WO2019233674A1

WO2019233674A1 - Method and reactor system for carrying out catalytic gas phase reactions

Info

Publication number: WO2019233674A1
Application number: PCT/EP2019/060461
Authority: WO
Inventors: Christian SCHUHBAUER; Dieter Verbeek; Rolf Bank
Original assignee: Man Energy Solutions Se
Priority date: 2018-06-08
Filing date: 2019-04-24
Publication date: 2019-12-12
Also published as: DE102018113737A1

Abstract

In a method for carrying out catalytic gas phase reactions in a reactor system having a first reactor stage, a condenser, a heating zone and a further reactor stage, which are all arranged in series in the flow direction of a reaction gas flowing through the reactor system, it is proposed to provide, as the first reactor stage, a tube bundle reactor which has a bundle of reaction tubes which are filled with a first catalyst charge and are cooled by a heat transfer medium, and to split the reaction gas flowing into each reaction tube into at least two part flows which are introduced into the first catalyst charge at different points in the axial direction of the reaction tube. Also proposed is a reactor system for carrying out the method.

Description

Process and reactor system for carrying out catalytic gas phase reactions

The invention relates to a method and a reactor system for carrying out catalytic gas phase reactions according to the preamble of claims 1 and 11, respectively. In particular, the invention relates to the performance of exothermic catalytic gas phase reactions with pronounced temperature maxima. An example of such reactions are methanation reactions.

Against the background of finite natural gas deposits, there is an increasing interest in the production of a natural gas substitute. This gas, referred to as "substitute natural gas" or "substitute natural gas", essentially contains methane and minor amounts of unreacted and other gases which must comply with the feed-in specification of the relevant gas network if the SNG is there is to be fed.

SNG can be produced on the basis of coal or other carbonaceous substances such as waste or biomass via synthesis gas - a mixture of CO and / or CO2, H2 and possibly other constituents such as water. In another method, methanation is used for chemical storage of e.g. Excess electricity used, which arises in the generation of electricity by renewable energy. By electrolysis hydrogen is generated with this stream. Carbon dioxide is preferably recovered from exhaust gases from industrial processes or from biogas plants. The methanation feed gas in this case consists almost exclusively of CO2 and H2.

The methanation is decisively determined by the following chemical reactions: CO methanation:

CO + 3H ₂ <-> CH ₄ + H ₂ O AH _r ° = -206 kJ / mol (1)

002 methanation ("Sabatier reaction")

CO2 + 4 H2 <- CH ₄ + 2 H ₂ O AH _R ° = -165 kJ / mol (2)

Water gas shift reaction ("WGS reaction")

CO + H ₂ O <- H ₂ + CO ₂ AH _r ° = -41 kJ / mol (3)

In addition, a variety of other side reactions take place, including the Boudouard reaction, in which carbon monoxide decomposes to carbon and carbon dioxide. Boudouard reaction:

2 CO <-> C + C0 ₂ DH _K ° = -173 kJ / mol (4)

The reactions are i.d.R. catalyzed with elements of VIII. Subgroup, preferably with nickel.

All reactions are in equilibrium with one another, whereby the position of the equilibria can be changed by suitable measures. From the reaction equations for the position of the equilibrium the dependencies of the essential process parameters pressure, temperature and concentrations of the reaction components arise:

Equations (1) and (2) result in a decrease in volume on the product side. Both reactions are also highly exothermic. This causes a high pressure and a low temperature to shift the reaction equilibrium to the product side. As pressure increases, however, its influence on the reaction equilibrium becomes smaller and smaller. In addition, the wall thicknesses of all pressurized parts of the apparatus are getting ever larger and demands on the tightness, reinforced by the small molecular size of the hydrogen, rise. The reaction temperature is limited downwards by the minimum working temperature of the catalyst.

Although an increasing reaction temperature increases the reaction rate and decreases according to the reactor. However, a disadvantage is a limitation of the conversion due to a shift in the reaction equilibrium on the educt side and increased aging of the catalyst.

The reaction equilibrium can be further shifted to the product side by removing the product component water H2O from the reaction system. However, this promotes soot formation by the Boudouard reaction according to equation (4). By a minimum content of water, this reaction is suppressed. Favorable is a feed gas with a nearly stoichiometric composition, otherwise the superstoichiometric component remains. This is unproblematic as long as this component is still within the specification of the feed gas. In the other case, a separation is necessary, which leads to increased costs.

In the case of heterogeneously catalyzed exothermic gas phase reactions, a reaction zone with particularly pronounced evolution of heat often forms in the initial region of the catalyst bed, as a result of which a more or less pronounced temperature profile with a maximum in this reaction zone forms in the catalyst bed in the direction of flow. The range of the temperature maximum is also called "hot spot". This maximum temperature is desirable per se, as long as it does not exceed a certain limit. The high reaction temperature leads to a high reaction rate, which shortens the length of the catalyst bed up to the reaction equilibrium. However, in exothermic equilibrium reactions at high temperature the conversion decreases. To increase the conversion then the temperature of the reaction gas must be reduced.

Especially in methanation reactions, this hot spot is very pronounced. If this is not sufficiently controlled, then it can lead to an excessive increase in temperature with the result of increased aging of the catalyst and associated sales reduction. This requires more frequent costly catalyst changes. In addition, there is a risk of damage to thermometers, if they are installed in individual reaction tubes.

Various methods and reactor systems were used to take account of these reaction-technical boundary conditions.

The gases participating in the chemical reactions are referred to below as "feed gas", as long as they are still before entering a reactor system. From entry into a reactor system, the gases are referred to as "reaction gas", from the outlet as "product gas".

A widely used method for methanation are reactor systems with at least two reactor stages, wherein the first reactor stage is a high-temperature methanation in which high reaction rates are achieved at high temperature. In a second reactor stage, the reaction is continued in a low-temperature methanation to the desired conversion.

Often, several adiabatic catalytic fixed bed reactors are connected in series with intermediate cooling stages. If the feed gas is fed completely to the first reactor, it is often diluted with circulating gas from downstream reaction stages to avoid excessively high reaction temperatures. Before entering the reactor, the feed gas is tempered with a preheater to the minimum working temperature of the catalyst. A corresponding method is shown in DE2549439A. In a process presented in EP21 10425A1, the feed gas is split and fed in parallel to the entry of a plurality of adiabatic fixed bed reactors connected in series. The feed gas can additionally be diluted by adding steam. The process is exemplarily operated at about 35 bar, the inlet temperature in the reactors is between 240 ° C and 300 ° C, the outlet temperature at about 600 ° C. The heat of reaction is used in-process for gas tempering and in other parts of the plant. It has also been proposed to embed coils in successive reactors into the catalyst bed of one of these reactors in order to cool the reaction gas more effectively. Such isothermal methanation reactors with boiling water cooling in combination with adiabatic reactors are described in DE2705141 A1 or in DE2940334A1. In the latter publication, the process is run with hydrogen excess. The unreacted hydrogen is removed from the product gas and returned to the feed gas.

For controlling the temperature of exothermic catalytic gas phase reactions in a shell and tube reactor with catalyst-filled reaction tubes, GB472629 proposes to continuously or stepwise alter one or more of the surface, thickness, or activity of the catalyst, gas velocity, turbulence, or flow area. So z. B. in areas with strong heat development, a correspondingly large reduction of the flow cross-section required, which in turn has correspondingly high pressure losses result. Such pressure losses necessitate more powerful compressors and thus lead to increased investment and operating costs.

In another approach to control the hot spot in a Rohrbündeiraktor proposed in US3268299, in each reaction tube coaxially einzu a metering tube in the catalyst bed and initiate the reaction gas mixture in the metering tube. The metering tube extends from the gas inlet-side end of the reaction tube into the catalyst filling, is free of catalyst material and has a plurality of axially distributed outflow openings, all of which open into the catalyst bed. As a result, the reaction is distributed over the length of the reaction tube, whereby the reaction intensity is reduced and the tem peratu ranstieg is reduced. On the other hand, the metering tube causes a reduction of the hydraulic diameter of the catalyst bed and thus a reduction of the radial transport path for the heat of reaction and thus their faster removal. In another embodiment, a reaction gas component A can be fed into the metering tube and a reaction gas component B in the beginning of the Katalysatorschüttu ng.

DE 102010040757 A1 proposes a tube bundle reactor whose catalyst-filled reaction tubes provided with metering tubes extend through two heat transfer zones, the outflow openings of the metering tubes all opening into the catalyst bed within the first heat transfer zone. The reaction gas is passed in each reaction tube simultaneously both in the inlet of the catalyst bed and in the metering. Although this construction allows an optimal reaction by the stepped reaction gas addition and optimum temperature control by independent heat transfer circuits, however it is quite expensive if in the second heat transfer zone only little heat is generated.

