DE102016125641B4 - Reactor and process for producing a natural gas substitute from hydrogen-containing gas mixtures - Google Patents

Abstract

Reactor for producing a natural gas substitute from higher hydrocarbons, hydrogen and/or carbon monoxide and/or carbon dioxide or any mixtures of these substances or gases containing these substances, characterized in that it is designed to be pressure-resistant and has a solid catalyst support with at least two continuous channels with a diameter of has at least 0.3 mm and a maximum of 3 mm and a surrounding jacket for accommodating the temperature control medium, and that it has at least two solid catalyst supports with different thermal conductivity arranged in series.

Description

The present invention relates to a reactor and a method for producing a natural gas substitute (SNG) from hydrogen-containing gas mixtures with carbon monoxide and/or carbon dioxide, which may contain impurities with higher hydrocarbons in the feed gas stream, and the use of the reactor.

The starting product for the synthesis is, in particular, synthesis gas derived from biomass (biomass gasification and pyrolysis: practical design and theory / Prabir Basu). That is, the origin of the biomass is primarily straw, wood pellets, etc. The processes for carrying out the gasification of these substances are known in principle. The methanation of feed gases containing carbon monoxide and carbon dioxide is a reaction which is chemically catalyzed: $$\begin{array}{l}\text{CO}+3{\text{<figure-callout id="2" label="H" filenames="DE102016125641B4\_0002.png" state="{{state}}">H</figure-callout>}}_2\rightleftharpoons {\text{CH}}_4+{\text{H}}_2\text{O}{\mathrm{\Delta }}_{\text{R}}{H}^0=-206\frac{\text{<figure-callout id="2" label="kJ mol CO" filenames="DE102016125641B4\_0002.png" state="{{state}}">kJ</figure-callout>}}{\text{mol}}& \\ {\text{CO}}_{}{}_2+4{\text{<figure-callout id="2" label="H" filenames="DE102016125641B4\_0002.png" state="{{state}}">H</figure-callout>}}_2\rightleftharpoons {\text{CH}}_4+2{\text{H}}_2\text{O}{\mathrm{\Delta }}_{\text{R}}{H}^0=-165\frac{\text{<figure-callout id="2" label="kJ mol CO" filenames="DE102016125641B4\_0002.png" state="{{state}}">kJ</figure-callout>}}{\text{mol}}& \\ {\text{CO}}_{}{}_2+{\text{<figure-callout id="2" label="H" filenames="DE102016125641B4\_0002.png" state="{{state}}">H</figure-callout>}}_2\rightleftharpoons \text{CO}+{\text{H}}_2\text{O}{\mathrm{\Delta }}_{\text{R}}{H}^0=41\frac{\text{kJ}}{\text{mol}}\end{array}$$

The CO ₂ methanation is therefore an exothermic equilibrium reaction. According to Le Chatelier's principle, low temperatures and high pressures shift the equilibrium towards the products and thus represent the thermodynamically preferable process conditions.

The heat of reaction that is released must be dissipated from the system as effectively as possible. According to the state of the art for the production of natural gas substitutes, adiabatic tray reactors with intermediate cooling or tube bundle reactors filled with catalyst beds or combinations of both systems are used. Since the exothermic methanation reaction heats up the gas mixture during the reaction, a sufficiently high conversion cannot be achieved in a single stage due to the thermodynamic equilibrium limitation. This necessitates several intermediate cooling stages and several reaction trays for sufficiently high conversions. As a result, the method is complex and expensive and limited in terms of its ability to change loads. In the tube bundle reactors filled with the catalyst beds, the heat transport capacity is limited by the poor heat conduction between the catalyst pellets/particles. In order to prevent catalyst damage from unwanted temperature peaks, the diameter of the tubular reactors is limited to 1 to a maximum of 1.5 inches (25 - 40 mm). This in turn leads to a large number (usually several hundred to thousands) of reaction tubes connected in parallel.

In exhaust gas purification in motor vehicles, tubular reactors are used which contain metallic, channel-shaped bodies as catalyst supports. Such reactors are, for example, in the WO-A 2007/020021 , WO-A 2009/315770 , EP-A-0245737 and US-A 5464679 described. These systems are primarily used for exhaust aftertreatment because of their low pressure loss. However, aspects of heat conduction and heat transport are not described.

Furthermore, the use of catalysts in honeycombs is known from the dissertation by M. P. Wolf "Minimization of the pressure gas loss by optimizing the CO removal stage for a stationary PEM fuel cell heating device", Engler/Bunte/Institut Dissertation 2011, KIT.

Due to their poor heat transport properties, the previous tube bundle reactors can only inadequately meet these conditions (transport of the heat of reaction released). Due to the poor heat transport properties, local temperature peaks occur which can lead to the deactivation of the catalyst. Tray reactors, in turn, have no possibility of temperature control and therefore cannot solve the problems according to the invention (transport of the heat of reaction released).

With their catalyst beds, tube bundle reactors have non-uniform heat transport properties that are difficult to predict and define. It is not possible to specifically adjust the temperature along the reaction tube. Attempts are therefore made to run the reaction tubes as isothermally as possible by embedding them in a cooling medium at a constant temperature.

With their temperature range of 200-300° C., the mode of operation of the usual tube bundle reactors is below the temperatures required for starting the activity for reforming the higher hydrocarbons. In tray reactors, these temperatures can occur for a short time, but they cannot be reliably controlled. In addition, these temperatures only occur for such a short time that the residence times are not long enough to convert higher hydrocarbons.

