DE102018113737A1 - Process and reactor system for carrying out catalytic gas phase reactions - Google Patents

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 arranged one behind the other 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, the one Has bundle reaction tubes, which are filled with a first catalyst bed and are cooled by a heat carrier, and to divide each flowing into a reaction tube reaction gas into at least two partial streams, which are introduced in the axial direction of the reaction tube at different locations in the first catalyst bed. A reactor system for carrying out the process is also proposed.

Description

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. 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 CO ₂ , H ₂ and possibly other constituents such as water. In another method, the methanation is used for the chemical storage of excess current, for example, which is generated by the generation of electricity by renewable energies. 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 CO ₂ and H ₂ .

• CO-Methanisierung: CO + 3 H2 ↔ CH₄ + H₂O ΔH_R ^O = -206 kJ/mol (1)
• CO2-Methanisierung („Sabatier-Reaktion“) CO₂ + 4 H₂ ↔ CH₄ + 2 H₂O ΔH_R ^O = -165 kJ/mol (2)
• Wassergas-Shift-Reaktion („WGS-Reaktion“) CO + H₂O ↔ H₂ + CO₂ ΔHR^O = -41 kJ/mol (3)

The methanation is decisively determined by the following chemical reactions:

• CO methanation: CO + 3 H2 ↔ CH ₄ + H ₂ O ΔH _R ^O = -206 kJ / mol (1)
• CO2 methanation ("Sabatier reaction") CO ₂ + 4H ₂ ↔ CH ₄ + 2H ₂ O ΔH _R ^O = -165 kJ / mol (2)
• Water gas shift reaction ("WGS reaction") CO + H ₂ O ↔ H ₂ + CO ₂ ΔHR ^O = -41 kJ / mol (3)

• Boudouard-Reaktion: 2 CO ↔ C + CO₂ ΔH_R ^O = -173 kJ/mol (4)

In addition, a large number of other secondary reactions take place, including the Boudouard reaction, in which carbon monoxide decomposes to carbon and carbon dioxide.

• Boudouard reaction: 2 CO ↔ C + CO ₂ ΔH _R ^O = -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 H ₂ O 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, then this is often diluted with cycle gas from downstream reaction stages to avoid too 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 in the DE2549439A shown. In an in EP2110425A1 In the process described, the feed gas is split up and fed in parallel to the inlet 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 the process for gas tempering and in other parts of the plant.

It has also been proposed to embed coils in successive reactors in the catalyst bed of one of these reactors to more effectively cool the reaction gas. Such isothermal methanation reactors with boiling water cooling in combination with adiabatic reactors are known in DE2705141A1 or in DE2940334A1 described. 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 make more powerful compressors necessary and thus lead to increased investment and operating costs.

Another approach to controlling the hot spot in a tube bundle tractor is in US3268299 proposed to embed a metering tube coaxially in the catalyst bed in each reaction tube and to introduce the reaction gas mixture into 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 to the length of the reaction tube, whereby the reaction intensity is reduced and the temperature rise 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 catalyst bed.

In DE102010040757A1 a tube bundle reactor is proposed, the catalyst-filled and provided with metering tubes reaction tubes extend through two heat transfer zones, wherein the outflow openings of the metering tubes all within the first heat transfer zone in the Opening catalyst bed. The reaction gas is passed in each reaction tube simultaneously both in the inlet of the catalyst bed and in the metering. Although this design allows optimum reaction by the stepped addition of reaction gas and optimum temperature control by independent heat transfer circuits, but it is quite expensive if only little heat is generated in the second heat transfer zone.

In DE1645840A For example, a two-stage methanation process is proposed in which after the first methanization 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 flowed through two methanization reactors. In the generic DE102009059310A1 In the case of two reactors connected in series with external cooling, such partial condensation is carried out between these two reactors.

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 from the gas discharge points of the metering tube - d. H. the corresponding part 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, therefore, several smaller temperature maxima occur in the flow direction of the reaction gas. 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 resulting absolute temperature maximum, d. H. the maximum temperature occurring in the catalyst bed, the turnover increases, whereby 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 CO ₂ and H ₂ 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, for example because the reactant gas is not exactly in stoichiometric composition is present. With non-stoichiometric feed gas composition, one or the other component remains after the reaction. The choice of the feed gas composition depends on the respective feed conditions for the SNG. If, for example, the hydrogen H _{2 is to} be reacted as completely as possible, the process is preferably run in a range between the stoichiometric composition of CO ₂ / H ₂ and a slight excess of CO ₂ . The preferred range of ratios of CO ₂ / H ₂ 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.

