of deposition and / or dwell time of particles in the fluidized bed growing zone can be adjusted by admission of gas such as hydrogen and / or feedstock hydrocarbon and / or one or more products of combustion through one or more suitable located ports. WO 2013 / 004398 also discloses a methane pyrolysis reaction. Hydrocarbons are introduced into a reaction chamber and are thermally decomposed into carbon and hydrogen in the presence of a carbon-rich granulated material, wherein at least a portion of the thermal energy necessary for the pyrolysis of the hydrocarbon is introduced into the reaction zone via a hot gaseous heat transfer medium. This transfer medium could be the produced hydrogen. WO 2019 / 145279 and WO 2020 / 200522 disclose a reactor and a method for carrying out an endothermic reaction like hydrocarbon pyrolysis by using a moving bed technology with electric heating of solid particles. In order to achieve optimal heat integration the heat capacity flows of the gas and solid particle flows are adapted to each other to generate an upper and a lower heat integration zone. US 2021 / 0331918 discloses a systems and methods for conversion of hydrocarbons to hydrogen by pyrolysis in a conversion reactor that contains a plurality of sequential fluidized beds. The fluidized beds are arranged so that the coke particles forming the fluidized bed move in a counter-current direction relative to the gas feed and / or product. A portion of the hydrogen- containing product gas is used as a pneumatic transport gas for the transport of the recycled coke particles. At the top of the reactor system, the coke particles are separated from the hydrogen-containing prod- uct gas flow in a cyclone separator before feeding the recycled coke particle into the fluidized beds. In reactor concepts, in which the solid carbon is deposited on a solid carrier material, like in fixed, fluidized or moving bed technologies, the deposition of carbon can result in agglomeration of carrier bed particles. Several negative as- pects are associated with this agglomeration process. In fixed bed reactors, stronger agglomerated beds require a more complex, slower procedure to remove the carrier bed. In moving bed reactors, a higher level of agglomeration between carrier particles can lead to a blockage of the solid flow. In fluidized bed reactors, large agglomerates may no longer be fluidized, thus disrupting the fluidization behavior of the bed ending with a collapse of the fluidized bed. Most of the disclosed experiments on methane decomposition are still conducted on a laboratory scale in a batch mode in fixed-bed reactors or fluidized bed for a very short period of time. The main problem to cross the gap be- tween the laboratory scale fixed bed and its industrial implementation in a fixed, moving or fluidized bed is the impact of deposition of carbon and other solids on the process in industrial dimension and time scale. The problem of carbon deposition in the reactor system, especially on the reactor wall, is already mentioned and partly solved in the state of the art: WO 2023 / 057242 discloses a method for carrying out hydrocarbon pyrolysis by using a moving bed technology with electric heating of carbonaceous material having a porosity of 30 to 70 vol.-% and a high carbon content of 99 wt.-% to 100 wt.-% as metal impurities tend to resolve from the carbonaceous material and deposit on the reactor walls and tends to form agglomerates. DE 102019130600 discloses a method for producing hydrogen and carbon from hydrocarbons in an electric heated reactor. An inert gas component is supplied in the reaction cross section between the electrodes. By introducing an inert gas component, a deposition of carbon in this region of the internal reactor wall is prevented. WO 2022 / 081170 discloses a methane pyrolysis process using stacked fluidized beds with electric heating of coke particles. It is disclosed that the electrical heating is performed above the reaction zone in a hydrogen-rich environ- ment in order to reduce, minimize, or eliminate formation of coke on the surfaces of the electrical heater. WO 2024 / 115488 discloses the use of by-products streams from a cracking process as feedstock for a methane py- rolysis process. Said methane pyrolysis is operated in a moving bed reactor at 800 to 1500°C, wherein the by-prod- ucts streams used as feedstock for the methane pyrolysis is preferably diluted by an internal recycled hydrogen prod- uct stream resulting in a total feedstock stream containing 70 to 99 vol.-% hydrogen and 1 to 30 vol.-% hydrocar- bons. No details are disclosed in view of the methane pyrolysis operation parameters of conversion and deposition rate. However, the problem that deposition of carbon can result in agglomeration of carrier bed particles is still unsolved. Hence, for these reactor types a higher level of agglomeration is associated with a reduced production capacity due to longer downtimes and / or reduced carbon deposition rates per cycle to mitigate agglomeration problems. Further- more, the reactor equipment can be damaged in the removal process of a strongly agglomerated carrier bed. The amount of carbon that can be deposited per batch / moving bed cycle without reaching a critical level of agglomer- ation / blockage of the solid flow is a crucial parameter of the overall economics in hydrocarbon pyrolysis. The in- crease in the achievable carbon deposition without reaching a critical level of agglomeration directly increases the production capacity with all associated advantages regarding the Capital Expenditures (CapEx) and the Operational Expenditures (OpEx) of the process. Task It is an object of the present invention to minimize the risk of agglomeration of carrier particles in cracking and pyroly- sis processes. Another objective of the invention is to prevent the formation of filaments on said carrier particles. An- other objective of the present invention is the regulation of a uniform deposition and thus the regulation of the local deposition rate, especially in a fixed and moving bed reactor. Invention The present invention comprises a process of producing hydrogen comprising introducing an external feed stream containing hydrocarbons into a reaction chamber and reacting said hydrocarbons in said reaction chamber in a fixed, fluidized or moving bed of solid materials having a particle size of 0.1 mm to 10 mm at a pressure of 1 to 30 bar, a temperature of 500 to 2000°C, a residence time of the hydrocarbon in the range of 0.1 to 50 s, a conversion rate of 20 to 90% and a carbon deposition rate of 0.1 to 15 wt.-% giving a hydrogen-containing product stream containing unreacted hydrocarbons and solid carbon depositing on said bed material, wherein 10 to 95 vol.