• - BAJOHR, S.; HENRICH, T.: Entwicklung eines Verfahrens zur Methanisierung von biomassestämmigem Synthesegas in Wabenkatalysatoren. In: GWF Gas Erdgas, 2009, Vol. 150, Nr. 1-2 S. 45-51. - ISSN 0016-4909
• - TRONCONI, Enrico [et al.]: Monolithic catalysts with ‚high conductivity‘ honeycomb supports for gas/solid exothermic reactions: characterization of the heat-transfer properties. In: Chemical engineering science, Vol. 59, 2004, No. 22-23, S. 4941-4949. - ISSN 0009-2509
• - GÖRKE, O. [et al.]: Highly selective methanation by the use of a microchannel reactor. In: Catalysis Today, 2005, Vol. 110, S. 132-139. - ISSN 1873-4308

As further prior art, the following can be mentioned:

• - BAJOHR, S.; HENRICH, T.: Development of a process for the methanation of biomass-derived synthesis gas in honeycomb catalysts. In: GWF Gas Erdgas, 2009, Vol. 150, No. 1-2 pp. 45-51. - ISSN 0016-4909
• - TRONCONI, Enrico [et al.]: Monolithic catalysts with 'high conductivity' honeycomb supports for gas/solid exothermic reactions: characterization of the heat-transfer properties. In: Chemical engineering science, Vol. 59, 2004, no. 22-23, pp. 4941-4949. - ISSN 0009-2509
• - GÖRKE, O. [et al.]: Highly selective methanation by the use of a microchannel reactor. In: Catalysis Today, 2005, Vol. 110, pp. 132-139. - ISSN 1873-4308

The object of the present invention is to solve the stated problems of the prior art and to provide a method and a reactor in which a natural gas substitute can be produced in one process stage. In particular, the task is to control exothermic reactions in a defined manner so that the reaction equilibrium is primarily on the side of the product. Here, the control of the heat dissipation for strongly exothermic reactions and thus a conversion as complete as possible should be achieved.

This problem is solved by the subject matter of the independent claims. The dependent claims represent developments of the invention.

The subject matter of the invention is also the use of the solid catalyst support.

The method according to the invention can be used in particular for CO ₂ methanation/CO methanation/MeOH synthesis. In these reactions, the equilibrium shifts towards the reactants at high temperatures. An advantage of the process used according to the invention is that the reaction via the particular solid catalyst support leads to complete CO ₂ conversion/CO conversion. At the entrance to the solid catalyst support, the temperatures are very high, but preferably not higher than 550° C., and the equilibrium is limited. The temperature quickly reaches a high value at the beginning of the introduction of the gas mixture into the respective channel. First, the temperature is at the same level as the surrounding temperature control medium. However, the internal temperature rises abruptly as soon as the first reaction starts.
The temperature drops over the reaction section of the channels of the solid catalyst support. The conversion is very high, ranging from 80% to 100%.

When the synthesis gas flows in, a clear hotspot with temperatures of preferably a maximum of 550° C. forms at the inlet of the solid-state catalyst support. Most of the reactants have preferably already reacted shortly after entering the tubular solid catalyst support. In the rest of the solid catalyst support, the speed of the reaction is so slow, since the concentration of the educt is significantly lower than entering the solid catalyst support, that less heat is released than is removed and the gas cools down. At the outlet of the solid catalyst support, the temperature of the temperature control medium surrounding the solid catalyst support should be almost reached. The hotspot at the entrance to the solid catalyst support ensures that higher-value hydrocarbons are reformed and converted into synthesis gas. The length of the channels of the solid catalyst support ensures that 80% to 100% conversion is achieved according to the invention.

Suitable gas mixtures are mixtures containing carbon monoxide and/or carbon dioxide, which can also contain higher hydrocarbons having two or more than two carbon atoms. Synthesis gases are preferred, ie a mixture of CO/CO ₂ /H ₂ /CH ₄ /H ₂ O. CO ₂ and CH ₄ are used for the CO ₂ methanation. CO/H ₂ /CH ₄ are preferably used for the carbon monoxide methanation. CO/H ₂ /CO ₂ /CH ₄ can be used for the methanol synthesis. In this case, not all components are absolutely necessary for the reactions. On the other hand, other components can certainly be added. These can contain impurities from nitrogen oxides, amines, ammonia and organic compounds with heteroatoms (H, O, P, Cl, F).

For the synthesis of methane, a ratio of H ₂ /CO = 3: 1 and H ₂ / CO ₂ = 4: 1 is aimed at, the CO / CO ₂ : ratio is of secondary importance for the synthesis as long as there is sufficient H ₂ for the respective C/CO ₂ component are available. Because the implementation depends primarily on the H ₂ content. A water vapor addition of 15-30% is advantageous in the methanation of CO. If a methanation of CO ₂ (without the presence of CO) is desired, the addition of steam is not necessary. The temperatures should be between 200-550°C, preferably between 220-550°C, particularly preferably 222-550°C. The temperature of the tempering medium is preferably from 200 to 240.degree. Preferably, nickel-based catalysts can be used. It preferably contains 65-75% by weight of NiO and a remainder of Al ₂ O ₃ as a binder. The pressure is preferably between 5-20 bar.

If the H ₂ /CO (CO ₂ ) ratio does not correspond to the desired ratio, a water-gas shift reaction (CO+H ₂ O ⇌ CO ₂ + _H2 ). However, there is a risk that too much CO ₂ will be produced for the H ₂ that is then present. H ₂ should therefore preferably be added, for example from an electrolysis (with or without a prior shift reaction).

The addition of steam is not necessary for the synthesis of methanol. The ratio of CO:CO ₂ is preferably 5:1. The proportion of H ₂ is preferably 50-70%. The temperature is preferably 200-550°C, more preferably 220-550°C. Cu-ZnO-Al ₂ O ₃ systems are preferably used as catalysts. The pressure is preferably 10-30 bar.

In a temperature range of 200-220°C, the reaction rate is so low that high modified residence times are necessary in order to achieve a high conversion, i.e. the catalyst mass used, based on the reactable reactant gas stream of CO and/or CO ₂ , achieves values of more than 200 kg*s/mol. This can be achieved using large amounts of catalyst >250 g/m ² and/or long channels in the solid catalyst support.

At a temperature below 200°C, gaseous Ni(CO) ₄ is formed. Here, nickel is dissolved out of the coating, so that the effective catalyst mass in the coating is reduced.

The temperature in the hotspot should not exceed 550°C (400°C on the inside of the outer wall of the solid-state catalyst support), otherwise there is a risk that the catalyst will sinter. The reaction temperature is therefore preferably kept below 550°C.