DE1645840A proposes a two-stage methanation process in which after the first methanation stage so much water vapor is removed from the reaction gas that carbon deposition on the catalyst is prevented. In US2009 / 0247653A1, such a partial condensation is carried out after the reaction gas has passed through two methanization reactors. In the generic DE102009059310A1 such partial condensation is carried out between these two reactors in two reactors connected in series with external cooling.

The invention has for its object to improve a generic method and reactor system so that catalytic gas phase reactions with pronounced temperature maxima safer, faster and can be carried out with a small size and that at the same time the life of the catalyst is extended.

According to the invention, this object is achieved by a method according to claim 1 and a reactor system according to claim 11.

By the measures according to the invention the effectiveness of catalytic gas phase reactions is considerably improved with increased reliability, because it is a more precise control or management of the reaction temperature and achieved a higher throughput, while the structural design of the second reactor stage can be significantly simplified. The basic idea is to optimize one reaction parameter in each individual reactor stage. The aim of the first reactor stage is a controlled high reaction rate. As a result, the reactor dimensions are minimized and extended the life of the catalyst. In the second reactor stage, the residual conversion takes place at low temperature. At this residual conversion only little heat of reaction is generated. To fulfill this function, a structurally simple adiabatic catalytic fixed bed reactor is usually sufficient. Because the first reactor stage is a tube bundle reactor, in which a metering tube is embedded in the catalyst bed of the reaction tubes, the interior of which is catalyst-free, not only is the radial heat transport path shortened, but also the flow area of the catalyst bed in the flow direction of the reaction gas is reduced and reduced Reduced flow cross-section reaction gas also introduced at axially different locations as part of streams in the catalyst bed, whereby the heat load of the catalyst bed is made uniform and thus the temperature maximum is reduced. In addition, the reaction gas flowing out of the gas outlet points of the metering tube - ie the corresponding the partial streams - diluted by the already flowing in the catalyst bed and upstream reaction gas. Instead of a particularly pronounced temperature maximum in the initial region of the catalyst bed so occur in the flow direction of the reaction gas several smaller temperature maxima on. These specific advantages of a tube bundle reactor with the division of the reaction gas into partial streams or with metering tubes arranged in the reaction tubes make use of the invention. By lowering the height of the absolute temperature maximum which occurs, ie the maximum temperature occurring in the catalyst bed, the conversion increases, as a result of which increased space velocities are possible.

Due to the increased amount of sales in the first reactor stage only a small residual conversion is required in the second reactor stage, which can be carried out in a structurally simple fixed bed reactor, z. In an adiabatic reactor.

A further advantage of the method according to the invention is that the need for recycling partial product streams for cooling or dilution is eliminated.

The method is suitable for all heterogeneously catalyzed exothermic gas phase reactions that form a hot spot in the beginning of the catalyst bed. In particular, the invention is suitable for those reactions in which this hot spot is particularly pronounced and is correspondingly difficult to control.

Preferably, the catalytic gas phase reaction includes methanation reactions as an example of such reactions. For the methanation reaction according to equation (2), the feed gas for carrying out the process according to the invention preferably consists of CO ₂ and H ₂ . When using CO2 and H2 as components of the feed gas in stoichiometric composition, after water separation at the end of the process, a product gas having a composition according to the equilibrium conditions at a given pressure and temperature is to be expected. In fact, the methane content will be slightly lower, perhaps because the reactant gas is not exactly in stoichiometric composition. If the feed gas composition is not stoichiometric, one or the other component will always remain after the reaction. The choice of the feed gas composition depends on the respective feed conditions for the SNG. If, for example, the hydrogen H2 is to be converted as completely as possible, the process is preferably run in a range between the stoichiometric composition of CO2 / H2 and a slight excess of CO2. The preferred range of ratios of CO2 / H2 in this case is between 0.25: 1 and 0.26: 1. In the design of the second reactor stage, the main objective is to achieve a methane concentration according to the feed specification. This is achieved by a high conversion at a relatively low temperature. The lower temperature limit is determined by the minimum working temperature of the catalyst, also referred to as "light-off temperature". The upper temperature limit is determined by the temperature resistance of the catalyst and the turnover to be maintained. Conveniently continues to have a high pressure.

As the material for the catalyst, all known unsupported catalysts and shell catalysts for methanation come into question. These usually contain elements of VIII. Subgroup, preferably nickel.

In a favorable embodiment of the invention, the reaction gas is cooled to a temperature below the dew point of water in step c) and at least a portion of the condensed water derived. Due to the partial condensation and dissipation of the reaction product water between the first reactor stage and the second reactor stage, the reaction equilibrium is shifted to the product side in a favorable manner.

In this case, the reaction gas after deriving at least a portion of the condensed water advantageously contains a residual water vapor content of 0% to 30% and preferably from 15% to 25% by keeping a residual water vapor in the reaction gas soot formation is prevented, thus stabilizing the process. Before entering the second reactor stage, the reaction gas is heated to something above the light-off temperature analogous to the feed gas in a preheater. With this residual steam content, an optimum is achieved with respect to the shift of the reaction equilibrium on the product side and the prevention of soot deposits.

In an advantageous development of the invention, the first catalyst bed has at least two catalyst layers of different activity, wherein the activity of the first catalyst layer in the flow direction of the reaction gas is less than the activity of the at least one further catalyst layer and in step b) a first partial flow into the first catalyst layer and each further partial stream is introduced past the first catalyst layer into the at least one further catalyst layer. The term catalyst activity is known to the person skilled in the art and is described in more detail, for example, in the article "Reschetilowski W .: Introduction to Heterogeneous Catalysis, Springer Verlag Berlin Heidelberg 2015, pp. 11-20, DOI 10.1007 / 978-3-662-46984-2J2 , These measures are based on the finding that, although the amount of reaction gas entering the catalyst bed is locally reduced by the partial streams and thus the heat of reaction liberated is controlled and also the downstream in the catalyst bed entering reaction gas partial streams are diluted by the already flowing in the catalyst bed and upstream reaction gas, however, there is in the partial flow, which occurs directly above the first downstream partial flow into the catalyst bed, such a reaction controlling effect not. Here it is so that the undiluted reaction gas reacts after a short inlet section under the influence of an undiluted catalyst and forms a pronounced hot spot. This hot spot can be controlled to a certain extent by appropriate dimensioning of the annular gap between the metering tube and the reaction tube. As the annular gap narrows, the radial path to the cooled reaction tube wall initially decreases. However, the reduced flow cross-section leads to a smaller hydraulic diameter and thus to a greater pressure loss. To maintain a constant flow rate, the inlet pressure must be increased and the gas flow in the dosing tube throttled. The result is increased investment costs and operating costs for the compressor. Since the catalyst volume is reduced in the ring space, the reaction tube must be extended accordingly to get to the equilibrium state. Alternatively, an increase in the number of reaction tubes is possible. In any case, the ratio of the hydraulic diameter of the annular gap to the particle size of the catalyst is smaller with the result of increased Randgängigkeit and correspondingly undefined reaction conditions. The disadvantages of a more powerful compressor and an excessive increase in the catalyst volume are now avoided by the present invention, the activity of the first catalyst layer is reduced. Since the chemical reactions take place under the influence of catalysts, as a precondition, first of all the conversion decreases. On the other hand, however, there are several influences that lead to an increase in sales and thus approximately cancel the sales-reducing effect of reduced catalyst activity. The reduced catalyst activity initially results in less heat of reaction, which leads to a reduction in the height of the temperature of the hot spot. A reduced gas temperature now reduces the volume flow and thus the dependent on the gas velocity pressure loss, which in turn has an increased mass flow through the first catalyst layer result. It establishes a balance with the parallel gas flow inside the metering tube. As described above, sales are now rising at the same time as the temperature drops. The increase of the gas fraction in the first catalyst layer also has the consequence that the reaction gas entering the catalyst from the first gas exit point undergoes an increased dilution, whereby the subsequent hot spot is controlled even more effectively. The same applies to the subsequent gas outlet points or reaction sections. In sum, the conversion at the end of the reaction tube thus remains essentially constant. In this way, the temperature control in the initial region of the catalyst bed is significantly improved, with the conversion, the flow cross-section and the pressure loss remain essentially unchanged. The decisive advantage of the measure according to the invention is that the improved Control the temperature of the hot spots in the first catalyst layer the life of the

Catalyst is significantly increased. Thus, at a same change interval of the catalyst, the tube bundle reactor of the reactor system according to the invention has a significantly higher catalyst activity than a tube bundle reactor with undiluted catalyst, so that in a given production period through the reactor system more product can be produced.