-% of said hydrogen- containing product stream is directly recycled internally to said reaction chamber, and wherein the total reactor feed stream containing said external feed stream and said internally recycled feed stream has a volume ratio of H2 / C1+ of 0.5 to 25. Said carbon deposition rate can be defined by reference time unit (see equation 7) or by weight-% (see equation 8), for the feature in the claims, the definition by weight-% is taken so that said carbon deposition rate is defined by: ఛ^^^Δ^^^,^^^ ^^^^^ ^^^ ⋅ ^^^^^^ = 1 ,ௗ^^^^^௧^^^ ^,ௗ^^^^^௧^^^ ^^^^^^ 00 ⋅^^ = 100 ⋅ ^ ^^^^ ^^^^ = 100 ⋅^ ^^^ The reactor specific reference residence time is defined as follows: • Batch / Fixed bed: ^^^^^= total pyrolysis time • Moving bed: ^^^^^= mean residence time of carrier particles in the pyrolysis zone • Fluidized bed: ^^^^^= mean residence time of carrier particles in the pyrolysis zone Reference mass in [g]: • Batch / Fixed bed: ^^^= initial carrier mass in the reactor •Moving bed: ^^^ = ^^^^^^^^^^^^ ∙ ^^^^^ with the reactor inlet mass flow ^^^^^^^^^^^^ in [g / s]• Fluidized bed: ^^ = ^^^^^ ∙ ^^ with the reactor in ^^^^ ^^^^^^^ ^^^ let mass flow ^^ ^^^^^^^ in [g / s]wherein^^^ ^,ௗ^^^^^௧^^^ is the carbon deposition rate [g / s]Δ^^^ is the carbon mass increase [g]^^^^^ is the reactor specific reference residence time [s]^^^^^^^ is the carbon deposition rate (per reference time unit) in [wt.%].Surprisingly, it was found that the risk of local high deposition rates leading to agglomeration can be controlled by a moderate carbon deposition rate, a moderate conversion rate and a volume ratio of H2 / C1+ of 0.5 to 25. In addition, it was found that the presence of C2+ components in the gas phase promotes the formation of an unde- sired type of pyrolytic carbon in the form of filaments. In some installments, for example in moving bed reactors, fixed bed reactors or fluidized bed reactors, the formation of this filamentous pyrolytic carbon is associated with a higher level of agglomeration of particles. In contrast, no formation of filaments is observed if the concentration of C2+ com- ponents is low. The concentration of C2+ components can be reduced in different ways: Firstly, the C2+ concentration of the fresh hydrocarbon feed can be reduced by purification, secondly, the C2+ concentration of the fresh hydrocarbon feed could be reduced by blending with a co-feed. The first option, the usage of a hydrocarbon feed with low C2+ concentrations essentially mean the usage of a high purity methane stream. This can either be realized by using high purity methane side streams of a chemical produc- tion side, like the methane cut of a stream cracker, or by purifying a hydrocarbon stream, for example natural gas, by means of a separation unit. Hence, such low C2+ streams are associated with additional costs, process steps and energy demand compared with the direct usage of non-purified hydrocarbon sources with higher C2+ concentrations like natural gas. It was found that blending the non-purified external hydrocarbon feed with parts of the hydrogen-containing product stream in a recycle mode reduces the ratio of C2+ components at the inlet of the reactor and thus reduces the for- mation of the undesired type of pyrolytic carbon in the form of filaments. In addition, a reduced conversion rate and reduced carbon deposition rate reduces the amount of C2+ components formed during the pyrolysis reaction. The conversion rate and the carbon deposition rate can be adjusted by temperature, by pressure, by dilution of the feed- stock with an inert gas like hydrogen, by reference time, and / or by porosity of the carrier. By reducing C2+ concentrations, associated problems like agglomeration can be reduced, a thus higher amounts of carbon can be deposited increasing the space time yield and thus the overall profitability of the process. None of the cited documents mention the influence of C2+ components either in the feedstock or formed in the pyrol- ysis reaction. In addition, none of the documents disclose to use a part of the produced hydrogen-containing product stream in a direct internal recycle mode for regulation both a reduced conversion rate and carbon deposition rate and a reduced amount of C2+ components in the feedstock. Definition C1 stands for methane. C1+ stands for all hydrocarbons having one or more C-atoms in the molecule. C2+ stands for all hydrocarbon having two or more C-atoms in the molecule. The (total) reactor feed stream includes both the external feed stream and the internally recycled hydrogen-contain- ing product stream. Detailed Description of the Invention Cracking or pyrolysis process of hydrocarbons The process of cracking or pyrolyzing hydrocarbons to produce hydrogen is well described in the state of the art. Typically the pyrolysis and cracking operate at 500 to 2000°C in a pressure range from atmospheric pressure to 30 bar. Typical hydrocarbon feedstocks for pyrolysis processes are gaseous hydrocarbons, like natural gas, methane, ethane, biogas, like biomethane. Typical feedstocks for cracking processes are gaseous and liquid hydrocarbons like ethane, propane, naphtha, bionaphtha. Methane Pyrolysis Preferably, a methane pyrolysis is conducted. The methane pyrolysis is conducted in a moving, fluidized or fixed bed reactor, preferably at temperatures ranging from 500 to 2000°C, preferably ranging from 800 to 1800°C, even more preferably ranging from 1000 to 1600°C, even more preferably ranging from 1100 to 1500°C, even more preferably ranging from 1150 to 1450°C and at pres- sures ranging from 1 to 30 bar, preferably from 2.5 to 25 bar, even more preferably 5 to 20 bar, particularly 5 to 15 bar. The conversion rate is preferably 25 to 85 %, more preferably 30 to 80 %, even more preferably 50 to 80 % (based on the hydrocarbons in the feed stream, see equation 3). The carbon deposition rate of solid pyrolytic carbon is preferably 0.1 to 15 wt.-%, more preferably 2 to 12 wt.-%, even more preferably 3 to 10 wt.-%, even more preferably 4 to 9 wt.%, even more preferably 4 to 8 wt.-%. The definition of the carbon deposition rate, sometimes also referred to as carbon growth rate, depends on the reac- tor type and is defined in the following: The symbols and indices used are shown in Table 1. Table 1. Symbols and indices used for describing carbon deposition and growth rates. Symbol Description^^^^,ௗ^^^^^௧^^^ Carbon deposition rate [molC / s]^^^^ுସ,^ Molar flow [molC / s]^^ Conversion rate [-]^^^ ^,ௗ^^^^^௧^^^ Carbon deposition rate [g / s]^^^ Molar mass [g / mol]^^^^ Carbon deposition [g / g]Δ^^^ Carbon mass increase [g]^^^ Carbon mass [g]^^^^^ Carbon deposition rate in [g / g*s]τ Time unit [s]^^^^^^^ Carbon deposition rate (per reference time unit) [-] or in [wt.%]V Reactor volume [m3]Indices Description 0 Particle before pyrolysis or state feed gas 1 Particle after pyrolysis or state in pyrolysis gas in Reactor inlet ref Reference (e.g. time unit) The carbon deposition rate in [mol,C / s] results from the molar flow of CH4 in the feed gas and the conversion rate: ^^^^,ௗ^^^^^௧^^^ = ^^^^ுସ,^ ⋅ ^^(^^,^^,^^^ , ^^^^^^%ுଶ,^, ^^^^^^%^ுସ,^) 1 Considering the molar mass of atomic carbon leads to the carbon deposition rate in [g,C / s]: ^^^ ^,ௗ^^^^^௧^^^ = ^^^ ^,ௗ^^^^^௧^^^ ⋅ ^^^ 2 The CH4 conversion rate is calculated using the molar flow of CH4 in the feed and in the pyrolysis gas: ^^^^ − ^^^ ^^^^^ = ுସ,^ ^ுସ,^ = ^ுସ,^ 3^^^ 1 −^ுସ,^ ^^^ ^ுସ,^ The carbon deposition in [g / g] is defined as: Δ^^^^^^ = ^ 4^^^ The carbon deposition rate in [g / (g*s)] follows from the initial mass and the carbon deposition rate in [mol,C / s]: ^^^^^^^^ = ^,ௗ^^^^^௧^^^ 5^^^ The total increase in carbon mass on zone: Δ^^^ = ^^^ ^,ௗ^^^^^௧^^^ ⋅ τ 6 Considering the reactor volume ^^ and the reactor type specific reference residence time leads to the carbon deposi- tion rate per reference time unit in [-]: ఛ^^^Δ^^ ^^^^ ^^^ ⋅ ^^^^^^ = ^,^^^ = ^,ௗ^^^^^௧^^^ ^,ௗ^^^^^௧^^^ ^^^ 7^^^ ^ ^^^^ ^^^^ = The carbon deposition rate per reference time unit can also be expressed in [wt.