A higher temperature can be achieved if a catalyst or catalyst system optimized for the reformation of higher hydrocarbons is additionally used. Then maximum temperatures of 800-900°C can be reached in the HotSpot.

• - Menge Katalysator in der Beschichtung
• - Strömungsgeschwindigkeit des Synthesegases
• - Durchsatzmenge des Synthesegases
• - Veränderung der Wärmeabtransportfähigkeit des Festkörperkatalysatorträgers z.B. durch Änderung der Struktur des Festkörperkatalysatorträgers, Änderung des für den Festkörperkatalysatorträger eingesetzten Materials, Einstellung des Leerraumanteils, Änderung des Verhältnisses m² Wärmetauscherfläche/m³ Festkörperkatalysatorträgerinnenwände, Änderung der Struktur der Form der Kanäle, der Dicke der Innenwände und der Durchmesser der Kanäle.

With regard to the design of the solid catalyst carrier, a "design point" must be determined taking the following points into account:

• - Amount of catalyst in the coating
• - Syngas flow rate
• - Synthesis gas flow rate
• - Changing the ability of the solid catalyst carrier to transport heat, e.g. by changing the structure of the solid catalyst carrier, changing the material used for the solid catalyst carrier, adjusting the proportion of empty space, changing the ratio m ² heat exchanger area/m ³ solid catalyst carrier inner walls, changing the structure of the shape of the channels, the thickness of the inner walls and the diameter of the canals.

The design point is determined by a numerical simulation of the solid catalyst carrier. Simulation software, preferably Comsol Multiphysics, is used for this. In this software, the geometry is post 3 modeled to the outside of 13. Heat dissipation is taken into account with a heat transfer coefficient on the outside of 13. A calculation grid with a sufficient number of cells is generated. Sufficient here means that the calculation result does not change significantly if the number of cells is further increased. The relevant energy and mass transport equations are implemented for each part of the geometry. In this case, the solid catalyst support is preferably mapped homogeneously in order to minimize the computational effort. Correlations for the effective thermal conductivity known from the literature are used to describe the heat transport in the solid catalyst support to implement in solid catalyst supports. The reaction rate is calculated using formal kinetics known from the literature. A mass transport limitation is preferably implemented via a catalyst efficiency in order to take into account the mass transport resistance of the catalyst layer. The calculation is preferably carried out for the stationary case in order to reduce the computational effort. The parameters relevant to the reactor design must be specified for the calculation. These are the heat transfer coefficient on the outside of 13, the length of the reactor, the amount of catalyst per internal volume of the solid catalyst support, the gas velocity, the temperature of the temperature control medium, the composition of the educt gas, the diameter of the solid catalyst support, the thermal conductivity of the various materials in question and the thickness of the walls. Possible combinations of these values result from the features according to the invention.

According to the invention, the solid catalyst supports can be arranged next to one another and/or one behind the other.

The corresponding specifications are made before the solid catalyst carrier is constructed. During the ongoing process, the speed of the synthesis gas and the temperature of the temperature control medium, as well as the pressure in the solid catalyst support, can be controlled and/or regulated via control or regulating valves and the delivery of the product stream. The load change rate, i.e. the change in the amount of synthesis gas supplied, should not exceed a maximum of 5%/min.

The advantages of the process in the solid catalyst support lie in the high load range. The system enables 100% conversion of the synthesis gas to methane/-ol regardless of the utilization if this fluctuates in a range of 20-100% of the maximum possible utilization. That is, the system makes it possible to ensure 100% conversion of the synthesis gas to methane/methanol, even if the amount of synthesis gas supplied changes, eg from 200m ³ /h to 1000m ³ /h. However, the system is designed for 100% utilization. If the system is run with the parameters of the design point and 100% of the synthesis gas is fed in, there is 100% conversion of the CO/CO ₂ . If the synthesis gas supply is reduced, the system continues to run with 100% conversion, provided at least 20% of the synthesis gas flows in and the load change rate of max. 5%/min is maintained.

According to the invention, inferior synthesis gases (with impurities from higher hydrocarbons = long-chain, mono- and polyaromatic) can also be used for the reaction, since they are reformed in the solid catalyst support under the conditions prevailing in the hot spot and then used as synthesis gas for the methane/ -ol synthesis are available. For example, hydrocarbons are broken down into CO/CO ₂ and H ₂ under steam. During methanation or methanol synthesis, water vapor is produced or added. There is therefore a three-fold effect: low-grade synthesis gases can be used, the higher hydrocarbons do not have to be separated off before the reaction, and the product stream has to be cleaned less or not at all. Overall, the efficiency is thus increased due to better utilization of the exothermic reaction and greater flexibility with regard to the synthesis gas (impurity, quantity).

The solid catalyst support according to the invention is preferably designed as a pressure-resistant reactor which has a large number of channels. The channels have continuous openings. Channels without corners are preferred. Because a coating of the round or curved ducts is advantageous in contrast to ducts with corners and sharp edges, since cracks occurring at the edges or incomplete coatings can be avoided. The channels should be small enough in diameter to ensure good exchange between the catalyst coating and synthesis gas and large enough for the gas to flow through the channels.

• - Die innere Oberfläche der Kanäle im Festkörperkatalysatorträger werden aufgeraut, indem z.B. die Wände der Kanäle mit Säure vorbehandelt werden (eintauchen in Säure für mindestens 3 min). Dadurch wird die Oberfläche angeätzt oder oxidiert.
• - Der Festkörperkatalysatorträger wird in eine Beschichtungssuspension eingetaucht, entnommen, getrocknet und kalziniert.
• - Eine Beschichtungssuspension aus Katalysator + Bindemittel (V_Bindemittel=1.545 m_kat) + dest. Wasser (Vwasser=2,5 VBindemittel) wird hergestellt. Z.B., wird eine Suspension aus dem Katalysator (für die Methansynthese: NiO; für die MethanolSynthese: CuO 40 Gew.-%, ZnO 20 Gew.-% der Gesamtkatalysatormasse) mit einem Oxidationsmittel (anorganische/organische Säuren, wie z. B. HCl, CH₃COOH) und kolloidales, disperses Aluminiumoxid als Bindemittel bei einer Temperatur von 20-40°C, vorzugsweise 30-35°C, hergestellt, die eine Viskosität von 50-200 mPas, vorzugsweise 120-140 mPas aufweist. Die durchschnittliche Korngröße x des Katalysators beträgt 10 µm≤x≤100 µm (Bestimmung mittels Laserbeugung).
• - Die Katalysatormasse liegt vorzugsweise bei 50-120 g/m² Oberfläche der inneren Oberflächen der Kanäle
• - Der pH-Wert wird auf 3-4, bevorzugt 3,5 - 3,8, eingestellt
• - Es folgt das Eintauchen des Kanals in die Suspension, ablaufen lassen der überflüssigen Suspension und Waben mit Druckluft ausblasen.
• - Danach wird bei mehr als 100°C getrocknet
• - und bei 450°C für mindestens 3 h kalziniert.
• - Die Menge an aufgetragenen Katalysator wird mittels Differenzwägung bestimmt