By replacing the catalyst material of the first catalyst layer, the reaction conditions can be adapted relatively easily to different requirements and in particular to changing conditions in the entire reactor system. Thus, a particularly effective temperature control is possible at any time. Thus, in the first region of the first catalyst layer, a controlled hot spot, damped according to a specification, is formed. This sounds out and the temperature in the catalyst layer changes only slightly. This last range of the small temperature change preferably accounts for a proportion of 10 to 40% of the first catalyst layer.

The conversion of the educts at the end of the first reactor stage is in the range of 90% to 99%. The residual conversion takes place in the second reactor stage. The conversion in the first reactor stage is limited by the reaction product water because of the equilibrium reaction. Due to the partial condensation after the first reactor stage, water is removed from the reaction system, so that the reaction equilibrium is further shifted to the product side. Due to the high conversion of the first reactor stage, the heat of reaction in the second reactor stage is so low that i.d.R. a simple adiabatic fixed bed reactor is sufficient for the remaining methanation.

The activity of the first catalyst layer is preferably adjusted to 5% to 90% and particularly preferably to 10% to 40% of the activity of the at least one further catalyst layer. By this measure, the temperature level of the hot spot is safely and defined damped, whereby the turnover remains at a high level and the life of the catalyst bed is extended. If temperature measuring devices are installed in reaction tubes, they are protected against damage. This ensures a secure and permanent process monitoring.

In a further embodiment of the invention, the first and the second catalyst bed can each be subdivided into even further catalyst layers having different catalyst activities, the catalyst activity preferably increasing from catalyst layer to catalyst layer. In an advantageous embodiment of the invention, the reaction parameters in the first reactor stage, in particular the heat carrier temperature, the bed heights of the individual catalyst layers, the axial distances of introduction points of the partial streams, the partial flow amounts and the catalyst activation are adjusted so that the maximum temperatures in the first and second catalyst layer in a range of 300 ° C to 900 ° C, preferably between 500 ° C and 700 ° C. The controlled high temperatures within the first catalyst bed lead to an optimum in terms of reaction rates and conversion.

As a result, this reaction section requires only a relatively small reaction space. The stepped arrangement of the introduction points in the first reactor stage achieves both a high controlled reaction rate with a correspondingly small reaction volume and a high conversion in the range of lower temperatures. In this way, the required total catalyst volume can be minimized.

The heat of reaction generated in the process is preferably used in the process and, more preferably, also in the apparatus.

As a heat carrier come because of the temperature ranges of several species in question. Preference is given to boiling water. However, the wall thicknesses in a tube bundle reactor with a cladding diameter in the range of 5 to 8 meters at a corresponding pressure can be so large that the tube bundle reactor can not be built economically. Therefore, on a case by case basis, the cooling may be more favorable with a heat transfer medium which is liquid under operating conditions. Here, the pressureless operation offers, for example with liquid salt, ionic liquids or a thermal oil.

As materials for the reactor system are usually all customary in pressure vessel construction heat-resistant steels into consideration, such. the material with the short name 16Mo3 or similar materials. Depending on the location and the type of stress, other materials may also be used.

To control the temperature of individual reaction tubes are equipped with thermometers, which allow a Temperaturmessu ng along the reaction tubes.

Preferably, the reaction gas is passed through exactly two reactor stages. As already mentioned above, the reaction parameters can be adjusted so that in the second reactor stage, the remaining reaction of the starting materials takes place and thus the total conversion of the reactants is completed and another reactor stage is no longer required. This means especially at Methanisierungsreaktionen that after the second reactor stage, the product gas after condensing out of the water content meets the feed specification of the respective gas network into which it is to be fed.

Preferably, the entire reactor system with a space velocity (GHSV) of 5000 1 / h to 20,000 1 / h, particularly preferably from 8000 1 / h to 15000 1 / h, operated. Here, the space velocity "GHSV" means the "gas hourly space velocity". It is based on the bulk volume of the total catalyst contained in the reactor system Normvolu menstrom, here refers to the bulk volume on catalyst beds of any kind - whether undiluted or diluted - relates. A high space velocity leads to a smaller reactor volume at the same reactor power. This results in reduced investment costs and constructive advantages by reducing space requirements.

Preferably, the reaction gas is introduced into the first reactor stage at a pressure of from 5 bara to 50 bara and more preferably from 10 bara to 30 bara and most preferably from 15 bara to 25 bara. This pressure range is optimal in terms of influencing the reaction equilibrium, taking into account the design effort, which consists primarily in the strength design, and the investment and operating costs for compression of the feed gas at the inlet of the reactor system.

Preferably, the heat carrier temperature of the first reaction stage is adjusted so that it is in the range of 240 ° C to 300 ° C. This temperature range is optimal with regard to influencing the reactions in terms of reaction rate and conversion. In some cases, the inventive method can also be carried out at temperatures in the range between 200 ° C and 350 ° C. Decisive for the lower temperature is always the light-off temperature of the catalyst, which should not be undercut.

The reactor system according to the invention has several components which influence one another. Preferably, the individual components are designed using a simulation program. Such a simulation program can be a commercially available program or a program which can be created with known process engineering contexts. The catalysts have different activities depending on the composition. The parameters of the catalyst to be used are determined for this purpose in laboratory experiments and thus adapted to the simulation parameters. The sizes of the first and second reactor stages influence each other. The larger the first reactor stage, ie the more catalyst volume is present there, the greater is the conversion there. The second reactor stage is correspondingly smaller. Conversely, a smaller first reactor stage has a larger second reactor level required. In optimization calculations, the reactor system can, for example, for a minimum total amount of catalyst under given boundary conditions to mass flow rate and

Methane concentration can be optimized.

If the conversion of the first reaction stage is too low, a critical operating case may occur in the second reactor stage in which so much heat of reaction is released that the temperature rises so high that first the catalyst is damaged and secondly the conversion below a specified one Value drops, which is normally at least equal to the feed-in specification. There is also the risk of damage to the reactor.

In order to create a reserve in the event of a decrease in the catalyst activity of the first reaction stage in the course of the operating time, the second reaction stage is therefore preferably dimensioned so that even with a decrease in the conversion in the first reaction stage to a worst-case expected value of the limit in the second Reaction stage, which leads to a critical operating case is not achieved.

When using the reactor system according to the invention for methanation reactions, reaction tubes having an outside diameter in the range from 20 mm to 100 mm, preferably from 20 mm to 40 mm, have proven particularly useful with wall thicknesses in the range of 1.5 mm to 3.0 mm. The metering tube preferably has an outer diameter in a range from 6 mm to 85 mm, particularly preferably 6 mm to 15 mm, with a wall thickness in the range from 1.0 mm to 2.0 mm.

The cross-sectional shape of the metering tubes is arbitrary. Thus, the cross section may for example be circular or oval or quadrangular. Several parallel metering tubes in a reaction tube are possible.

The shape of the usable particles for the catalyst is not particularly limited. In principle, all known shaped bodies can be used, e.g. Balls, pellets, saddles or cylinder rings. The particles should have good flow behavior and not settle after filling, i. The height of the bed should ideally not change during operation. They should be such that the flow behavior in the annular space between the reaction tube inner wall and Dosierrohraußenwand only slightly deviates from the flow behavior of a particle bed with a large flow area. With the flow behavior in particular the pressure loss and the Randgängigkeit meant.

Particularly preferred are cylinder pieces with a diameter in the range of 1, 2 mm to 3.0 mm, a length in the range of 3.0 mm to 8.0 mm and a length / diameter ratio in the range of 1: 1 to 8: 1. In this embodiment, the particles are preferred for use as a full catalyst. For the second reactor stage, an adiabatic catalytic fixed bed reactor is preferably used. Its simple structural design increases the efficiency of the reactor system.

Preferably, the first reactor stage is flowed through from bottom to top, d. H. the reaction tubes are flowed through from bottom to top. Accordingly, the metering tube of a reaction tube is attached at the lower end and ends within the catalyst bed. By this arrangement, the filling of the catalyst is greatly simplified in the reaction tubes, since the upper ends of the reaction tubes in this arrangement are free of internals.

In a favorable development of the invention, the first catalyst bed has at least two catalyst layers, wherein the activity of the first catalyst layer in the direction of flow of the reaction gas is less than the activity of the at least one further catalyst layer and the first gas outlet point in the flow direction of the reaction gas opens into the at least one further catalyst layer , As already mentioned above, by reducing the activity of the first catalyst layer, the chemical reactions and thus the resulting heat of reaction are further attenuated in a manner which, by exchanging the catalyst material of the first catalyst layer, relatively easily meets different requirements and in particular changing conditions throughout Reactor system is customizable. Thus, a particularly effective temperature control is possible at any time. Thus, in the first region of the first catalyst layer, a controlled hot spot, damped according to a specification, is formed.