%] as follows: ఛ^^^Δ^^ ^^^^ ^^ ^,^^^ ^,ௗ^^^^^௧ ^^ ⋅ ^^^^^ ^^^ ^,ௗ^^^^^௧^^^ ^^^ 8^^^ = 100 ⋅^^ = 100 ⋅ ^ ^^^^ ^^^^ = 100 ⋅^ ^^^ The reactor specific reference residence as • Batch / Fixed bed: ^^^^^= total pyrolysis time • Moving bed: ^^^^^= mean residence time of carrier particles in the pyrolysis zone • Fluidized bed: ^^^^^= mean residence time of carrier particles in the pyrolysis zone Reference mass in [g]: • Batch / Fixed bed: ^^^= initial carrier mass in the reactor •Moving bed: ^^ = ^^^^^ ∙ ^^ wit ^^^ ^^^^^^^ ^^^ h the reactor inlet mass flow ^^^ ^^^^^^^ in [g / s]• Fluidized bed: ^^ ∙ ^^^^^ with the reactor inlet mass in [g / s] Depending on the available methods, different ways for the practical application of the described theoretical framework for the calculation of ^^^^^^^exist. It is important to note that the specific material property must be measured before the pyrolysis (index “0”) and after a pyrolysis time of ^^^^^(index “1”). Exemplary ways of application: • Using the Sauter mean diameter of the particle size distribution and the measured real density from helium or xylene pycnometer:^^^^ = ^^ଷ ^^⋅ ^^ , ^^ ଷ^ 6 ௌ^௨௧^^,^ ு^,^ ^ =6 ^^ௌ^௨௧^^,^ ⋅ ^^ு^,^9 •Using a reference volume ^^^^^ which contains the particle mass ^^^௨^^,^,^ to determine the bulk density:^^ ^^^௨^^,^ ^^^௨^^,^^௨^^,^ =^^ , ^^^௨^^,^ =^^^ ^^^^^11 Δ^^^,^^^ = ^^^^^ ⋅ (^^^௨^^,^ − ^^^௨^^,^) 12 • Using the measured or calculated initial carrier weight ^^^, using gas analytics for the calculation of conver- sion rate, using the measured of calculated molar flow rates and applying the detailed formulas given above. Moving bed: The solid carbon mass from hydrocarbon pyrolysis deposited on one solid granule in one pass through the reactor divided by the mass of the said one solid granule before entering the reactor. The mass of the said solid particle before and after the pyrolysis is calculated by the Sauter mean diameter of the particle size distribution and the real density of the used granule. The reference residence time according to equations 7-8 is preferably between 0.1 and 15 h, preferably between 1 and 10 h and more preferably between 2 and 8 h. Fixed bed: The solid carbon mass from hydrocarbon pyrolysis deposited on one solid granule during its residence time in the reactor divided by the mass of the said one solid granule before entering the reactor. The mass of the said solid particle before and after the pyrolysis is calculated by the Sauter mean diameter of the particle size distribution and the real density of the used granule. The reference residence time according to equations 7-8 is preferably 0.1 to 90 h, more preferably 0.25 to 50 h, more preferably 0.5 to 25 h and even more preferably 1.0 to 10 h. Fluidized bed: The solid carbon mass from hydrocarbon pyrolysis deposited on one solid granule during one inner circulation through the reactor divided by the mass of the said one solid granule before entering the reactor. The mass of the said solid particle before and after the pyrolysis is calculated by the Sauter mean diameter of the particle size distribution and the real density of the used granule. In a fluidized bed, one particle can circulate multiple times through the reactor during its residence time inside the reactor. The reference residence time according to equations 7-8 is preferably 0.1 to 90 h, more preferably 0.25 to 50 h, more preferably 0.5 to 25 h and even more preferably 1.0 to 10 h. The methane pyrolysis process can be heated in different ways known to the persons skilled in the art: via plasma, microwave, heated carrier gas, resistance heating, induction, liquid metal processes or autothermal. Preferably the methane pyrolysis process is heated electrically, even more preferably by resistive heating (Joule heating) of the solid material as described for example in US 2982622, WO 2019 / 145279 and WO 2020 / 200522. Reactor feed stream The feed of the cracking or pyrolysis process contains at least two feed streams: (i) external feed stream containing hydrocarbons, preferably gaseous and liquid hydrocarbons, more preferably gaseous hydrocarbons, even more pref- erably methane and / or other light hydrocarbons feed stream, preferably natural gas and (ii) an internally recycled hy- drogen-containing product stream. External feed stream Typically, natural gas is used as external feed stream. The composition of natural gas is very depending on the loca- tion of extraction. Beside methane, other hydrocarbon gases like ethane, propane, and butane may be present in varying amounts. Additionally, the presence of impurities such as carbon dioxide, nitrogen, and sulfur compounds can impact the quality and usability of natural gas. Natural gas typically has the following composition: methane from 50 to 99 Vol.-%, preferably 85 to 99 Vol.-%, ethane from 0.5 to 25 Vol.-%, preferably from 1 to 15 Vol.-%, propane from 0 to 10 Vol.-%, preferably from 0.5 to 5 Vol.-%, butane from 0 to 10 Vol.-%, pentane from 0 to 5 Vol.-%, CO2 from 0 to 5 Vol.-% and N2 from 0 to 10 Vol.-%, preferably from 1 to 5 Vol.-%, sulfur components from 0 to 10 Vol.-%, preferably from 0 to 2 Vol.-%, of the total raw natural gas. Typically, natural gas contains hydrocarbons with a ratio of C1 / C2+ of 3 to 100, more preferably 3 to 50, even more preferably 3 to 25, even more preferably 4 to 20, even more preferably 5 to 15. In addition or alternatively to natural gas, synthetic methane, biogas, bio methane, flare gas and / or industrial off gas, e.g. cracker off gas, can be used. Such external feed stream also preferably contains hydrocarbons with a ratio of C1 / C2+ of 3 to 100, more preferably 3 to 50, even more preferably 3 to 25, even more preferably 4 to 20, even more preferably 5 to 15. Internally recycled hydrogen-containing product stream The methane pyrolysis results in a hydrogen-containing product stream. Depending on the reaction conditions, the hydrogen content of the hydrogen-containing product stream is preferably between 30 vol% and 99 vol%, more pref- erably between 50 vol% and 98 vol%, and in particular between 60 vol% and 98 vol%, related to the total volume of the hydrogen-containing product stream. The content of methane is preferably between 1 and 60 vol%, more preferably between 2 and 40 vol% and in particu- lar between 2 and 30 vol%. The sum of the contents of all C2+ hydrocarbon components comprising e.g. C2H6, C2H4, C2H2, C3H8, C3H6, C3H4, C4H8, C4H6, C6H6, C7H8, C8H10 in the hydrogen-containing product stream is preferably between 0 and 1 mol%, more preferably between 0 and 0.5 mol% and in particular between 0 and 0.1mol%. Preferably, the content of C2H6 is preferably between 0.1 and 5000 ppm, more preferably between 0.1 and 1000 ppm and in particular between 0.1 and 500 ppm. Preferably, the content of C3H8 is preferably between 0.1 and 3000 ppm, more preferably between 0.1 and 1000 ppm and in particular between 0.1 and 200 ppm. Preferably, the content of C2H4 is preferably between 0.1 and 5000 ppm, more preferably between 0.1 and 1000 ppm and in particular between 0.1 and 500 ppm. Preferably, the content of C2H2 is preferably between 0.1 and 5000 ppm, more preferably between 0.1 and 1000 ppm and in particular between 0.1 and 500 ppm. Preferably, the content of C6H6 is preferably between 0.1 and 3000 ppm, more preferably between 0.1 and 1000 ppm and in particular between 0.1 and 