The inner surfaces of the channels are coated with at least one catalyst. The gas mixture flows through the channels and comes into contact with the catalyst. Any compounds can be considered for the coating on the inner surface of the channels. This means that both porous and non-porous catalysts can be used. However, porous catalysts are particularly preferred. Metal compounds are primarily suitable for this purpose. Catalysts that can be used can contain Fe, Co, Ni, Cu, Ru, Rh, Pd, Ag, Re, Os, Ir, Pt, Ti, Au, or V. The metals used are hydrogenation-active and can also be present in the catalyst in small proportions. The catalysts mentioned can be used individually or in any combination with one another. These can be present, for example, in a mixture with colloidal, disperse aluminum oxide. For example, in the case of the conversion tion of methane and methanol, NiO or Co/Zn can be used as the catalyst. The application of the coating on the inner surface of the channels can be carried out as follows:

• - The inner surface of the channels in the solid catalyst support are roughened, for example by pretreating the walls of the channels with acid (immersion in acid for at least 3 minutes). This will etch or oxidize the surface.
• - The solid catalyst support is immersed in a coating suspension, removed, dried and calcined.
• - A coating suspension of catalyst + binder (V _binder = 1,545 m _cat ) + dist. Water (Vwater=2.5 Vbinder) is produced. For example, a suspension of the catalyst (for methane synthesis: NiO; for methanol synthesis: CuO 40% by weight, ZnO 20% by weight of the total catalyst mass) with an oxidizing agent (inorganic/organic acids such as HCl , CH ₃ COOH) and colloidal, disperse aluminum oxide as a binder at a temperature of 20-40° C., preferably 30-35° C., which has a viscosity of 50-200 mPas, preferably 120-140 mPas. The average particle size x of the catalyst is 10 μm≦x≦100 μm (determined by means of laser diffraction).
• - The catalyst mass is preferably 50-120 g/m ² surface area of the inner surfaces of the channels
• - The pH is adjusted to 3-4, preferably 3.5-3.8
• - The channel is then immersed in the suspension, the excess suspension is drained off and the honeycomb is blown out with compressed air.
• - Then it is dried at more than 100°C
• - and calcined at 450°C for at least 3 hours.
• - The amount of applied catalyst is determined by differential weighing

The individual cavities of the channels can vary in the geometry of the inner walls by selecting different channel diameters if at least two solid catalyst supports are connected in series. The wall thicknesses of the solid catalyst supports are freely selected within the technical strength limits and designed for the thermal conductivity adapted to the reaction to be carried out.

The wall thickness of the channels of the solid catalyst support preferably has a thickness of about 1-2 mm. Such solid catalyst carriers are pressed into an additional pressure-resistant jacket tube (wall thickness 2-6 mm) in order to generate the required pressure resistance. According to the invention but this is not necessary if z. B. two solid catalyst supports are welded together. The only requirement is that the outer wall of the solid catalyst carrier or the outer wall of the solid catalyst carriers arranged one behind the other in a row is sufficiently pressure-resistant. The outer wall (with or without an additional pressure-resistant metal tube) derives the temperature in the direction of the surrounding temperature control medium and along the tubular solid catalyst carrier. The inner wall thickness of the channels can be 10 μm to 10 mm, preferably 10 μm to 2 mm.

The length of the passage opening of the channels is at most 3 mm, preferably 0.3-2 mm, particularly preferably 0.7-1 mm. The diameter or the opening is preferably at least 0.3 mm wide in each direction, since an approx. 150 μm thick coating is applied to the inner walls and the synthesis gas must still be able to pass in between.

According to the invention, solid catalyst supports with channel structures of any shape can be used. You can be configured differently in their cross section. These cross sections can preferably be circular or elliptical. Polygons, e.g. hexagonal shapes, are also conceivable.

These are preferably tubular configurations with an outer jacket tube and the solid catalyst support arranged therein with two, preferably a large number of channels with structured inner surfaces. According to the invention, at least two channels are required. According to the invention, the use of more than 200 channels is preferred. The inner surfaces can be made of metal. Metals or alloys that have good thermal conductivity, e.g. steels and/or aluminum, are preferably suitable for this purpose. Graphites and/or carbides, which can also be used in combination with the aforementioned metals, are also conceivable.

The multiplicity of preferably more than 200 channels takes up the gas flow. The catalysts are arranged on the inner surfaces of these individual channels. According to the invention, the channels are as fine as possible with a diameter of at least 0.3 mm and at most 3 mm.

The more channels there are in fine detail, the larger the surface area. This in turn provides good heat transfer and contact with the catalyst coated on the inner surfaces of the channels. The heat transport can therefore be influenced by the channel diameter and the number of channels.

The reaction is tempered by a tempering agent in the jacket flow, which envelops the solid catalyst support. This flow can be carried out in co-current or counter-current. Countercurrent is preferred. A high radial dissipation of the heat is also preferred, above all at the beginning of the reaction at the channel entrance. Since the highest temperature gradients occur there. A minimum temperature at the outlet is specified by the tempering medium.

According to the invention, the temperature can be controlled by setting a suitable flow direction of the temperature control medium. In the case of direct current, there must be a maximum difference in temperature for heat to be transported away at the inlet. In the countercurrent process, the temperature difference must be high enough so that the heat of the reaction can be dissipated right from the start when the gas enters.