The reduction of the catalyst activity is preferably achieved by diluting the catalyst with an inert material. The dilution ratio of catalyst to inert material is preferably in a range of 1: 1 to 1:10, and more preferably in a range of 1: 2.5 to 1: 4. The ratios relate to the bulk volumes of the individual components.

By diluting the catalyst with inert particles to achieve a particular activity, one is very flexible in choosing a suitable catalyst.

If a catalyst change should be necessary, it will often be restricted to the area with the highest reaction temperatures, ie to the area with the dilute catalyst. Here the likelihood of injury is greatest.

The degree of dilution of the first catalyst layer with the aim of a reduced catalyst activity depends in particular on the basic activity of the undiluted catalyst, the flow rate, the GHSV and the inlet temperature. The size and shape of the catalyst particles and inert particles are usually about the same. As a result, segregation of the two types of particles is prevented.

In a further embodiment of the invention, catalyst particles and inert particles are different in size and / or in shape. This embodiment can be particularly advantageous when e.g. Inert particles in the nature of catalyst particles are difficult to obtain or inert particles of a different kind are significantly more economical. There may also be combinations of different features with which properties such as pressure loss, heat conduction or heat transfer can be specifically optimized. The individual parameters of this training can not be determined in advance. They must be examined and determined separately in individual cases.

In the case of deviations from the shape and size of the two particle types to one another, a filling method is preferably used, which is particularly designed to prevent segregation.

In the reactor system according to the invention, the proportion of the first catalyst layer on the first catalyst bed is preferably 5% by volume and 50% by volume.

The proportion of the first catalyst layer with reduced activity on the total catalyst volume is the result of the process simulation. In general, a proportion in a range of 10 vol .-% to 35 vol .-% of the total catalyst volume of the first catalyst bed. In the first catalyst layer and in the initial region of the second catalyst layer, a high conversion is achieved at high temperature. The following hot spots following each additional gas discharge point are preferably weaker in their height.

The at least two catalyst layers preferably contain the same catalyst material and the first catalyst layer contains 5% by volume to 90% by volume, preferably 10% to 40% by volume, of the catalyst material of the at least one further catalyst layer. In this way, only a single type of catalyst needs to be procured. With this one catalyst, the catalyst activity of the first catalyst layer can be adjusted flexibly by dilution with inert material. In practice, process optimization is first performed by numerical simulation. This results in a first, possibly even relatively further optimal range of the degree of dilution. Further optimization can be achieved through pilot plant trials. Here, the first, already determined area is checked and, if necessary, further optimized, ie defined more precisely. Advantageously, at least two of the units of the first reactor stage, condenser, heating zone and second reactor stage form a structural unit. If the first and second reactor stages are arranged in a reactor housing, then the two reactor stages can be arranged either sequentially or in parallel. The combination of individual units in a constructive unit leads to a compact production plant. More compact design units increase the clarity of the system and lead to cost savings. In internal heat utilization heat losses are minimized. In some cases, the second reactor stage can also be designed as an isothermal catalytic reactor. This can be advantageous if both reaction stages are arranged in a reactor housing.

In a favorable embodiment of the invention, the second reactor stage is a cooled reactor whose average heat transfer temperature is 0 K to 30 K lower than the heat carrier temperature of the first reactor stage. This temperature range is optimal in terms of reaction rate, recoverable conversion and overall size of the individual units of the reactor system.

In this case, the second reactor stage is preferably a tube bundle reactor with a bundle of reaction tubes which are filled with the second catalyst bed and, during operation, are flowed through by a reaction gas and cooled by a heat carrier. This variant increases the flexibility of the overall system, because it is possible to reduce the conversion of the first reactor stage and to increase the heat of reaction in the second reactor stage accordingly. This procedure may be useful when e.g. two identical reactors represent an alternative to two different reactors for constructional or economic reasons or offer common heat utilization.

Particularly preferably, the reaction tubes of the two reactors are located in a common heat carrier space. With the arrangement of the two reactor stages in a common heat transfer space, a reactor is saved. The reactions of both reactor stages can be controlled simultaneously. Nevertheless, very different adapted reaction conditions can be set by the metering tubes in the first reactor stage.

In an advantageous embodiment of the invention, the ratio of the annular gap between the inner wall of the reaction tube and the outer wall of the metering tube to the particle diameter of the first catalyst layer in the range of 2 to 6 in each reaction tube. This ratio is particularly advantageous in terms of handleability, heat transfer and catalytic activity. Preferably, the axial distance between the at least one gas inflow and the first gas outflow in the flow direction of the reaction gas, the axial distance between the Gasausströmstellen and the axial distance between the last Gasausströmstelle and the end of the metering tube and the number chosen so that a heating surfaces belastu ng by the liberated reaction heat between the Gaseinström- or Gasausströmstellen in the range of 10 kW / m ² to 150 kW / m ² , preferably in the range of 20 kW / m ² to 50 kW / m ² results. Reference surface here is the tube outer surface. With this boundary condition with regard to the heat production, a favorable ratio to the heat removal is achieved, which in turn leads to advantageous reaction temperatures.

In a favorable development of the invention, the quadratic ratio of the inner diameter of the reaction tube to the outer diameter of the metering tube is in a range from 2 to 6. With these preferred ratios, an annular gap results, which particularly meets the requirements for a reaction control.

The invention will be explained in more detail by way of example with reference to the drawings.

Show it:

FIG. 1 shows a diagram of a first exemplary embodiment of a reactor system according to the invention,

FIG. 2 shows, on an enlarged scale, a sectional view through a first exemplary embodiment of a reaction tube from a tube reactor of the first reactor stage,

3 shows a view similar to FIG. 2 of a second exemplary embodiment of a reaction tube from a tube reactor of the first reactor stage,

FIG. 4 shows a diagram of a second exemplary embodiment of a reactor system according to the invention, with first and second reactor stages in the same reactor housing,

FIG. 5 shows a vertical section through a reactor housing of a third exemplary embodiment of a reactor system according to the invention, with first and second reactor stages in the same reactor housing,

6 shows a vertical section through a tube bundle reactor of a fourth embodiment of a reactor system according to the invention and

Figure 7 is a vertical section through a reactor housing of a fifth embodiment of a reactor system according to the invention, with first and second reactor stage and capacitor and heating zone in the same reactor housing.

The illustrated in Figure 1 embodiment of a reactor system 1 according to the invention for carrying out the method according to the invention comprises a preheater 2, a first reactor stage 3 with a first catalyst bed 4, a condenser 5, a compressor 6, a Aufheizzone 7 and a second reactor stage 8 with a second catalyst bed 9, which are all arranged in the flow direction of a reactor system 1 by flowing through the feed gas 10 and the reaction gas 1 1 one behind the other.

The first reactor stage 3 is a tube bundle reactor 12 which has a bundle 13 with a plurality of catalyst-filled reaction tubes 14 which extend vertically between an upper tube plate 15 and a lower tube plate 16 and are enclosed by a cylindrical jacket 17. The two ends of the reaction tubes 14 are gas-tight welded to the respective tube sheet 15,16. In operation, the reaction tubes 14 of the reaction gas 1 1 - in the illustrated embodiment, from top to bottom - flows through and cooled by a heat carrier 18. The upper tube sheet 15 is spanned by a gas inlet hood 19 and the lower tube plate 16 be of a gas outlet haul 20.

In FIGS. 2 and 3, catalyst-carrying reaction tubes 14 with metering tube 21 are shown. The reaction tubes 14 and metering tubes 21 are not shown to scale there. The length / diameter ratio of the reaction tubes 14 and metering tubes 21 is much larger in reality.

In Figure 2, the reaction tubes 14 are filled with the first catalyst bed 4, wherein in each reaction tube 14 coaxial with this a metering tube 21 is arranged, the interior 22 is free of catalyst. The metering tube 21 extends by a predetermined length into the reaction tube 14 and is fixed by spacers 23 in its central position. The metering tube 21 is attached in a manner not shown here at its gas inlet end 24 on the reaction tube 14 or on the adjacent upper, gas inlet side tube plate 15. Between the inner wall 25 of the reaction tube 14 and the outer wall 26 of the metering tube 21, an annular gap 27 is formed whose size is dimensioned so that the square ratio of the inner diameter 28 of the reaction tube 14 to the outside ßend diameter 29 of the metering tube 21 in a range of 2 to 6 lies.

In the gas outlet end region 30 of each reaction tube 14, a catalyst holder 31 is arranged, on which the catalyst bed 4 rests. This extends from the catalyst holder 31 to the metering tube 21 and then further into the annular space 32 between the reaction tube 14 and metering 21 into a predetermined distance from the gas inlet end 33 of the reaction tube 14. At the gas inlet end of the reaction tube 33, a catalyst-free space 34 is formed , This prevents the way bubbles of catalyst material, because at the gas inlet of the reaction tubes 14, turbulence in the reaction gas stream 1 1 can form. In addition, in the manner not shown here in the lower region of the free space 34 a gas-permeable member, such as a perforated plate or a wire mesh, are arranged to fix the catalyst in its position.