200 ppm. The nitrogen content of the hydrogen-containing product stream is preferably between 0 and 20 vol%, more prefera- bly between 0 and 10 vol% and in particular between 0 and 5 vol%. The carbon monoxide content of the of the hydrogen-containing product stream is preferably between 0 and 2 vol%, more preferably between 0 and 1 vol% and in particular between 0 and 0.5 vol%. The carbon dioxide content of the of the hydrogen-containing product stream is preferably between 0 and 2 vol%, more preferably between 0 and 1 vol% and in particular between 0 and 0.5 vol%. The water content of the hydrogen-containing product stream is preferably between 0 and 2 vol%, more preferably between 0 and 1vol% and in particular between 0 and 0.5 vol%. The hydrogen-containing stream preferably contains hydrocarbons with a ratio of C1 / C2+ of 4 to 10000, more prefer- ably 15 to 1000, even more preferably 20 to 500, even more preferably 30 to 250, even more preferably 50 to 200. Recycling the hydrogen-containing product stream The hydrogen-containing product stream leaving the reaction chamber preferably enters a separation unit. This sepa- ration unit typically includes a valve, a flow meter, and a meter to determine the H2 concentration. Preferably, 20 vol% to 90 vol% of the hydrogen-containing product stream is separated and recycled to the reaction chamber, more preferably 30 vol % to 85 vol %, more preferably 40 vol % to 85 vol %, even more preferably 50 vol % to 80 vol %. The vol% of the hydrogen-containing product stream taken for the internal recycle is preferably adjusted to a ratio H2 / C1+ ratio of the total reactor feed stream ranging from 0.5 to 25, more preferably from 0.75 to 10, even more preferably from 1 to 8, even more preferably 2 to 6. Preferably, a direct recycle of the hydrogen-containing product stream is used, that means, a portion of the (raw) hy- drogen-containing product stream is recycled without any purification steps. Said recycled hydrogen-containing prod- uct stream may or may not be cooled, heated and / or compressed. The recycled hydrogen-containing product stream is mixed with the external feed stream, e.g. natural gas, inside or outside the reaction chamber. Before mixing, the temperature of the internally recycled hydrogen-containing product stream is preferably between 10°C and 200°C, even more preferably between 15 and 150°C and in particular be- tween 20 and 100°C. Preferably, the external feed stream and the recycled hydrogen-containing product stream are mixed before entering the reaction chamber, outside of the reactor. Alternatively, the streams are separately intro- duced into the reaction chamber. Ratio of internally recycled hydrogen-containing product stream and external hydrocarbon stream The (total) reactor feed stream includes both the external feed stream and the internally recycled hydrogen-contain- ing product stream. Preferably the ratio of the volume of the internally recycled hydrogen-containing product stream to the volume of the external feed stream is between 0.1 to 20, more preferably between 0.3 and 8, even more preferably between 0.8 and 6. Typically, the total reactor feed stream contains preferred 50 to 95 Vol.-% hydrogen, more preferred 65 to 85 Vol% hydrogen, 5 to 50 Vol.-% hydrocarbons, more preferred 35 to 15 vol% hydrocarbons, and 0 to 3 Vol.-% inert gases, particularly nitrogen and argon. The total reactor feed stream preferably contains hydrocarbons with a ratio of C1 / C2+ of 15 to 500, more preferably 17 to 100, even more preferably 20 to 50. Preferably, said volume ratio of the total reactor feed stream is at least increased by a factor of 1.1 to 2.5, more pref- erably 1.25 to 2, compared to said volume ratio of C1 / C2+ of the external feed. Preferred Process Conditions: Moving Bed Preferably the solid material (also called substrate) is guided in form of a moving bed through the reaction chamber, with methane and / or other light hydrocarbons being passed advantageously in countercurrent to the substrate. For this purpose, the reaction chamber is preferably designed as a vertical reactor, which means that the movement of the moving bed is gravity driven. Flow through the moving bed is taking place, advantageously, homogeneously, and uniformly (see for example WO 2013 / 004398, WO 2019 / 145279 and WO 2020 / 200522). Methane and / or other light hydrocarbons are preferably introduced via the bottom of the reactor, preferably having a temperature of 10 to 200°C. The substrate is preferably introduced via the top of the reactor, preferably having a temperature of 10 to 200°C. The hydrogen-containing product gas is preferably taken off via the top of the reactor, preferably having a temperature of 10 to 200°C. The granular pyrolytic carbon is preferably taken off via the bottom of the reactor, prefer- ably having a temperature of 10 to 200°C. The discharged carbon is preferably at least partly recycled and introduced into the reactor again using it as sub- strate for the moving bed. As a result, a continuous process of the moving bed is achieved. The recycled carbon can optionally be mixed with fresh, initial solid material. The flow velocity of the substrate is advantageously in the range of 0.005 to 6.0 cm / s, more preferably 0.005 to 0.5 cm / s. The flow velocity of the gas flow is advantageously in the range of 0.025 to 10 m / s, more preferably 0.025 to 2 m / s. The gas residence time in the reactor is advantageously between 0.1 and 50 s, preferably between 1 and 20 s. The residence time of the substrate is preferably between 0.1 and 15 hours, preferably between 1 and 10 hours and more preferably between 2 and 8 hours. Solid material The solid material, also called substrate or carrier material, is preferably a granular material. The particle size of a preferred substrate is in the range of 0.1 to 10 mm, preferably 0.3 to 8 mm. The preferred substrate is a carbon-containing substrate, for example calcined petcoke (CPC), acetylene coke, green coke, anode butts, bio char, recycled or recovered carbon black (from polymer recycling via pyrolysis), black mass (battery recycling) or the pyrolytic carbon itself. The carbon content is preferably in the range of 70 to 99 wt.%, more preferably 90 wt.% to 99.9 wt.%, more preferably 99 to 99.95 wt.%, even more preferably 99 to 99.99 wt.%. The oxygen content is preferably in the range of 20 to 0.001 wt.%, more preferably 5 wt.% to 0.1 wt.%, more prefera- bly 1 wt % to 0.5 wt%. The ash content, in particular Fe, Ni, Ca, Si, Al, Cr, K, Mg, Mn, Pb, V and Zn, is preferably in the range 20 to 0.001 wt.%, more preferably 5 wt.% to 0.1 wt.%, more preferably 1 wt % to 0.5 wt%. The Sulphur content is preferably in the range 5 to 0.0001 wt.%, more preferably 0.5 wt.% to 0.01 wt.%, more prefer- ably 0.1 wt % to 0.01 wt%. The chlorine content is preferably in the range 1 to 0.00001 wt.%, more preferably 0.1 wt % to 0.001 wt%, even more preferably 0.01 wt.% to 0.1 wt.%. Preferably, the metal-content on the outer surface (defined by the outer shell of 5 µm thickness), is preferably in the range 20 to 0.001 wt.%, more preferably 5 wt.% to 0.001 wt.%, more preferably 1 wt % to 0.001 wt%, even more preferably 0.5 wt.% to 0.001 wt.