Various solid catalyst supports can be arranged in series. This can involve catalyst supports made of different materials. These therefore have different thermal conductivity or different internal configurations. I.e. the channel diameter, the inner wall thickness, the channel geometry and the catalyst coating have an influence on the thermal conductivity.

According to the invention, the solid catalyst support can also be provided with an intermediate volume. I.e. behind a solid catalyst support there can be an empty space and then a subsequent solid catalyst support. This reduces the axial heat dissipation via the subsequent solid catalyst carrier. In this case, with regard to the intermediate volume, a pressure-resistant jacket tube must be arranged so that the temperature control medium can be guided in the jacket flow over the entire length of the solid catalyst support system.

Accordingly, the temperature control medium in the jacket flow is preferably guided axially in opposite directions. The temperature of the tempering medium is preferably 200-240° C., for example in the synthesis of methane and methanol. The temperature throughout the solid catalyst support is at least 200°C. In this case, a maximum heat dissipation at the channel inlet of the catalyst support is adjusted axially in opposite directions. This is where the so-called hotspot temperature occurs and a large amount of C/CO ₂ + H ₂ reacts in the presence of the catalyst.

The material is therefore selected according to the heat dissipation properties of the solid catalyst support. The thermal conductivities of the materials that can be used, for example, according to the invention are at the following values: steel approx. 15 W/mK, aluminum approx. 230 W/mK, graphite approx. 120 - 160 W/mK.

The first selection criterion is the thermal conductivity of the material. If this is suitable, the second criterion checked is whether it can be used in the temperature range according to the invention. According to another criterion, coatability with the catalyst must be guaranteed, e.g. by preparing the surface accordingly. Finally, shape, mechanical strength and coefficient of thermal expansion must be retained when the coating is applied. Metals, graphite, carbides, steels, alloys which are not hydrogenation-active/shift-active are therefore preferably considered as the material of the solid-state catalyst support or its pressure-resistant additional jacket tube. If steel is used, it may contain a component which is hydrogenation or shift-active on its own, but which in the form of the steel has no or only negligible hydrogenation or shift activity. The material of the solid catalyst support and its outer jacket tube do not have to be the same.

The design of the solid catalyst support can be selected on the basis of empirical values and/or simulations. Since, for example, a high temperature is to be expected in the entry area of the solid-state catalyst support, a material with high thermal conductivity can be considered for the front part. This can prevent the temperature in the HotSpot from rising above 550°C. The more The solid catalyst support can dissipate heat, the more educt gas can be converted in the solid catalyst support volume/time without exceeding the temperature of 550°C. In other words, the required volume becomes smaller the greater the thermal conductivity of the solid catalyst support.

The solid catalyst support has a specific m ² heat exchanger surface area/m ³ solid catalyst support internal volume ratio, depending on how the reaction is to be designed. The heat exchanger surface is the outer wall that is in contact with the surrounding temperature control medium. The better the thermal conductivity of the selected material, the smaller the value for the ratio m ² heat exchanger surface area/m ³ internal volume of solid catalyst support. The value is preferably 30-130 m ² /m ³ .

Fluids of all kinds can be considered as temperature control media. Liquids are particularly preferred. That is to say, compounds known to those skilled in the art can be used as temperature control agents for heat dissipation or supply. These include, for example, liquids such as dibenzyl toluene and silicone oils. Boiling water circuits or molten salts are also conceivable for cooling. The heat transfer coefficient is preferably 100-300 W/m ² K max. 500 W/m ² K. The heat transfer coefficient can be achieved by the speed of the flow against or around the pipe.

Honeycomb configurations are particularly preferred for the solid catalyst support with a multiplicity of channels. Under the honeycomb will be in the EP0245737 A1 and also in DE 4303950 C1 systems described. Such honeycomb structures are also described in MP Wolf's dissertation cited above. According to the invention, therefore, preferably come in the DE 4303950 C1 mentioned systems into consideration. Reference is made here in particular to column 1, line 49 to column 4, line 38. That is to say that metallic solid catalyst carriers are preferably considered, in particular catalyst carriers such as can also be used in exhaust gas aftertreatment systems of motor vehicles with internal combustion engines. These are bodies that, together with the walls, form a large number of channels through which a fluid can flow.

The solid catalyst supports are preferably structures which, in an alternative, are designed in such a way that their smooth and straight transverse surfaces can be fitted together with those of other channels of the solid catalyst support in order to create a connection. However, this is not absolutely necessary.

The outer surfaces of the channels can be connected using all methods familiar to a person skilled in the art. I.e. clamping methods, gluing methods or welding methods can be considered. Because the channels are preferably metal or contain metal, welding is a preferred form of connection. However, it is not absolutely necessary that the walls of the channels fit together exactly. It is essential that the channels can be flowed through and that no gas can escape laterally at a pressure of up to 30 bar. This is particularly important for the weld seams. The channels themselves are exposed to almost no differential pressure. The channel density is preferably 100-600 cpsi (channels per square inch), preferably 200-400 cpsi.

The void fraction of the solid catalyst support preferably corresponds to the fraction that is filled with the synthesis gas and is preferably 40-85%, particularly preferably 65%. According to the invention, the thermal conductivity of the solid catalyst carrier can also be adjusted by varying the empty spaces, the inner and outer wall thickness of the channels, the material of the solid catalyst carrier and/or its channel densities. The absorption and/or dissipation of the heat can be adjusted by means of the material or the filling of the inner walls of the channels of the solid catalyst support

According to the invention, the solid catalyst support is either designed to be pressure-resistant, or it is arranged in a pressure-resistant jacket tube. The syntheses take place under pressure at preferably 5-30 bar. The solid catalyst support has to withstand this. If this is arranged in a pressure-resistant casing pipe, there is preferably a planar connection from the inside to the pressure-resistant casing pipe. If such a pressure-resistant jacket pipe is present, there is a planar connection to the temperature control medium flow. This achieves a maximum temperature transfer from the solid catalyst support to the optionally arranged pressure-resistant jacket tube and to the temperature control medium.