The ratio of the annular gap 27 between the inner wall 25 of the reaction tube 14 and the outer wall 26 of the metering tube 21 to the particle diameter of the first catalyst bed 4 is in the range of 2 to 6.

The gas inlet-side end 24 of the metering tube 21 forms an inflow opening 35 for the reaction gas 11. At a predetermined axial distance from the gas inlet end 36 of the first catalyst bed 4, d. H. downstream of the free space 34, a first Gasausström- point 37 A is arranged and in the example shown at further predetermined axial distances to this still a second and a third, last Gasausströmstelle 37 B, 37 C. A gas outflow point is formed by one or more gas outflow openings 38, which are preferably distributed uniformly around the circumference of the metering tube 21. The Gasausströmöffnungen 38 a Gasausströmstelle 37 A, 37 B, 37 C may be offset in the circumferential direction against each other. The size and number of these Gasausströmöffnungen 38 are dimensioned so that sets a predetermined gas flow. Furthermore, 21 throttle openings may be arranged in the manner not shown here within the metering tube.

The metering tube 21 has three functions. It reduces the flow cross-section of the catalyst bed 4 in the reaction tube 14 to the cross-section of the annular space 32 between the reaction tube 14 and metering tube 21, whereby the heat of reaction is reduced. Further, the reaction gas 11 is introduced axially stepped, whereby the total heat of reaction is divided into several smaller proportions. And finally, transverse to the flow direction of the reaction gas 11, the heat dissipation path is shortened to the size of the annular gap 27 between the reaction tube 14 and metering tube 21, so that not only less reaction heat is created, but this is also derived even faster.

The axial distance between the gas inlet-side end 36 of the first catalyst bed 4 and the first gas outlet 37A, the axial distances between the Gasausströmstellen 37A, 37B, 37C and the axial distance between the last Gasausströmstelle 37C and the downstream end 39 of the metering tube 21 are set so that a Heizflächenbelastung by the liberated heat of reaction between the Gaseinström- or Gasausströmstellen 35, 37A, 37B, 37C in the range of 10 kW / m ² to 150 kW / m ² , preferably in the range of 20 kW / m ² to 50 kW / m ² , yields.

In particular, the distance between the gas inlet-side end 36 of the first catalyst bed 4 and the first gas outlet 37A is calculated using simulation calculations set such that a reaction temperature profile with a temperature maximum, the hot spot, forms within this distance.

The flow path of the reaction gas 11 through the first reactor stage 3 will now be described with reference to FIGS. 1 to 3.

First, the feed gas 10 before it enters the first reactor stage 3, d. H. in the tube bundle reactor 12, heated in the preheater 2 to at least the so-called light-off temperature of the catalyst of the first catalyst bed 4. This is the temperature necessary for a catalytic reaction to take place. Preferably, the feed gas 10 is adjusted to a temperature which is 5 K to 30 K higher than the light-off temperature. It is also possible to heat in the tube bundle reactor 12 directly in front of the first catalyst bed 4. As a result, a part of the system is saved.

The preheated feed gas 10 enters the gas inlet hood 19 of the tube bundle reactor 12 and is distributed there - now referred to as reaction gas 11 - to the reaction tubes 14th

In each reaction tube 14, the reaction gas stream entering there is subdivided into a first substream 11.1, which enters directly into the first catalyst bed 4 located in the annular space 32 between metering tube 21 and reaction tube 14, and into a second substream 11.2 which flows into the inflow opening 35 of the metering tube 21 enters and is there bypassed to the first catalyst bed 4, until it from the Gasausströmstellen 37A, 37B, 37C as part streams 1 1.2A,

11.2B, 1 1.2C exits into the first catalyst bed 4. The gas inlet-side end 36 of the catalyst bed 4 thus forms the point of introduction for the partial stream 11.1 and the gas outlet 37A, 37B, 37C form the points of introduction for the partial streams 1 1 2A, 11 2B, 11 2C in the catalyst bed 4th

The first, directly introduced into the annulus 32 Teilstroml 1.1 flows through the space 34 - reaction tube section 14.1 - and then enters the reaction tube section 14.2 in the first catalyst bed 4 a.

The reaction of the first substream 11.1 is well advanced to the downstream end of the reaction tube section 14.2, ie, before the first gas outflow point 37A, which is to be understood as a degree of conversion in the range of about 70% to 85%. The required axial extent of the reaction tube section 14.2 is determined by means of simulation calculations. The subsequent reaction tube section 14.3 is a series of gas outflow points 37A, 37B, 37C and subsequent reaction sections. It starts with the first gas outlet 37 A and ends with the downstream end 39 of the metering tube 21. The added in the Gasausströmstellen 37A, 37B, 37C partial streams 11 .2A, 11 .2B, 11.2C of the reaction gas 11 react in the respective subsequent catalyst bed 4th to the next Gasausströmstelle 37 B, 37 C and to the downstream end 39 of the dosing 21st The axial distances of the Gasausströmstellen 37 A, 37 B, 37 C are dimensioned so that the reaction of the respective added partial streams 11 .2A, 11 .2B, until the beginning of the next Gasau sströmstel le 37 B, 37 C has advanced, including here also a degree of conversion in the range of about 70% to 85%. In this case, the conversion of the reaction sections at their downstream ends in the downstream direction from reaction section to reaction section increases. This increase in sales is the result of the increasing degree of dilution and the increasing heat dissipation due to the ever-increasing flow velocity. Both effects favor the temperature control. At the end of the last reaction section between the last Gasausströmstelle, here 37C, and the downstream end 39 of the metering tube 21, the degree of conversion is in the range of about 80% to 90%.

After the downstream metering tube end 39, a last reaction tube section 14.4 connects with the last part of the first catalyst bed 4, in which a residual conversion takes place up to the reaction equilibrium. Because of the small amount of heat of reaction generated there, the effects of the dosing tube 21 are no longer needed. Due to the larger flow cross-section, the residence time of the reaction gas 11 is extended, which makes it all the more possible to carry out the still required residual conversion. At the end of the reaction tube section 14.4 then the conversion is in the range of 90% to 99%.

Figure 3 shows an embodiment in which the first catalyst bed 4, d. H. the catalyst bed of the first reactor stage 3, is divided into two catalyst layers 4a, 4b. The first catalyst layer 4a in the flow direction of the reaction gas 11 has a lower activity than the subsequent second catalyst layer 4b.

The first catalyst layer 4a begins at the gas inlet-side end 36 of the first catalyst bed 4 in the annular space 32 and ends upstream of the first gas outlet 37A. Compared to FIG. 2, the reaction tube section 14.2 there is thus subdivided into a reaction tube section 14.2a, which contains the first catalyst layer 4a (with lower activity), and a reaction tube section 14.2b, which contains part of the second catalyst layer 4b.

Preferably, the downstream end 40 of the first catalyst layer 4a is at a predetermined distance upstream of the first gas discharge point 37A second catalyst layer 4b by this distance upstream of the first Gasausströmstelle 37A and forms the reaction tube section 14.2b. This reaction tube section 14.2b fulfills the main function of reacting the reaction gas 11 almost to the reaction equilibrium until it reaches the first gas discharge point 37A. Another function is to provide a safety zone in the event of possible settling of the catalyst, so that it is ensured that the reaction gas 11, which flows through the first gas outlet 37 A in the catalyst bed 4, flows into an undiluted catalyst layer 4 b and so on optimal conditions.

The reaction forms a temperature profile within the first catalyst layer 4a with a hot spot and has progressed far to the downstream end 40 of the first catalyst layer 4a, which is to be understood as a degree of conversion in the range of about 70% to 85%. In the subsequent second catalyst layer 4b, the reactions are carried out as described with reference to FIG.

The two catalyst layers 4a, 4b contain the same catalyst material, the first catalyst layer 4a being a mixture comprising 5% by volume to 90% by volume, preferably 10% by volume to 40% by volume of the catalyst material of the second catalyst layer 4b and otherwise contains inert material. The particle sizes of the catalyst material and the inert material are preferably the same in order to avoid segregation, in particular when filling the reaction tubes 14.

In addition, with reference to FIG. 1, the reaction gas 11 flows out of the reaction tubes 14 into the gas outlet hood 20 of the tube bundle reactor 12 and from there out of the first reactor stage 3 to the condenser 5.

In the condenser 5, the reaction gas 11 is cooled to a temperature below the dew point of at least part of the constituents of the reaction gas 11, and a part of the condensed constituents 41 are discharged. In the case of methanation, the water formed during the reactions in the tube bundle reactor 12 is partially condensed out and discharged into lines.