%. The BET surface area of the substrate is preferably between 0.001 and 100 m2 / g, preferably 0.01 and 50 m2 / g, in particular 0.05 to 30 m2 / g. Typically, the porosity of the solid material is between 0.1% to 90%, preferably between 1% to 80%, even more preferably 30% to 70% according to WO 2023 / 057242, even more preferably 30% to 60% (Hg porosimetry, DIN66133). Preferably, the density of the substrate is in the range of 1.5 to 2.5 g / cc (real density in xylene, ISO 8004). Preferably, the bulk density of the substrate is in the range of 0.5 to 1.5 g / cc. Pyrolytic granular carbon Typically, the density of the deposited pyrolytic carbon produced via the described methane pyrolysis process is in the range of 1.5 to 2.5 g / cc, preferably 1.7 to 2.4 g / cc, even more preferably 2.0 to 2.3 g / cc (real density in xylene, ISO 8004). Typically, the bulk density of the pyrolytic carbon is in the range of 0.5 to 1.5 g / cc, more preferably 0.7 to 1.3 g / cc. Typically, the ash content of the pyrolytic carbon is in the range of 0.001 to 1 weight-%, preferably 0.01 to 0.2 weight- %. Typically, the carbon content is in the range of 98 to 100 weight-%, more preferably 99.5 to 100 weight-%, even more 99.75 to 100 weight-%, even more 99.9 to 100 weight-%. Typically, the porosity of the granular pyrolytic carbon is between 0% to 15%, preferably 0.2% to 10%, most prefera- bly 0.2% to 5% (Hg porosimetry, DIN66133). Typically, the specific surface area of the pyrolytic carbon measured by Brunauer-Emmett-Teller (BET) is in the range of 0.001 to 20 m2 / g, preferably 0.001 to 10 m2 / g, even more prefera- bly 0.05 to 5 m2 / g. Preferred Moving Bed Method Preferably a moving bed technology is used for the methane pyrolysis of the present invention comprising: - a plurality of solid substrates is guided into the first heat integration zone and from there into the reaction zone, - the solid substrates are heated in the reaction zone, - the solid substrates are guided from the reaction zone into the second heat integration zone and are with- drawn from the second heat integration zone, - methane is introduced into the second heat integration zone and from there into the reaction zone, wherein methane in the second heat integration zone is heated against solid substrates coming from the reaction zone, wherein the solid substrates are cooled, and wherein the methane is contacted with the heated solid substrates in the reaction zone, wherein heat from the heated solid substrates is transferred to methane in order to heat methane in the reaction zone, wherein methane is decomposed to hydrogen and granular pyro- lytic carbon in the reaction zone, - hydrogen produced is guided from the reaction zone into the first heat integration zone, wherein the solid sub- strates in the first heat integration zone are preheated against hydrogen coming from the reaction zone, wherein hydrogen is cooled, and wherein - hydrogen is withdrawn from the first heat integration zone - granular pyrolytic carbon produced is deposited on solid substrates and withdrawn with the substrates - withdrawn substrate is preferably at least partly recycled and introduced into the reactor - withdrawn hydrogen is preferably at least partly recycled and introduced into the reactor. Preferred Reactor Preferably, the reactor for carrying out the methane pyrolysis, in which hydrogen and pyrolytic carbon are produced from hydrocarbons, preferably methane, comprises: - a reactor surrounding a reactor interior - the reactor is configured to provide a gravity-driven moving bed in a reaction zone of the reactor interior, which gravity-driven moving bed comprises a large number plurality of solid substrates, wherein the reactor is also configured to guide methane into the reaction zone, wherein, in order to heat methane, the reactor is configured to heat the solid substrates in the reaction zone by generating an electric current in the solid sub- strates between a pair of first and second electrodes such that, by transferring heat from the solid sub- strates to methane, - methane in the reaction zone can be heated to a reaction temperature to produce hydrogen and granular pyrolytic carbon, - and wherein the reactor interior also comprises a first heat integration zone in which heat from hydrogen produced in the reaction zone can be transferred to solid substrates of the reactor gravity driven moving bed which are to be guided into the reaction zone, and wherein the reactor interior also comprises a second heat integration zone in which heat from solid substrates of the reactor gravity driven moving bed coming from the reaction zone can be transferred to methane in order to preheat methane, wherein said reaction zone is arranged between said pair of first and second electrodes and said first heat integration zone is arranged above said first electrode and said second heat integration zone is arranged below said second electrode. The volume of the reaction section is preferably 1 m3 to 1000 m3, preferably 5 m3 to 750 m3, more preferably 0.5 m3 to 500 m3. The height of the reaction section is preferably 0.1 m to 50 m, preferably 0.5 to 20 m, more preferably 1 m to 10 m. The said moving bed reactor has several advantageous compared to other reactor types. - Heat integration enables a high efficiency since gas and solid are leaving the reactor at temperatures be- tween 10 °C – 200 °C - No additional equipment for heat recovery or special equipment for handling of gas solids with temperatures above 200 °C required. - Reproduceable carbon quality: Each particle is treated under the same conditions in the moving bed. Same residence time for each solid particle in the reaction zone. Preferred Fixed Bed Method Preferably a fixed bed technology is used for the methane pyrolysis of the present invention comprising: - a plurality of solid substrates having a particle size of 0.5 to 10 mm is introduced into the reactor and fixed in the reaction zone, preferably the substrates are introduced via the top of the reactor, - the solid substrates are heated in the reaction zone, preferably the solids are electrically heated, more prefer- ably the solids are heated via a direct electric Joule heating, - methane is introduced into the reaction zone, wherein the methane is contacted with the heated solid sub- strates in the reaction zone, wherein heat from the heated solid substrates is transferred to methane in order to heat methane in the reaction zone and wherein methane is decomposed to hydrogen and pyrolytic carbon that is deposited on the solid substrates in the reaction zone, preferably methane is introduced via the top of the reactor, - hydrogen is withdrawn from the reaction zone, preferably from the bottom of the reactor - By feeding hydrocarbon from the top into the reactor, the gas velocity is not limited by the fluidization regime of the solid particles, which allows higher space time yields compared to moving bed or fluidized bed reactors. Gas velocities in the reaction zone (calculated for an empty reactor) are 0,5 – 15 m / s and more preferred 1 – 10 m / s. In addition, the residence time of the substrate can be increased to 2 to 24 hours. - batchwise the solid substrates with pyrolytic carbon deposited on it are withdrawn from the reaction zone - withdrawn substrate is preferably at least partly recycled and introduced into the reactor - withdrawn hydrogen is preferably at least partly recycled and introduced into the reactor. Fixed bed reactors are known in the state of the art, e.g. [1] Eigenberger, Ruppel, Ullmanns’s encyclopedia of