According to the invention, the pressure-resistant jacket tube can also be made of a material with a lower or the same thermal thermal expansion compared to the material of the solid catalyst support, or be made of the same material as the solid catalyst support.

The solid catalysts arranged in series can consist of one material or different solid catalysts can consist of either different materials or the same material but with a different structure. The structure of the channels with regard to the proportion of voids, channel density and channel shape depends on the desired heat dissipation performance.

In an alternative, the channels of solid catalyst supports arranged one behind the other in a row have the same diameter.

The proportion of empty space, the shape of the channels, the channel density, the channel diameter do not have to be the same in the case of two solid catalyst supports arranged one behind the other in a row. However, the prerequisite is that the channels can be freely flowed through. A number of solid catalyst supports can be arranged next to one another in an overall system. This increases the surface area and favors the amount of synthesis gas that can be converted.

In the method according to the invention, a precise temperature profile can be generated in the channels of the solid catalyst support by arranging different solid catalyst supports with different heat conduction properties, despite the constant temperature of the surrounding temperature control medium. A wide load range can be achieved by arranging solid catalyst carriers made of different materials or with different thermal conductivities in a row.

Due to the possibility of tailoring the temperature along the solid catalyst support, the space-time yield can be maximized, which leads to a reduction in the reactor volume and the amount of catalyst. Furthermore, the diameter of the channels of the solid catalyst carrier can be increased due to the improved heat transport, which leads to a reduction in the number of parallel solid catalyst carriers. Both effects (reduction of the reaction volume, fewer parallel solid catalyst carriers) reduce the apparatus costs and thus increase the economics of the process. Due to the tolerance of the process to the higher hydrocarbon contamination of the syngas, which is achieved by setting a reforming zone in the initial region of the channel of the solid catalyst support with a temperature range of -350 to ~550°C, lower quality feed gases can be used. This reduces upstream gas cleaning process steps.

Temperature distributions in the reaction system can be specifically preset with the described solid catalyst supports arranged one behind the other. This enables the reaction system to operate at the level of the optimum reaction rate. Surprisingly, a gas mixture which contains higher hydrocarbons as impurities can also be used in this way according to the invention. In this way, additional synthesis gas can also be produced from these substances, which is hydrogenated in-situ.

• In der 1 ist ein Rohrreaktor nach dem Stand der Technik dargestellt. Über den Katalysator 6 wird in Strömungsrichtung 5 das Gas abgeführt. Der Reaktor weist die Höhe 8 und den Durchmesser 7 auf.
• In 2 ist ein Festkörperkatalysatorträgersystem nach der Erfindung dargestellt. Hier sind die einzelnen Festkörperkatalysatorträger 1 unterschiedlicher Geometrie dargestellt. Sie sind gemäß Ziffer 2 miteinander verbunden. Die Festkörperkatalysatorträger haben den Durchmesser 3 und in Strömungsrichtung 4 verlässt das Gas den Reaktor.
• In 3 ist ein weiterer Reaktor nach der Erfindung dargestellt. Hierbei ist gemäß Ziffer 9 der Einlauf vorgesehen, während in Ziffer 11 der Auslauf vorhanden ist. In Ziffer 10 ist der erfindungsgemäße Festkörperkatalysatorträger vorhanden. Die Pfeile deuten die Bewegungsrichtung des Synthesegases an. Durch Ziffer 15 wird der Radius des Festkörperkatalysatorträgers, welcher durch die Ziffern 13 und 14 begrenzt ist, dargestellt. Die Kanalwand setzt sich aus der inneren Wand 13 und äußeren Wand 14 mit einem dazwischenliegenden Leerraum zusammen. Zwischen der inneren Wand 13 und der äußeren Wand 14 wird das Temperiermittel 12 geführt.
• 4 zeigt eine Gesamtanlage in schematischer Darstellung. Hierbei ist in Ziffer 17 ein Shift-Reaktor vorgesehen, der eine Wasser-Gas-Shift Reaktion gemäß Stand der Technik durchführen kann, vorgesehen. Gemäß Ziffer 18 ist der erfindungsgemäße Festkörperkatalysatorträger im Reaktor vorgesehen. Ziffer 19 bezeichnet den Thermostaten. In Ziffer 20 ist ein Separator angeordnet. Das Abwasser läuft über das Element gemäß Ziffer 21 ab. Die Verdampfung erfolgt über Ziffer 22. Das Produktgas 23 wird von dem Separator 20 abgeführt. D.h. in Ziffer 20 erfolgt eine Auftrennung in Abwasser 21 und Produktgas 23. Im Einzelnen wird demgemäß das Synthesegas eingespeist. Unter Zugabe von Wasserstoff und einer Shiftreaktion im Shift-Reaktor 17 wird das Verhältnis CO/CO₂:H₂-Verhältnis angepasst. Der Shift-Reaktor 17 wird mit Wasserdampf für die Shiftreaktion angereichert. Das Synthesegas wird mit dem Wasserstoff durch die Kanäle des Festkörperkatalysatorträgers 18 geleitet, das von dem Temperiermittel auf der voreingestellten Temperatur mittels des Thermostaten 19 gehalten wird. Im Separator 20 ist das Abscheiden von Wasser 21 und des Produktes 23 vorgesehen. D.h. dort kann das Produkt 23 entnommen werden.
• In 5 ist in der Seitenansicht ein Reaktor mit Festkörperkatalysatorträger dargestellt. Gemäß Ziffer 9 ist wieder der Einlauf erkennbar. In Ziffer 10 ist der Festkörperkatalysatorträger dargestellt. Ziffer 11 bezeichnet den Auslauf. Gemäß Ziffer 21 verlässt das Produktgas das Reaktorsystem. Gemäß Ziffer 13 ist die innere Wand des Formkörperkatalysatorträgers dargestellt. Die äußere Begrenzung 14, d. h. die äußere Wand 14 umschließt den Lauf des Temperiermittels. In 5 ist dies im Gegenstrom dargestellt.
• Gemäß 6 ist ein Reaktor mit einem Festkörperkatalysatorsystem, bestehend aus zwei Festkörperkatalysatorträgern dargestellt. D.h. es sind die beiden Festkörperkatalysatorträger 10 a und 10 b vorhanden. Ansonsten ist der Gang des Verfahrens wie in 5.
• 7: Hier ist der Festkörperkatalysatorträger 10 a und 10 b durch einen Zwischenraum getrennt. Dieser Zwischenraum bedingt eine verbesserte axiale Abfuhr der Wärme.