The reactions in the first reactor stage 3 and the temperature decrease in the condenser 5 lead to a decrease in volume. To produce optimum reaction conditions to maximize the conversion in the second reactor stage 8, the pressure is subsequently increased in the compressor 6. In the heating zone 7, the compressed reaction gas 1 1 is heated to at least the light-off temperature of the second catalyst bed 9 of the second reactor stage 8. In the illustrated embodiment, the second reactor stage 8 is an adiabatic reactor 42. In the adiabatic reactor 42, the reaction of the reaction gas 11 is completed by achieving a high conversion at a relatively low temperature. In the case of methanation, the product gas 43 now consists largely of methane CH ₄ with a methane concentration corresponding to a given feed specification.

As a rule, following the second reactor stage 8, the water produced in the process is almost completely condensed out.

In the exemplary embodiment of a reactor system 101 according to the invention shown in FIG. 4, the first and second reactor stages 3, 8 are designed as a structural unit in the form of a so-called combined reactor 100.

Combined reactor 100 has in this embodiment both for the first reactor stage 3 reaction tubes 114A and for the second reactor stage 8 reaction tubes 114B, which are combined in a single tube bundle 113 and distributed mixed there. The tube bundle 113 is enclosed by a reactor jacket 117, so that the reaction tubes 114A,

114B of the first and the second reactor stage 3, 8 are located in a common heat carrier space 118A in which they are cooled by a heat carrier 118. Further, the reaction tubes 114A, 14B of both the first and second reactor stages 3, 8 are filled with catalyst material and are flowed through by reaction gas 11 during operation, with the reaction tubes 114A of the first reactor stage 3 in fluid communication with another gas distribution space and gas collection space Reaction tubes 114 B of the second reactor stage eighth

The reaction tubes 1 14A of the first reactor stage 3 are filled with the first catalyst bed 4 and are flowed through from top to bottom. They are attached at their ends with a first upper and a first lower tube sheet 1 15A, 116A gas-tight. In each reaction tube 114A of the first reactor stage 3, a metering tube 21 is arranged coaxially, which is at least partially embedded in the first catalyst bed 4. The metering tubes 21 and this first catalyst bed 4 can be designed, for example, as shown in FIGS. 2 and 3.

The reaction tubes 114B of the second reactor stage 8 are filled with the second catalyst bed 9 and are flowed through from bottom to top. They are at their ends with a second upper tube sheet 1 15B, which is spaced above the first upper tube sheet 115A, and in a second lower tube sheet 116B, which is spaced below the first lower Tube bottom 116A is arranged gas-tight. The reaction tubes 14B of the second reactor stage 8 are filled exclusively with the second catalyst bed 9 and contain none

Dosing tubes.

The second upper tube sheet 1 15B is spanned by an upper reactor hood 1 19 and the second lower tube sheet 116B by a lower reactor hood 120.

The space between the first upper and second upper tubesheet 1 15A,

115B, the (first) gas inlet space 119A and the space between the first lower and second lower tubesheet 16A, 16B form the (first) gas collection space 120A for the reaction tubes 114A of the first reactor stage 3.

The lower reactor hood 120 forms the (second) gas inlet space 1 19B and the upper reactor hood 119 forms the (second) gas collecting space 120B for the reaction tubes 14B of the second reactor stage 8.

The exemplary embodiment of a reactor system 101 according to the invention shown in FIG. 4 for carrying out the method according to the invention is operated as follows:

A gas stream with a first feed gas component 10a and a gas stream with a second feed gas component 10b, in the case of methanation, for example, a gas stream of CO2 and a gas stream of H2, are combined into a feed gas 10 in a mixer, not shown here. This is heated to the light-off temperature of the first catalyst bed 4 before it enters the first reactor stage 3 during ongoing stationary operation in countercurrent in a preheater 102 by the product gas 43 coming from the second reactor stage 8. When starting the reactor system 101, however, there is still no hot product gas 43. The feed gas 10 is therefore heated during the starting process by a preferably electrically operated starting preheater 102.1. The operation of this Anfahrvorwärmers 102.1 is possible in any way - for example, with steam, if a steam network is available. The diversion of the feed gas stream 10 takes place by corresponding opening or closing of shut-off valves 150 in the pipeline.

Subsequently, the preheated feed gas 10 flows into the (first) gas distribution chamber 119A of the first reactor stage 3 of the combined reactor 100 and from there into the reaction tubes 114A with the metering tubes 21. After emerging from these reaction tubes 14A, the reaction gas 11 is passed into the first gas collection chamber 120A. From there it is fed into a heat exchanger 151, where it gives off a first part of its heat.

Subsequently, the water 41 contained in the reaction gas 11 - in the case of methanation - is partially condensed and discharged in a condenser 105, so that the reaction gas 11 only contains a water content of about 20% by volume. The cooled reaction gas 11 is then heated again in the heat exchanger 151 before it is conducted into the (second) gas distribution chamber 1 19B of the second reactor stage 8 of the combined reactor 100. From there, the reaction gas 11 flows into the reaction tubes 14B of the second reactor stage 8, where the residual conversion takes place towards the desired product gas composition. The reaction gas 11 then enters from these reaction tubes 14B into the second gas collection chamber 120B in the upper reactor hood 119, from where it is led out of the combined reactor 100 as product gas 43.

In the subsequent preheater 102 mentioned above, the product gas 43 heats the feed gas 10 entering the first reactor stage 3 and is subsequently conducted into a second condenser 152 where the water 153 still formed in the second reactor stage 8 is condensed out and discharged. The dry product gas 43 is then transferred to the gas feed unit. In the case of methanation, the degree of conversion is in the range of 98.0% to 99.6%, so that the product gas 43 corresponds to a feed specification with a given methane concentration. The same product quality can be achieved with other embodiments, as shown for example in FIG. 1 or subsequently in FIG. 7.

The combination reactor 100 shown is a boiling water reactor. The operating pressure is 55 bara, resulting in an operating temperature of 270 ° C. In order for the heat management to be as efficient as possible, at least some of the feed water is used as the cooling medium for the condensers. In the illustrated embodiment, the feed water 155 supplied by means of a pump 154 is used as a coolant in the condenser 105 between the first and second reactor stages 3, 8 before being supplied to the steam drum 156 of the coolant circuit 157 of the boiling water reactor 100.

The reactor system 101 is further equipped with a measuring, control and control system, not shown here, as is customary for systems of this type. In particular, the composition of the reaction gas 1 1 is analyzed between the first and second reactor stage 3.8. In this case, it is checked whether the conversion of the first reactor stage 3 is sufficiently high, so that the residual conversion in the second reactor stage 8 does not lead to unacceptably high temperatures. Temperatu rmessstellen are still appropriate in all process streams and in the second reactor stage 8. The process streams include the feed gas components 10a, 10b before and after the mixture, the reaction gas 11, the heat carrier 118, condensed out water 41,

43, feedwater 155 for the boiling water circuit 157 and any other auxiliary currents.

FIG. 5 shows a variant 200 of the combined reactor 100 shown in FIG. 4. The difference is that the reaction tubes 214A, 214B of the first and second reaction stages 3, 8 are not distributed uniformly over the reactor cross section, but are arranged in separate regions , Thus, the reaction tubes 214A of the first reactor stage 3 are arranged in an annular tube bundle 213A, which encloses a central tube bundle 213B with the reaction tubes 214B of the second reactor stage 8.

As in the exemplary embodiment according to FIG. 4, the reaction tubes 214A of the first reactor stage 3 each contain a coaxially arranged metering tube 21 which is at least partially embedded in the first catalyst bed 4. The reaction tubes 214B of the second reactor stage 8 are filled exclusively with the second catalyst bed 9 and contain no metering tubes.

The upper reactor hood 219 forms the (first) gas distributor space 219 A of the first reactor stage 3 and has a gas inlet nozzle 250 A for the feed gas 10. The lower reactor hood 220 forms the (first) gas collecting space 220A of the first reactor stage 3 and has a gas outlet nozzle 251 A for the reaction gas 11.

The (second) gas distribution chamber 219B of the second reactor stage 8 is formed by a distributor hood 252, which within the lower reactor hood 220, d. H. within the first gas collection chamber 220A, and is connected to a second gas inlet nozzle 250B that extends through and out of the lower reactor cap 220. The second gas collection chamber 220B is formed by a collecting hood 253 which is located within the upper reactor hood 219, d. H. within the first gas distribution space 219A, and is connected to a second gas exit port 251B which extends through and out of the upper reactor cap 219.

The distributor hood 252 is fastened to the side of the lower tube bottom 216 facing the lower reactor hood 220. The collecting hood 253 is fastened on the side of the upper tube bottom 215 facing the upper reactor hood 219.