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 amount 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 / 978-3-662-46984-2_2 , 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 the reaction gas substreams entering the catalyst bed downstream pass through the upstream and downstream catalyst bed Reacted gas to be diluted, but there is no part of the partial flow, which is above the first downstream partial flow directly into the catalyst bed, such a reaction controlling effect. 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 in the annulus is reduced, the reaction tube must be correspondingly extended to reach 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 lifetime of the catalyst is significantly increased by the improved control of the temperature of the hot spot in the first catalyst layer. 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 rates and the catalyst activation are set 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 internally in the process and more preferably also inside 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 shell and tube reactor with a shell diameter in the range of 5 to 8 meters at become 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. The unpressurized operation, for example, with liquid salt, ionic liquids or a heat transfer oil offers here.

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 temperature measurement 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, in particular in methanation reactions, that after the second reactor stage the product gas, after condensing out 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 volume flow, here 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 in terms of 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 limit is always the light-off temperature of the catalyst, which should not be undershot.

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, i. the more catalyst volume is present there, the greater is the turnover there. The second reactor stage is correspondingly smaller. Conversely, a smaller first reactor stage requires a larger second reactor stage. In optimization calculations, the reactor system can be optimized, for example, to a minimum total amount of catalyst given given boundary conditions for mass flow rate and methane concentration.

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 have an outer diameter in the range of 20 mm to 100 mm, preferably from 20 mm to 40 mm especially proven for wall thicknesses in the range of 1.5 mm to 3.0 mm. The metering tube preferably has an outer diameter in a range of 6 mm to 85 mm, more preferably 6 mm to 15 mm with a wall thickness in the range of 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 if, for example, inert particles such as catalyst particles are difficult to obtain or inert particles of a different type 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 scheduled 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 possibly further optimized, d. H. specified 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.

The axial distance between the at least one gas inlet point and the first gas outlet point in the flow direction of the reaction gas, the axial distance between the gas outlet points and the axial distance between the last gas outlet point and the end of the metering tube and their number are preferably chosen such that there is a heating surface load due to the released heat of reaction between the gas inflow or gas outflow points 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.

• 1 ein Schema eines ersten Ausführungsbeispiels eines erfindungsgemäßen Reaktorsystems,
• 2 in vergrößertem Maßstab eine Schnittansicht durch ein erstes Ausführungsbeispiel eines Reaktionsrohrs aus einem Rohrbündelreaktor der ersten Reaktorstufe,
• 3 eine der 2 ähnliche Ansicht eines zweiten Ausführungsbeispiels eines Reaktionsrohrs aus einem Rohrbündelreaktor der ersten Reaktorstufe,
• 4 ein Schema eines zweiten Ausführungsbeispiels eines erfindungsgemäßen Reaktorsystems, mit erster und zweiter Reaktorstufe im selben Reaktorgehäuse,
• 5 einen Vertikalschnitt durch ein Reaktorgehäuse eines dritten Ausführungsbeispiels eines erfindungsgemäßen Reaktorsystems, mit erster und zweiter Reaktorstufe im selben Reaktorgehäuse,
• 6 einen Vertikalschnitt durch einen Rohrbündelreaktor eines vierten Ausführungsbeispiels eines erfindungsgemäßen Reaktorsystems und
• 7 einen Vertikalschnitt durch ein Reaktorgehäuse eines fünften Ausführungsbeispiels eines erfindungsgemäßen Reaktorsystems, mit erster und zweiter Reaktorstufe sowie Kondensator und Aufheizzone im selben Reaktorgehäuse.

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

• 1 a diagram of a first embodiment of a reactor system according to the invention,
• 2 on an enlarged scale a sectional view through a first embodiment of a reaction tube from a tube reactor of the first reactor stage,
• 3 one of the 2 Similar view of a second embodiment of a reaction tube from a tube reactor of the first reactor stage,
• 4 a diagram of a second embodiment of a reactor system according to the invention, with first and second reactor stage in the same reactor housing,
• 5 3 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 a vertical section through a tube bundle reactor of a fourth embodiment of a reactor system according to the invention and
• 7 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.

This in 1 illustrated embodiment of a reactor system according to the invention 1 for carrying out the method according to the invention has a preheater 2 , a first reactor stage 3 with a first catalyst bed 4 , a capacitor 5 , a compressor 6 , a heating zone 7 and a second reactor stage 8th with a second catalyst bed 9 on, all in the flow direction of the reactor system 1 flowing through the feed gas 10 or reaction gas 11 arranged one behind the other.