The invention is described in more detail below with reference to the figures:

• In the 1 a tubular reactor according to the prior art is shown. The gas is discharged in the direction of flow 5 via the catalyst 6 . The reactor has the height 8 and the diameter 7.
• In 2 a solid catalyst support system according to the invention is shown. The individual solid catalyst supports 1 of different geometry are shown here. They are linked to one another in accordance with clause 2. The solid catalyst supports have the diameter 3 and the gas leaves the reactor in the direction of flow 4 .
• In 3 a further reactor according to the invention is shown. In this case, the inlet is provided according to number 9, while the outlet is available in number 11. In number 10 the solid catalyst support according to the invention is present. The arrows indicate the direction of movement of the synthesis gas. Numeral 15 represents the radius of the solid catalyst support, which is delimited by numerals 13 and 14. The channel wall is composed of inner wall 13 and outer wall 14 with a void space therebetween. The temperature control medium 12 is guided between the inner wall 13 and the outer wall 14 .
• 4 shows an overall system in a schematic representation. In this case, a shift reactor is provided in number 17, which can carry out a water-gas shift reaction according to the prior art. According to item 18, the solid catalyst support according to the invention is provided in the reactor. Numeral 19 designates the thermostat. In number 20 a separator is arranged. The waste water runs off via the element according to paragraph 21. Evaporation takes place via point 22. The product gas 23 is of the separator 20 discharged. Ie in section 20 there is a separation into waste water 21 and product gas 23. In detail, accordingly, the synthesis gas is fed. The CO/CO ₂ :H ₂ ratio is adjusted by adding hydrogen and a shift reaction in the shift reactor 17 . The shift reactor 17 is enriched with steam for the shift reaction. The synthesis gas is conducted with the hydrogen through the channels of the solid catalyst support 18, which is kept at the preset temperature by the thermostat 19 by the temperature control means. The separation of water 21 and the product 23 is provided in the separator 20 . Ie the product 23 can be removed there.
• In 5 a reactor with a solid catalyst support is shown in the side view. According to number 9, the inlet can be seen again. Numeral 10 shows the solid catalyst support. Numeral 11 designates the spout. According to paragraph 21, the product gas leaves the reactor system. According to number 13, the inner wall of the shaped body catalyst support is shown. The outer boundary 14, ie the outer wall 14 encloses the course of the temperature control medium. In 5 this is shown in countercurrent.
• According to 6 is a reactor with a solid catalyst system, consisting of two solid catalyst carriers shown. That is, the two solid catalyst supports 10 a and 10 b are present. Otherwise the course of the procedure is as in 5 .
• 7 : Here the solid catalyst carrier 10 a and 10 b is separated by a gap. This gap results in improved axial heat dissipation.

exemplary embodiments

Example 1:

Synthesis of Methane

Parameter set for the simulation in Comsol Multiphysics 5.2a simulation for experiment 1 2 3 4 parameter value value value value Unit heat transfer coefficient 236 124 122 112 W/m ² K Diameter of the channels of the solid catalyst support (10) 33 50 76.2 105 mm Diameter solid catalyst support + pressure pipe (10+13) 40 60.3 88.9 114.3 mm catalyst loading 53 gsm ² comb length 101.5 mm cell density 200 cpsi void fraction 0.65 - thickness of the inner walls 0.11 mm

A coated solid catalyst (length: 101.5 mm) is placed in a reactor according to 3 built-in.

• Temperiermittel: Dibenzyltoluol
• Temperatur: 240°C
• Druck: 8 bar
• Geschwindigkeit Eingangsstrom/Synthesegas bei Normbedingungen (1,013 bar, 0°C): 0,05 m/s
• Synthesegaszusammensetzung: H₂:CO₂:N₂ = 80:20:5
• Innenwände Material: Cr Al 20 5; Außenrohr: X2CrTiNb18
• Leerraumanteil: 65%
• Kanaldichte: 200 cpsi
• Wandstärke Innenwände 110 µm
• Druckrohr Material: Stahl X6CrNiMoTi17-12-
• Laufzeit: >1000 h

Reaction parameters:

• Tempering agent: Dibenzyltoluene
• Temperature: 240°C
• Pressure: 8 bars
• Velocity of input stream/synthesis gas at standard conditions (1.013 bar, 0°C): 0.05 m/s
• Synthesis gas composition: H ₂ :CO ₂ :N ₂ = 80:20:5
• Inner walls Material: Cr Al 20 5; Outer tube: X2CrTiNb18
• Void percentage: 65%
• Channel Density: 200cpsi
• Wall thickness inner walls 110 µm
• Pressure pipe Material: Steel X6CrNiMoTi17-12-
• Running time: >1000 hours

The system is first tested for pressure tightness at a pressure of 8 bar. The reactor is then insulated and the outer reactor tube is filled with the temperature control medium. Then the catalyst has to be reduced. For this purpose, the reactor is operated at an oil temperature of 300° C. and a pressure of 8 bar with an H ₂ :N ₂ molar ratio of 10:3.