To compensate for manufacturing tolerances and thermal expansions, the second gas inlet nozzle 250B and the second gas outlet nozzle 251 B each have a strain compensator 254. The operation is as follows:

The feed gas 10 is introduced into the first gas distribution chamber 219A and from there enters the reaction tubes 214A of the first reactor stage 3.

From the reaction tubes 214A of the first reactor stage 3, the reaction gas 11 enters the first gas collection chamber 220A and is discharged from the combination reactor 200 by means of the first gas flow strut 251A.

Subsequently, the reaction gas 1 1 is cooled so far in a condenser, not shown here, that in the reaction gas 1 1 - in the case of methanation - contained water is condensed and a portion of the water is discharged. Thereafter, the reaction gas 1 1 is heated in a heating zone, also not shown here again to the light-off temperature of the second catalyst bed 9 of the second reaction stage 8.

The heated reaction gas 11 re-enters the combination reactor 200 through the second gas inlet port 250B and is supplied to the second gas distribution space 219B (the distributor hood 252).

From the distributor hood 252, the reaction gas 1 1 enters the reaction tubes 214B of the second reactor stage 8.

From these reaction tubes 214B, the reaction gas 11 enters the second gas collecting space 220B (the collecting hood 253), and is discharged from the combination reactor 200 by means of the second gas discharge port 251B as the product gas 43.

It is transferred to the gas feed unit and possibly before a further condenser for condensing out of the water still contained in the reaction gas 1 1 after the second reactor stage 8 and in the now existing product gas 43.

By the embodiments of the second reactor stage 8 presented in FIGS. 4 and 5 as a tube bundle reactor integrated in the combined reactor 100, 200, it is possible to control the overall reaction more effectively and to optimize the size of the entire reactor system 101. The particular advantage of these reactor systems 101 is that overheating in the second reactor stage 8 can be ruled out by the cooling and the conversion is kept at a stable and high level, at the same time achieving a compact construction. FIG. 6 shows an exemplary embodiment of the first reactor stage 3. Here is the

Shell and tube reactor 12 a radiator 350 directly connected so that together form a constructive unit 300 together.

The tube bundle reactor 12 has a bundle 13 reaction tubes 14 which are filled with the first catalyst bed 4 and in each of which a metering tube 21 is at least partially embedded in the first catalyst bed 4. The reaction tubes 14 with the metering tubes 21 are flowed through from top to bottom. They correspond to the embodiment shown in Figure 3.

The upper, gas inlet-side ends of the reaction tubes 14 are in an upper tube sheet

Attached gas-tight, which is spanned by an upper reactor hood or gas inlet hood 19 with a gas inlet nozzle 351.

The lower, gas inlet-side ends of the reaction tubes 14 are in a lower tube sheet

16 attached gas-tight, on which on the side facing away from the reaction tubes 14 side of the radiator 350 is attached.

The reaction tubes 14 are enclosed by a cylindrical reactor jacket 17, which together with the upper and lower tubesheet 15, 16 forms a heat carrier chamber 18A.

In the heat carrier chamber 18A, the reaction tubes 14 are flowed around by a heat carrier 18, the heat carrier 18A supplied from a lower annular channel 352 and in the heat transfer chamber 18A through disc and annular baffles 353 meandering in each radial direction from outside to inside and vice versa and from below above - d. H. transverse to the reaction tubes 14 and in countercurrent to the reaction gas 1 1 - passed through the tube bundle 13 and is derived from an upper annular channel 354 again from the heat transfer chamber 18A. The annular channels 352, 354 enclose the reactor jacket 17 on the outside thereof and are in fluid communication with the heat carrier chamber 18A through jacket openings 355. The heat transfer medium 18 is circulated via a heat exchanger and a heat transfer pump, both of which are not shown here. Preferably, the heat carrier 18 is a well-known liquid salt, but other heat transfer can be used, such as thermal oil or ionic liquids.

The radiator 350 has a bundle 356 of cooling tubes 357, the number of which may be smaller than the number of reaction tubes 14. The cooling tubes 357 extend from an upper radiator tube bottom 358 vertically to a lower radiator tube bottom 359 and are cylindrical Cooler coat 360 enclosed. The ends of the cooling tubes 357 are connected to the respective radiator Tube bottom 358, 359 gas-tight welded. The cooling tubes 357 are catalyst-free and are flowed through by the reaction gas 11 from top to bottom.

Between the lower tube sheet 16 of the tube bundle reactor 12 and the upper cooler tube bottom 358, a gas transfer chamber 361 is arranged, into which the reaction gas 11 exits the reaction tubes 14 and from which it enters the cooling tubes 357.

The cooling tube bundle 356 is traversed transversely by a heat carrier 362 in a heat carrier circuit independent of the heat transfer medium circuit of the tube bundle reactor 12. The heat transfer fluid enters through an in the lower end region of the radiator 350 arranged inlet nozzle 363 in the radiator 350 and through an upper end of the radiator 350 arranged outlet nozzle 364 from this again. Halfway up the radiator 350, a baffle 365 is arranged between the inlet and the outlet nozzle 363, 364, which extends horizontally through the entire cooling tube bundle 356, so that the heat carrier flow coming from the inlet nozzle 363 and after its deflection to the outlet nozzle 364 out across the entire cooling tube bundle 356 is performed.

The lower radiator tube plate 359 is spanned on its side facing away from the cooling tubes 357 by a lower reactor hood or a gas outlet hood 20 with a gas outlet nozzle 366.

From the cooling tubes 356, the reaction gas 11 enters the gas outlet hood 20 and is discharged through the gas outlet nozzle 366 from the first reactor stage 3 and passed to a capacitor, not shown here.

In Figure 7, a combined reactor 400 is shown as an exemplary embodiment, in which all components of the reactor system are housed in a housing. The reaction gas 11 flows through this combined reactor 400 from bottom to top. In detail, the combi-reactor 400 contains the following components:

A lower reactor hood 420 forms a gas distributor space 419A, into which the feed gas preheated in a preheater 450 enters.

The gas distribution chamber 419A is followed by a shell-and-tube reactor 412 as the first reactor stage 3. The reaction tubes 414 of the tube bundle reactor 412 contain the first catalyst bed 4, in which metering tubes 21 are at least partially embedded, similar to that shown in FIG. 2 or 3. However, the metering tubes 21 are rotated relative to Figure 2 or 3 by 180 °, so that their Gas inlet opening in each case with the gas distribution chamber 419A in flow communication. A heat transfer medium 18 is hereby conducted in cocurrent with the reaction gas 11 through the heat transfer space 418A via a heat carrier inlet connection 451 and a heat transfer outlet connection 452 in a heat transfer circuit. The heat carrier 18 may, for. As liquid salt or boiling water or pressurized water.

Above the tube bundle reactor 412 of the first reactor stage 3, a working space 453 is arranged for a mechanic. The working space 453 is accessible through a manhole 454 and serves to fill the reaction tubes 414 with catalyst material during the initial filling or to change the catalyst material and for other work, such. B. the installation of thermometers.

Above this working space 453 there is a device 455 for collecting condensed constituents 456 of the reaction gas 11 - in the case of methanation, ie water - and for diverting the condensed constituents 456. The said device 455 is penetrated by at least one covered passage 457 for reaction gas 11 so that the reaction gas 11 can flow from the first reactor stage 3 in the combined reactor 400 further up to the still following components. The roof 458 prevents condensed components from entering the working space 453 or the first reactor stage 3.

Above this area, a condenser 405 connects to a bundle 459 catalyst-free cooling raw re 460, which are traversed by the reaction gas 11, which is cooled so far that a portion of the components - such as water - condenses. The condensed components 456 flow down into the cooling tubes 460 into the collecting and diverting regions 455, 456. The condenser 405 has an inlet connection 461 and an outlet connection 462 for the coolant 463, which here is passed countercurrently to the reaction gas flow through the condenser 405 , As the coolant 463, for example, water can be used.

Connected to the condenser 405 is a heating zone 407, in which the reaction gas 11 leaving the condenser 405 is heated to the reaction temperature of the second catalyst bed 9 in the second reactor stage 8. In the illustrated embodiment, the tubes 460 extend continuously through the condenser 405 and the heating zone 407 and are free of catalyst material. Via an inlet connection 464 and an outlet connection 465, heat transfer medium 466 is passed through the heating zone 407 in countercurrent to the reaction gas flow. The heat transfer medium 466 may, as in the first reactor stage 3 z. As liquid salt or boiling water or pressurized water. From the heating zone 407, the reaction gas 11 enters the second reactor stage 8, which is an adiabatic reactor 442 in the illustrated embodiment. It contains the second catalyst bed 9 in which the residual reaction of the reaction gas 11 takes place.