The first reactor stage 3 is a tube bundle reactor 12 who is a bunch 13 with a variety of catalyst-filled reaction tubes 14 which extends between an upper tube sheet 15 and a lower tube sheet 16 extend vertically and from a cylindrical shell 17 are enclosed. 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 from the reaction gas 11 - In the illustrated embodiment, from top to bottom - flows through and from a heat transfer medium 18 cooled. The upper tube sheet 15 is from a gas inlet hood 19 and the bottom tube sheet 16 from a gas outlet hood 20 spans.

In the 2 and 3 are catalyst filled reaction tubes 14 with dosing tube 21 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 bigger in reality.

In 2 are the reaction tubes 14 with the first catalyst bed 4 filled, with each reaction tube 14 coaxial with this a metering tube 21 is arranged, whose interior 22 is catalyst-free. The dosing tube 21 extends for a predetermined length in the reaction tube 14 in and out through spacers 23 fixed in its central position. The dosing tube 21 is 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 sheet 15 attached. Between the inner wall 25 of the reaction tube 14 and the outer wall 26 of the dosing tube 21 is an annular gap 27 formed, whose size is dimensioned such that the square ratio of the inner diameter 28 of the reaction tube 14 to the outside diameter 29 of the metering tube 21 is in a range of 2 to 6.

In the gas outlet end area 30 each reaction tube 14 is a catalyst holder 31 arranged on the catalyst bed 4 rests. This extends from the catalyst holder 31 to the metering tube 21 and then into the annulus 32 between reaction tube 14 and metering tube 21 in up to a predetermined distance from the gas inlet end 33 of the reaction tube 14 , At the gas inlet end of the reaction tube 33 thus becomes a catalyst-free space 34 educated. This prevents the blowing away of catalyst material, because at the gas inlet of the reaction tubes 14 can turbulence in the reaction gas stream 11 form. In addition, in a manner not shown in the lower part of the clearance 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 dosing tube 21 to the particle diameter of the first catalyst bed 4 is in the range of 2 to 6.

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

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

The axial distance between the gas inlet end 36 the first catalyst bed 4 and the first gas discharge point 37A , the axial distances between the Gasausströmstellen 37A . 37B . 37C and the axial distance between the last gas discharge point 37C and the downstream end 39 of the dosing 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 from 10 kW / m ² to 150 kW / m ² , preferably in the range from 20 kW / m ² to 50 kW / m ² .

In particular, the distance between the gas inlet end 36 the first catalyst bed 4 and the first gas discharge point 37A determined with the help of simulation calculations such that within this distance a reaction temperature profile with a temperature maximum, the hot spot, is formed.

In the following, the flow path of the reaction gas 11 through the first reactor stage 3 based on 1 to 3 described.

First, the feed gas 10 before entering the first reactor stage 3 ie in the tube bundle reactor 12 in the preheater 2 to at least the so-called light-off temperature of the catalyst of the first catalyst bed 4 heated. This is the temperature necessary for a catalytic reaction to take place. The feed gas is preferred 10 set at a temperature which is 5 K to 30 K higher than the light-off temperature. It is also a warming in the tube bundle reactor 12 directly in front of the first catalyst bed 4 possible. 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 distributes there - now referred to as the reaction gas 11 - on the reaction tubes 14 ,

In every reaction tube 14 is the reaction gas stream entering there into a first partial stream 11.1 that goes directly into the annulus 32 between metering tube 21 and reaction tube 14 located first catalyst bed 4 enters, and in a second partial flow 11.2 split that into the inflow opening 35 of the dosing tube 21 enters and there in the bypass to the first catalyst bed 4 is guided until it is out of the gas outflow 37A . 37B . 37C as partial streams 11.2a . 11.2b . 11.2C in the first catalyst bed 4 exit. The gas inlet end 36 the catalyst bed 4 thus forms the point of introduction for the partial flow 11.1 and the gas outlets 37A . 37B . 37C form the introduction points for the partial flows 11.2a . 11.2b . 11.2C in the catalyst bed 4 ,

The first, directly in the annulus 32 initiated partial flow 11.1 flows through the open space 34 - Reaction tube section 14.1 - And then enters the reaction tube section 14.2 in the first catalyst bed 4 on.