indus- trial chemistry Catalytic Fixed-Bed Reactors, 2012 Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim. Typically, several fixed bed reactors are combined in a cyclic mode operation e.g. One reactor is in production mode and carbon is deposed in the fixed bed until the maximum carbon deposition is reached. While the carbon is re- moved from this reactor, the next reactor starts with production and so on. With this cyclic operation mode, a continu- ous production is ensured. Preferred Fluidized Bed Method Preferably a fluidized bed technology is used for the methane pyrolysis of the present invention comprising: - a plurality of solid substrates having a particle size of 0.1 to 10 mm is introduced in the fluidization zone, the reaction zone, and placed on a distributor o optionally the substrates are heated externally and are introduced into the reactor at a temperature of 800 °C to 1500 °C - optionally the solid substrates are heated in the reaction zone, preferably the solids are electrically heated, - methane is introduced into the reactor via the bottom, distributed by the distributor fluidizing the solid sub- strates and guided through the reaction zone wherein methane is contacted with the heated solid substrates in the reaction zone, wherein heat from the heated solid substrates is transferred to methane in order to heat methane in the reaction zone, wherein methane is decomposed to hydrogen and pyrolytic carbon that is de- posited on the solid substrates in the reaction zone, o The Hydrocarbon concentration of the feed gas ranges preferably from 20 vol% - 100 vol %, more prefera- bly form 60 vol% to 100 vol % o Due to the relative motion of the particles in a fluidized, carbon deposition rates from the pyrolyzed gas on the solid granule per pass through the reactor are preferred above 10 w% and more preferred above 20 w%. The relative deposition is related to the mean solid particle mass flux out of the reactor. o The pressure is preferably 1 to 10 bar. - hydrogen is withdrawn from the reaction zone, preferably from the top of the reactor - preferably continuously, the solid substrates with pyrolytic carbon deposited on it are withdrawn from the reac- tion zone - withdrawn substrate is preferably at least partly recycled and introduced into the reactor - withdrawn hydrogen is preferably at least partly recycled and introduced into the reactor. Fluidized bed reactors are known in the state of the art, e.g[2]. Werther, Ullmanns’s encyclopedia of industrial chem- istry Fluidized Bed-Reactors, 2012 Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim. A fluidized bed reactor may have more than one stage as described in WO2022081170 and US20210331918. H2 Purification The part of the hydrogen-containing product stream that is not internally recycled into the reaction chamber is typi- cally purified in a removal apparatus, e.g. using adsorption or membrane technologies. For example, the gaseous product stream can still contain hydrogen sulfide. Hydrogen sulfide can preferably be re- moved from the gaseous product stream via gas scrubbing, e.g. via ZnO, CuZnO, Fe(OH)3, Zeolites, MOFs, as known in the state of the art. Typical process conditions of the gas scrubbing are 60 to 180°C and 1 to 100 bar. For example, the gaseous product stream can still contain carbon oxide and inert gases, e.g. nitrogen. Correspond- ing hydrogen recovery apparatuses are common knowledge to the person skilled in the art, for example pressure swing adsorption or permeation. Recycle of purification off-gas Preferably, a part of the off-gas from the hydrogen recovery is additionally internally recycled to the reaction cham- ber, preferably 5 to 100 vol.-% of the total off-gas, more preferably 10 to 98 vol.-%, even more preferably 25 to 95 vol.-%, even more preferably 50 to 95 vol.-% even more preferably 70 to 95 vol.-% even more preferably 80 to 95 vol.-%. Preferably, the remaining part of said total off-gas from the hydrogen recovery is withdrawn. The preferred percentage of recycle and withdrawal is dependent on the amount of inert gases in the hydrogen-containing product stream, especially on the nitrogen content. Advantages By the present invention, the formation of carbon filaments in the cracking process is significantly reduced or even prevented. The formation of filaments is associated with a higher risk of agglomeration or a reduced level of agglom- eration of the carrier particles in the cracking process. Hence, by reducing or preventing the filament formation, the risk of agglomeration of particles is strongly reduced, which allows the deposition of higher amounts of carbon per batch / cycle without reaching a critical level of agglomeration / blockage of the moving beds solid flow. The amount of deposited carbon determines the production rate in methane pyrolysis and is thus a crucial parameter of the overall economics of a commercial methane pyrolysis process. In a fixed bed reactor, the amount of deposited carbon di- rectly limits the maximum production capacity per batch, while it determines the continuous carbon production rate for a given carrier flow rate in a moving bed reactor or fluidized bed reactor. Hence, the increase in the achievable carbon deposition without reaching a critical level of agglomeration directly increases the production capacity with all associated advantages regarding the capital expenditures and the operational expenditures of the process. Further- more, by the pretreating process also initially not preferred materials can be prepared for the usage as carrier mate- rial in methane pyrolysis. For example, carbon-based materials with increased metal contents and / or increased oxy- gen contents (for example bio-based carbon, low grade CPC, anode butts, green coke etc.) can be pretreated and thus be made available for the usage as carrier materials in cracking processes of hydrocarbons. This way low-cost alternatives to high purity carbon carriers can be upcycled for cracking processes of hydrocarbons and production loops in customer markets producing low grade carbons as unwanted side products can be closed, having ad- vantages both regarding commercial as well as sustainability aspects. Description of the Figures Figure 1 presents the block diagram of a general embodiment of the process of the invention (1a), a moving bed re- actor (1b), a fixed bed reactor (1c), cited in [1], and a fluidized bed reactor (1d), cited in [2]. Figure 2a and 2b presents SEM pictures of carbon bed particles after pyrolysis of a hydrocarbon feed with a volume ratio C1 / C2+ of 12.25 (Experiment 1, (2a)) compared to a hydrocarbon feed containing solely methane (Experiment 2, (2b)). Figure 3 presents the observation of the continuous operation by the filling level of the carbon bed particles of the reactor. Figure 4a: The diagram shows the local carbon deposition along the reactor height for four CH4 concentrations in the feed gas, all at 1250°C in the reaction zone. Figure 4b: Normed reaction rate along the reactor height for four CH4 concentrations in feed gas and 1250°C in the reaction zone. Figure 4c: The diagram shows the local carbon deposition along the reactor height for 50% CH4 in the feed gas and different temperatures. Figure 5: The diagram shows the isolines of carbon deposition (GR) depending on temperature and CH4 concentra- tion in feed gas, where CDi < CDii < CDiii.