From the second reactor stage 8, the reaction gas 11 enters the upper reactor hood 419, which forms a gas collection chamber 420A, from which the finished reaction gas emerges as product gas 43. If the product gas 43 is still too moist, for example during methanation, it can be fed to a further condenser 467, in which this residual moisture 468 is condensed out.

The embodiments are not limiting in the embodiments of the invention. In particular, individual features of a particular embodiment in an adapted embodiment can be advantageously used in other embodiments. For example, it is also possible for the crude reactor to flow through from bottom to top in FIG. 1, as shown in FIG. 7.

Claims

claims

1

Process for carrying out catalytic gas phase reactions in a reactor system having a first reactor stage, a condenser, a heating zone and a second reactor stage, which are arranged one behind the other in the flow direction of a reaction gas flowing through the reactor system, comprising the following steps:

a) providing an isothermal reactor having a first catalyst bed as the first reactor stage;

b) passing the reaction gas through the first catalyst bed;

c) cooling the reaction gas effluent from the isothermal reactor to a temperature below the dew point of at least a portion of the constituents of the reaction gas and diverting at least a portion of the condensed constituents;

d) heating the reaction gas, from which part of the condensed constituents was derived in step c);

e) providing a reactor with a second catalyst bed as the second reactor stage; and

f) passing the heated reaction gas through the second catalyst bed; characterized in that

step a) comprises providing a tube bundle reactor (12, 100, 200, 412) comprising a bundle (13, 13, 213A) of reaction tubes (14, 114A, 214A, 414) which fills with the first catalyst bed (4) and are cooled by a heat carrier (18), and in step b) the respective reaction gas (11) flowing into a reaction tube (14, 114A, 214A, 414) is divided into at least two partial flows (11, 1, 1, 1.2), in the axial direction of the reaction tube (14, 114A, 214A, 414) are introduced at different locations in the first catalyst bed (4).

Second

A method according to claim 1, characterized in that the catalytic gas phase reaction includes methanation reactions.

Third

A method according to claim 1 or 2, characterized in that in step c) the reaction gas (11) is cooled to a temperature below the dew point of water and at least a portion of the condensed water is discharged.

4th

A method according to claim 3, characterized in that the reaction gas (11) after deriving at least a portion of the condensed water (41, 456) contains a residual water vapor content of 0% to 30% and preferably from 15% to 25%.

5th

Method according to one of the preceding claims, characterized in that the first catalyst bed (4) has at least two catalyst layers (4a, 4b) of different activity,

wherein the activity of the first catalyst layer (4a) in the flow direction of the reaction gas (11) is less than the activity of the at least one further catalyst layer (4b) and in step b) a first partial flow (11.1) into the first catalyst layer (4a) and each further Partial flow (1 1.2) is introduced past the first catalyst layer (4a) into the at least one further catalyst layer (4b).

6th

A method according to claim 5, characterized in that the activity of the first catalyst layer (4a) is adjusted to 5% to 90%, preferably to 10% to 40%, of the activity of the at least one further catalyst layer (4b).

7th

A method according to claim 5 or 6, characterized in that the reaction parameters in the first reactor stage (3), in particular the heat carrier temperature, the bed heights of the individual catalyst layers (4a, 4b), the axial distances of introduction points (36, 37A, 37B, 37C) the partial flows (11.1, 11 2A, 1 1 2B, 11 .2C), the partial flow rates and the catalyst activation are set so that the maximum temperatures in the first and second catalyst layer (4a, 4b) in a range of 300 ° C to 900 ° C, preferably from 500 ° C to 700 ° C, lie.

8th.

Method according to one of the preceding claims, characterized in that the entire reactor system with a space velocity (GHSV) of 5000 1 / h to 20,000 1 / h, preferably from 8000 1 / h to 15000 1 / h, operated.

9th

Method according to one of the preceding claims, characterized in that the reaction gas (11) in the first reactor stage (3) at a pressure of 5 bara to 50 bara and preferably from 10 bara to 30 bara and more preferably from 15 bara to 25 bara initiated becomes.

10

Method according to one of the preceding claims, characterized in that the heat carrier temperature of the first reactor stage (3) is adjusted so that it is in the range of

240 ° C to 300 ° C.

11

Reactor system for carrying out the method according to one of claims 1 to 10, comprising a first reactor stage, a condenser, a heating zone and a second reactor stage, which are arranged one behind the other in the flow direction of a reactor system flowing through the reaction gas, wherein

the first reactor stage is an isothermal reactor with a first catalyst bed, through which the reaction gas flowing into the reactor system can be passed,

the condenser is set up to cool the reaction gas flowing out of the isothermal reactor to a temperature below the dew point of at least a part of the constituents of the reaction gas and to remove at least part of the condensed constituents,

- The heating zone is adapted to heat the reaction gas from which a part of the condensed components has been removed in the condenser, and

the second reactor stage is a reactor with a second catalyst bed, through which the reaction gas flowing out of the heating zone can be passed,

characterized in that

the first reactor stage (3) is a tube bundle reactor (12, 100, 200, 412) having a bundle (13, 113, 213A) reaction tubes (14, 114A, 214A, 414) which are filled with the first catalyst bed (4) and in the operation of the reaction gas (11) flows through and are cooled by a heat carrier (18), and that

in each reaction tube (14, 114A, 214A, 414) coaxially to this a metering tube (21) is at least partially embedded in the first catalyst bed (4) whose interior (22) is catalyst-free and which from the gas inlet end (36) of the first Catalyst bed (4) extends a predetermined length into this and outside of the catalyst bed (4) at least one Gaseinströmstelle (35) and in the region of the first catalyst bed (4) at least one Gasausströmstelle (37A), wherein in the flow direction of the reaction gas (11) first gas outlet (37 A) is arranged at a predetermined distance to the gas inlet end (36) of the first catalyst bed (4), so that there is a partial stream (11.2) of each in a reaction tube (14, 114A, 214A, 414) flowing reaction gas (11 ) flows in the bypass to the first catalyst bed (4).

12

Reactor system according to claim 11, characterized in that the first catalyst bed (4) has at least two catalyst layers (4a, 4b) of different activity, wherein the activity of the flow direction of the reaction gas (11) first catalyst layer (4a) is lower than the activity of at least a further catalyst layer (4b) and in the flow direction of the reaction gas (11) first gas discharge point (37A) in the at least one further catalyst layer (4a) opens.

13th

Reactor system according to claim 12, characterized in that the at least two catalyst layers (4a, 4b) contain the same catalyst material and the first catalyst layer (4a) 5 vol.% To 90 vol.%, Preferably 10 to 40 vol.%, of the catalyst material of the at least one further catalyst layer (4b).

14th

Reactor system according to one of Claims 11 to 13, characterized in that at least two of the units of the first reactor stage (3), condenser (5), heating zone (7) and second reactor stage (8) form a structural unit (100, 200, 400).

15th

Reactor system according to one of claims 11 to 14, characterized in that the second reactor stage (8) is a cooled reactor whose average heat transfer temperature is 0 K to 30 K lower than the heat transfer peratu r of the reactor (12) of the first reactor stage ( 3).

16th

Reactor system according to claim 15, characterized in that the second reactor stage (8) is a tube bundle reactor with a bundle of reaction tubes, which are filled with the second catalyst bed (9) and in operation by a reaction gas (11) flows through and cooled by a heat transfer medium.

17th

Reactor system according to claim 16, characterized in that the reaction tubes (114A, 114B; 214A, 214B) of the two reactors of the first and second reactor stages (3, 8) are located in a common heat carrier space (11a).

18th

Reactor system according to one of claims 11 to 17, characterized in that at each

Reaction tube (14, 114A, 214A, 414) the ratio of the annular gap (27) between the inner wall (25) of the reaction tube (14, 114A, 214A, 414) and the outer wall (26) of the metering tube (21) to the particle diameter of the first Catalyst bed (4) is in the range of 2 to 6.

19th

Reactor system according to one of Claims 1 1 to 18, characterized in that the axial distance between the at least one gas inflow point (35) and the first gas outlet point (37A) in the flow direction of the reaction gas (11), the axial distance between the gas outflow points (37 A, 37B, 37C) and the axial distance between the last Gasausströmstelle (37C) and the end (39) of the metering tube (21) and their number are chosen so that a Heizflächenbelastung by the released heat of reaction between the Gaseinström- or Gasausströmstellen (35 , 37A, 37B, 37C) in the range of 10 kW / m ² to 150 kW / m ² , preferably in the range of 20 kW / m ² to 50 kW / m ² results.

20th

Reactor system according to one of Claims 11 to 19, characterized in that the square ratio of the inside diameter (28) of the reaction tube (14, 114A, 214A, 414) to the outside diameter (29) of the metering tube (21) is within a range of 2 to 6 lies.