The conversions are determined from the volumetric compositions of the product gas mixture measured in the µ-gas chromatograph (GC) in comparison to the educt gas composition. The µ-gas chromatograph used separates the product gas into its components using two columns, the quantitative analysis is carried out by comparing the thermal conductivity with that of the reference gas (column 1: molecular sieve column with argon as carrier gas for separating H ₂ , N ₂ , O ₂ , CO and CH ₄ ; Column 2: Plot-U column for separating CO ₂ and C ₂ and C ₃ molecules with helium as carrier gas). experiment 1 2 3 4 Diameter channels in mm 33 50 76.2 105 Sales XCO ₂ 83.5% 80.4 88.9% 90.5% Produced Methane 0.3kW 0.6kW 1.6kW 3.0kW Applied coating in g/m ² 9.6 g total amount of coating → 53 g/m ² (NiO) 21.1 total amount of coating → 52.8 g/m ² (NiO) 51.4 total amount of coating → 53.2 g/m ² (NiO) 101.4 total amount of coating → 53.1 g/m ² (NiO) Wall thickness pressure pipe in mm 2.0 3.6 4.4 4.4

Example 2:

• 6 Festkörperkatalysatorträgersysteme mit jeweils 20 Festkörperkatalysatorträgern (aneinandergeschweißt)
• Kanäle: Gesamtlänge 2 m; Durchmesser außen 42,4 mm; Durchmesser Matrix 37,2 mm -> Wandstärke Druckrohr 2,6 mm
• Temperiermittel: Dibenzyltoluol
• Temperatur: 240°C
• Druck: 5-20 bar
• Synthesegaszusammensetzung:
  • Synthesegas:
    • H₂ 52,8 Vol.-%
    • CO 29,2 Vol.-%
    • CO₂ 16,2 Vol.-%
    • CH₄ 1 Vol.-%
    • N₂ 0,8 Vol.-%
  • H₂-Zugabe bis Verhältnis für CO und CO₂-Methanisierung erreicht wurde
• Innenwände Material: Cr Al 20 5
• Außenrohr: Stahl X6CrNiMoTi17-12-
• Außenrohr = Druckrohr
• Leerraumanteil: 65%
• Kanaldichte: 200 cpsi
• Wandstärken: Außenwand 2,6 mm
• Laufzeit: >1000 h

Reaction parameters:

• 6 solid catalyst carrier systems, each with 20 solid catalyst carriers (welded together)
• Channels: total length 2 m; outer diameter 42.4 mm; Diameter matrix 37.2 mm -> wall thickness pressure tube 2.6 mm
• Tempering agent: Dibenzyltoluene
• Temperature: 240°C
• Pressure: 5-20 bars
• Synthesis gas composition:
  • synthesis gas:
    • H ₂ 52.8% by volume
    • CO 29.2% by volume
    • CO ₂ 16.2% by volume
    • CH ₄ 1% by volume
    • N ₂ 0.8% by volume
  • H ₂ addition until ratio for CO and CO ₂ methanation was reached
• Inner walls material: Cr Al 20 5
• Outer tube: steel X6CrNiMoTi17-12-
• Outer pipe = pressure pipe
• Void percentage: 65%
• Channel Density: 200cpsi
• Wall thickness: outer wall 2.6 mm
• Running time: >1000 hours

The solid catalyst supports are loaded with 53 g/m ² NiO. The system is first tested for pressure tightness at a pressure of 20 bar. Then the catalyst has to be reduced. For this purpose, the reactor is operated at an oil temperature of 300° C. and a pressure of 20 bar with an H ₂ :N ₂ molar ratio of 10:3.
After completion of the reduction, the synthesis gas is compressed in a compression stage to a pressure of 20 bar, and H ₂ is added to the compressed synthesis gas. Before entering the reactor, the gas is preheated to a temperature of 220°C. The required amount of H ₂ is constantly calculated via a routine of the process control system for the respective gas composition and amount of gas that is pumped into the system.

The conversions are determined from the volumetric compositions of the product gas mixture measured in the µ-gas chromatograph (GC) in comparison to the educt gas composition. The µ-gas chromatograph used separates the product gas into its components using two columns, the quantitative analysis is carried out by comparing the thermal conductivity with that of the reference gas (column 1: molecular sieve column with argon as carrier gas for separating H ₂ , N ₂ , O ₂ , CO and CH ₄ ; Column 2: Plot-U column for separating CO ₂ and C ₂ and C ₃ molecules with helium as carrier gas). Syngas velocity [m/s] 0.5 1 2 pressure 12 12 12 Conversion xCO ₂ [%] 98.9 97.5 98.2 Methane produced [kW] 71 71 95 Number of tubes in parallel 6 3 2

Claims (12)

Reactor for producing a natural gas substitute from higher hydrocarbons, hydrogen and/or carbon monoxide and/or carbon dioxide or any mixtures of these substances or gases containing these substances, characterized in that it is designed to be pressure-resistant and has a solid catalyst support with at least two continuous channels with a diameter of at least 0.3 mm and a maximum of 3 mm and a surrounding jacket for accommodating the temperature control medium, and that it has at least two solid catalyst supports with different thermal conductivity arranged in series. reactor after claim 1 characterized in that the solid catalyst support has at least 200 channels. reactor after claim 1 or 2 characterized in that it has at least two solid catalyst carriers arranged next to one another, so that they form a solid catalyst carrier matrix. Reactor according to one of the preceding claims , characterized in that the density of the channels is between 200 and 600 cpsi. Reactor according to one of the preceding claims , characterized in that the ratio of heat exchanger surface area to channel internal volume is 130 - 30 m ² /m ³ . Reactor according to one of the preceding claims , characterized in that it is surrounded by a jacket to accommodate temperature control medium. Reactor according to one of the preceding claims , characterized in that it contains materials which are not hydrogenation active. Reactor according to one of the preceding claims , characterized in that either the solid catalyst support is designed to be pressure-resistant or it is surrounded by a pressure-resistant jacket tube. Process for the production of a natural gas substitute, wherein at least one gas mixture is introduced into the solid catalyst carrier, wherein the temperature of the gas mixture is between 550 and 220 ° C, over the length by means of a temperature control agent surrounding the solid catalyst carrier a drop in the temperature profile and a thermal conductivity through the selection of the material and geometry of the channels is defined, characterized in that a reactor according to one of Claims 1 until 8th is used. procedure after claim 9 characterized in that higher hydrocarbons, hydrogen and/or carbon monoxide and/or carbon dioxide-containing mixtures or any desired mixtures of these substances are used as gas mixtures. Method according to one of the preceding claims , characterized in that the gas mixture of carbon monoxide and/or carbon dioxide is converted into methane or methanol. Use of a reactor according to any one of Claims 1 - 8th characterized in that it is used for the production of methane and methanol.

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