[0002] Example 1. Impact of carbon deposition rate on agglomeration Methane pyrolysis experiments were carried out in a fix bed lab reactor. The reactor had an inner ceramic reaction tube with an inner diameter of 50 mm and a height of 560 mm. The tube was filled with calcinated petroleum coke (CPC) with a particle size distribution of 2-4mm. The exact mass and the filling height of the CPC were evaluated before the start of the experiment to calculate the initial bulk density. For the experiment the Reactor was wall heated to 1200 °C and the reactor was fed with 100 % CH4. The experi- ment was carried out for 75 min while the conversion rate of methane was 80 vol%. The produced solid carbon was deposed on the initially filled CPC. The determination of carbon deposition was done by offline carbon analysis. The carbon bed was removed in six fractions (according to height) from the cold reactor tube after the experiment. The bulk density of each fraction was evaluated according to the mass of the fraction and its volume. The bulk density before and after the experiment was used to calculate the carbon deposition rate according to equations 8, 11 and 12. While the fractions were removed it was evaluated: if the reactor was blocked in the respective fraction or if the carbon particles didn’t flow easily out of the reactor, the fraction was assessed as blocked. In Table 1, the results are summarized. With a carbon deposition of 18 wt-% the bed was blocked and with 12 wt-% the bed was loose and carbon could easily be removed out of the reactor. Table 1: Impact of carbon deposition rate on agglomeration Position Reactor above Local carbon deposition Bed conditions inlet [mm] rate [wt %] 125 mm 18 Blocked by agglomeration 290 mm 12 Loose 458 mm 2 Loose 2. Impact of temperature, H2 dilution, pressure, reference time, and carrier porosity on carbon deposition rate Methane pyrolysis experiments were carried out in the fixed bed lab reactor described in Example 1. Several experi- ments with parameter variation were performed to evaluate their influence on carbon deposition rate; the results are summarized in Table 2. As stated in the description, the main parameters influencing the carbon deposition rate are temperature, pressure, H2 dilution, reference time, and carrier porosity; therefore, these were varied within the ranges allowed by the lab reactor. For experiments nr.1 – 9, the reactor was filled with particles consisting of calci- nated petroleum coke (CPC) core and pyrolytic carbon shell, with a particle size distribution of 2-4 mm and low po- rosity, reflected on its low specific surface area <0.1 m2 / g. In experiment nr.10, a carrier was used, consisting of pure CPC without pyrolytic carbon, and having a higher porosity than the carrier in experiments nr.1 – 10, reflected in its higher specific surface area of 11.7 m2 / g. In all experiments shown, the bed condition afterwards was “loose”. From the results, the following relations are seen: • An increasing temperature leads to an increase in carbon deposition rate. • An increasing H2 / C1 ratio, which means a decrease in CH4 concentration, leads to a decrease in carbon deposition rate. • An increasing pressure leads to a decrease in carbon deposition rate. • An increasing reference time, which in case of a fixed bed amounts to the total pyrolysis time, leads to an increase in carbon deposition rate. Table 2: Parameter variation and its impact on carbon deposition rate Nr. Temperature [°C] H2 / C1 Ratio Pressure [bar] Reference time [min] Carbon deposition rate [wt %] 1 1400 4 10 50 4.8 2 1300 4 10 75 2.14 3 1300 3 10 75 3.65 4 1300 2.33 10 61 5.16 5 1200 4 10 90 0.31 6 1300 4 5 75 2.88 7 1300 4 10 75 2.95 8 1300 4 10 75 2.07 9 1300 1 10 11 2.15 10 1300 4 10 75 10 3. Impact of H2 / C1+ ratio and impact of recycling hydrogen-containing product stream on filament formation 3.1 Use of a feed with C1 / C2+ ratio of 12.25 vs. feed containing only C1 (CH4) and no C2+ Pyrolysis of two different hydrocarbon feed stream was carried out in two experiments. The utilized reactor was in both experiments a fixed bed reactor filled over a total height of ~560 mm with solid carbon granule of calcined petro- leum coke (CPC) with a particle size distribution of 2-4mm in a ceramic tube with an inner diameter of 50 mm. The same process conditions were applied in both experiments 1 and 2 during the pyrolysis step, except for only one dif- ference: the composition of the gaseous feed stream: For experiment 1 a mixture of 20 % hydrocarbon feed with a C1 / C2+ ratio of 12.25, in 80% hydrogen was used as feed mixture. For experiment 2 a mixture of 20 % CH4 (only C1, no C2+ in feed), in 80% hydrogen was used as feed mixture. Except the inlet gas composition, the same pyrolysis step was carried out in both experiments 1 and 2. In the pyroly- sis step the feed mixture was dosed in the reactor operated at a pressure of 1 bara. A temperature ramp was applied during the pyrolysis: 1. Starting temperature 600 °C, 2. Increase in temperature from 600 °C up to 1250 °C at a rate of 150 °C / h, 3.1 h operation at 1250 °C, 4. Cooling down to 1000 °C at a rate of 200 °C / h, afterwards the flow of reactants was turned off. The resulting carbon beds of both experiments were analyzed via SEM. In Figure 2 the resulting SEM pictures of ex- periment 1 (Fig.2a) were compared with SEM pictures of experiment 2 (Fig.2b) On the SEM pictures of experiment 2 (no C2+ components in the feed) no filament formation was detected, while numerous filaments were visible in the particles of experiment 1 (C1 / C2+ ratio 12.25). Hence, filament formation was successfully prevented by the usage of a feed gas with a lower C1 / C2+ ratio. The results show that by reducing the C2+ concentration in the reaction feed the formation of filaments can be pre- vented. 2.2 Long-time Experiment Pyrolysis of natural gas was carried out in a moving bed reactor. The reactor had an inner diameter of 190 mm and was 3m heigh. The reactor was filled with carbon granule that comprises a mixture of 70 -80 w% calcinated petro- leum coke and 20 w% - 30 w% of pyrolysis coke produced in the said process. The granule had particle sizes rang- ing from 1 mm to 4 mm. The reactor was operated at 1250 °C. The reactor was direct electrically heated. To com- pensate heat losses, the reactor walls were additionally heated electrically. For experiment 3 the reactor was fed with a mixture of natural gas and a direct recycle of the hydrogen containing product gas. After mixing, the reactor was fed with a gas mixture containing 9 vol% hydrocarbons and 91 vol% hydro- gen. The composition of the hydrocarbons and the volume ratio C1 / C2+ ratio is shown in Table 3. The moving bed was operated over 800 hours without any signs for blocking or disturbance of the moving bed. The continuous opera- tion was observed by the filling level of the reactor. This filling level continuously decreased by discharging the reac- tor at the bottom. After a certain lower level was reached the reactor was refilled from the top. For Experiment 3 this resulted in a regular sawtooth function Table 3: C+ Composition of natural gas and the mixture of natural gas plus internally recycled hydrogen-containing product stream of experiment 3 Composition C+ Natural gas + Natural gas Feed [vol%] direct recycle CH4 95,05 96,94 C2H6 4,04 2,45 C3H8 0,91 0,55 C2H2 0,00 0,06 C2H4 0,00 0,00 C3+ 0,00 0,00 sum 100,00 100,00 C1 / C2+ 19,18 31,72 For experiment 4 the same reactor as used in experiment 3 which was with a mixture of natural gas and a direct re- cycle of the hydrogen containing product gas. After mixing, the reactor was fed with a gas mixture containing 16 vol% hydrocarbons and 84 vol% hydrogen. During this experiment, the reactor was suddenly blocked, and the moving bed was stocked after 9 h operation with this setting. In Figure 3 the blocking is illustrated by the filling level of the reactor. For experiment 4, the filling re- mained suddenly constant due to the blocking. The blocking was due to filament formation by the higher C2+ con- centration as shown on Table 3. Table 3: C+ Composition of natural gas used for experiment 4 Composition C+ Natural gas Feed [vol%] CH4 94,27 C2H6 4,58 C3H8 1,15 C2H2 0,00 C2H4 0,00 C3+ 0,00 sum 100,00 C1 / C2+ 16,44 4. Influence of H2 / C1+ ratio on reaction rate and carbon deposition rate For the examination of the H2 / C1+ ratio´s influence on reaction rate and carbon deposition, the following parameters are used: 1250°C in reaction zone, 1.4 bar(ü), 20 kg / h carbon (gas flow rate adjusted for heat integration). A 2D model implemented in COMSOL is used for the simulations; this model is based on and has been validated with the reactor described in Example 2.2. Four cases are studied at this temperature: (i) 15% CH4 in gas feed, (ii) 35% CH4 in gas feed, (iii) 50% CH4 in gas feed and (iv) 80% CH4 in gas feed, shown in Figure 4.a. Two more cases with 50% CH4 in gas feed, at 1170°C and 1120°C, were simulated, in order to show the influence of temperature on the achievable local carbon deposition, shown in Figure 4.c. The simulation was performed for a moving bed reactor op- erated in counter flow with solid particles moving from top to the bottom. Hence, the highest carbon deposition is at the bottom of the reactor. For the simulation, the pyrolysis was limited to the reaction zone and stopped in the heat integration zone. For the specific carbon deposition rate, it can be seen that the preferred range is reached with 15 vol% and 35 vol% at 1250°C. Higher CH4 concentration exceed the preferred range. By lowering the reaction temperature, the CH4 conversion decreases and the preferred specific deposition is reached with 50 vol% CH4 feed concentrations as well. Another advantage of the H2 dilution is revealed by the normed local reaction rate (Figure 4.b). By decreasing the CH4 concentration, a more homogeneous distribution of the reaction rate along the reactor length is reached, which lowers the probability for blocking. 5. Influence of temperature and H2 dilution on carbon deposition rate In Example 2, the influence of temperature, pressure, H2 dilution, reference time and carrier porosity on carbon dep- osition rate has been shown based on experimental results obtained from a lab reactor operated in fixed bed modus. The interconnections between these parameters and their influence on the carbon deposition rate are complex. In order to give a visual representation of the relation between carbon deposition rate, temperature, and H2 dilution, a parametric variation has been performed using the model described in Example 4. The porosity and reference time are the same for the simulated operating points. The results can be seen in Figure 5, which shows isolines of carbon deposition rate in dependency of temperature and H2 dilution, where CDi < CDii < CDiii. These confirm the trends observed from the experimental results in Example 2: an increasing temperature leads to an increase in carbon dep- osition rate, and a decreasing H2 / C1 ratio, which means an increase in CH4 concentration, leads to an increase in carbon deposition rate.

Claims

Claims:

1. A process of producing hydrogen comprising introducing an external feed stream containing hydrocarbons into a reaction chamber and reacting said hydrocarbons in said reaction chamber in a fixed, fluidized or mov- ing bed of solid materials having a particle size of 0.1 mm to 10 mm at a pressure of 1 to 30 bar, a tempera- ture of 500 to 2000°C, a residence time of the hydrocarbon in the range of 0.1 to 50 s, a conversion rate of 20 to 90% and a carbon deposition rate of 0.1 to 15 wt.-% giving a hydrogen-containing product stream contain- ing unreacted hydrocarbons and solid carbon depositing on said bed material, wherein 10 to 95 vol.-% of said hydrogen-containing product stream is directly recycled internally to said reaction chamber, and wherein the total reactor feed stream containing said external feed stream and said internally recycled feed stream has a volume ratio of H2 / C1+ of 0.5 to 25.

2. Process according to claim 1, wherein said external feed stream containing hydrocarbons has a volume ratio of C1 / C2+ of 3 to 100, said hydrogen-containing product stream containing unreacted hydrocarbons has a volume ratio of C1 / C2+ of 4 to 10000 and said total reactor feed stream containing said external feed stream and said internally recycled feed stream has a volume ratio of C1 / C2+ of 17 to 500.

3. Process according to claim 1 or 2, wherein 40 to 85 vol.-% of said hydrogen-containing product stream is re- cycled internally.

4. Process according to any claim 1 or 3, wherein said carbon deposition rate is 2 to 12 wt.-%, wherein said carbon deposition rate is defined by: ఛ^^^Δ^^ ^^^^^ = 100 ⋅ ^,^^^ ^^^ ^^^= 100 ⋅ ^,ௗ^^^^^௧^^^ ^^ ^,ௗ^^^^^௧^^^ ⋅ ^^^^^^^^ ^^ ^ ^^ ^^^^ = 100 ⋅^௧ୀ^ ^^^ ^^^wherein the reactor specific reference residence time is defined as follows: • Batch / Fixed bed: ^^^^^= total pyrolysis time • Moving bed: ^^^^^= mean residence time of carrier particles in the pyrolysis zone • Fluidized bed: ^^^^^= mean residence time of carrier particles in the pyrolysis zone Reference mass in [g]: • Batch / Fixed bed: ^^^= initial carrier mass in the reactor •Moving bed: ^^ = ^^^^^ ∙ ^^ with the reactor inle ^^^ ^^^^^^^ ^^^ t mass flow ^^^ ^^^^^^^ in [g / s]• Fluidized bed: ^^ = ^ ^^ ^^^ ^^ ^^^^^^^ ∙ ^^^^^ with the reactor inlet mass flow ^^^ ^^^^^^^ in [g / s]wherein^^^ ^,ௗ^^^^^௧^^^ is the carbon deposition rate [g / s]is the carbon mass increase [g]is the reactor specific reference residence time [s]^^^^^^^ is the carbon deposition rate (per reference time unit) in [wt.%].

5. Process according to any of claims 1 to 4, wherein said total reactor feed stream has a volume ratio of H2 / C1+ of 1 to 8.

6. Process according to any of claims 1 to 5, wherein said conversion rate is 30 to 80 %.

7. Process according to any of claims 1 to 6, wherein said external feed stream containing hydrocarbons has a volume ratio of C1 / C2+ of 3 to 15, said hydrogen-containing product stream containing unreacted hydrocar- bons has a volume ratio of C1 / C2+ of 20 to 500 and said total reactor feed stream containing said external feed stream and said internally recycled feed stream has a volume ratio of C1 / C2+ of 20 to 100.

8. Process according to any of claims 1 to 6, wherein said internally recycled feed stream has a volume ratio of C1 / C2+ of 20 to 50.

9. Process according to any of claims 1 to 8, wherein said hydrogen process is conducted at a temperature ranging from 1150 to 1400°C and a pressure ranging from 5 to 15 bar 10. Process according to any of claims 1 to 9, wherein a carbon-containing substrate is used as said solid mate- rial and the carbon content of said solid material is in the range of 90 wt.% to 99.9 wt.%.

11. Process according to any of claims 1 to 10, wherein said hydrogen-containing product stream is fed into said reaction chamber at a temperature of 10 to 200°C.

12. Process according to any of claims 1 to 11, wherein said external feed stream and said recycled hydrogen- containing product stream are mixed before entering the reaction chamber.

13. Process according to any of claims 1 to 12, wherein said hydrogen process is a methane pyrolysis and said methane pyrolysis is conducted in a moving bed reactor and wherein the substrate is guided in countercurrent to the total reactor feed stream and wherein the reactor specific reference residence time ^^^^^is between 1 and 10 h.

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