The reaction of the first partial flow 11.1 is to the downstream end of the reaction tube section 14.2 , ie before the first gas discharge point 37A , well advanced, which means a turnover in the range of about 70% to 85%. The required axial extent of the Reaction tube section 14.2 is determined with the help of simulation calculations.

The following reaction tube section 14.3 is a sequence of gas outlets 37A . 37B . 37C and subsequent reaction sections. It starts with the first gas discharge point 37A and ends with the downstream end 39 of the dosing tube 21 , The in the gas outflow points 37A . 37B . 37C added partial streams 11.2a . 11.2b . 11.2C of the reaction gas 11 react in the respective subsequent catalyst bed 4 until the next gas discharge point 37B . 37C or to the downstream end 39 of the dosing tube 21 , The axial distances of the Gasausströmstellen 37A . 37B . 37C are dimensioned so that the reaction of each added streams 11.2a . 11.2b , until the beginning of the next Gasausströmstelle 37B . 37C is well advanced, and this is also to be understood as 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 dosing tube 21 the conversion rate is in the range of about 80% to 90%.

After the downstream Dosierrohrende 39 closes a last reaction tube section 14.4 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 there only small resulting heat of reaction, the effects of the metering tube 21 no longer needed. Due to the larger flow cross-section, the residence time of the reaction gas is prolonged 11 , whereby all the more can be done of the still required residual sales. At the end of the reaction tube section 14.4 then the turnover is in the range of 90% to 99%.

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

The first catalyst layer 4a starts at the gas inlet end 36 the first catalyst bed 4 in the annulus 32 and ends upstream of the first gas discharge point 37A , Compared to 2 So is the local reaction pipe section 14.2 in a reaction tube section 14.2a , the first catalyst layer 4a (with less activity) and a reaction tube section 14.2b structured, which forms part of the second catalyst layer 4b contains.

Preferably, the downstream end is located 40 the first catalyst layer 4a at a predetermined distance upstream of the first gas discharge point 37A or protrudes the second catalyst layer 4b by this distance upstream over the first gas discharge point 37A and forms the reaction tube section 14.2b , This reaction tube section 14.2b fulfills the main function, the reaction gas 11 to react almost to the reaction equilibrium until it is the first Gasausströmstelle 37A reached. Another function is to provide a safety zone in the event of possible settling of the catalyst, so as to ensure that the reaction gas 11 passing through the first gas outlet 37A in the catalyst bed 4 flows into an undiluted catalyst layer 4b flows and can react under optimal conditions.

The reaction forms within the first catalyst layer 4a a temperature profile with a hot spot and is to the downstream end 40 the first catalyst layer 4a well advanced, which is understood to mean a degree of conversion in the range of about 70% to 85%. In the subsequent second catalyst layer 4b the reactions take place as based on 2 described.

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

Further flows with respect to 1 the reaction gas 11 from the reaction tubes 14 in the gas outlet hood 20 of the tube bundle reactor 12 and from there the first reactor stage 3 out to the capacitor 5 ,

In the condenser 5 becomes the reaction gas 11 to a temperature below the dew point of at least a portion of the constituents of the reaction gas 11 cooled and part of the condensed components 41 derived. In the case of methanation this becomes during the reactions in the tube bundle reactor 12 resulting water partially condensed and derived in 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 optimal reaction conditions to maximize conversion in the second reactor stage 8th the pressure is then in the compressor 6 elevated.

In the heating zone 7 becomes the compressed reaction gas 11 to at least the light-off temperature of the second catalyst bed 9 the second reactor stage 8th heated. In the illustrated embodiment, the second reactor stage 8th an adiabatic reactor 42 , In the adiabatic reactor 42 becomes the reaction of the reaction gas 11 completed by a high conversion is achieved at relatively low temperature. In the case of methanation, the product gas is 43 now largely methane CH ₄ with a methane concentration according to a given feed specification.

In general, following the second reactor stage 8th the water produced in the process almost completely condensed out.

At the in 4 illustrated embodiment of a reactor system according to the invention 101 are the first and the second reactor stage 3 . 8th as a constructive unit in the form of a so-called combined reactor 100 educated.

The combined reactor 100 indicates in this embodiment, both for the first reactor stage 3 reaction tubes 114A as well as for the second reactor stage 8th reaction tubes 114B on that in a single tube bundle 113 summarized and mixed there are distributed. The tube bundle 113 is from a reactor jacket 117 enclosed, so that the reaction tubes 114A . 114B the first and the second reactor stage 3 . 8th in a common heat transfer room 118A in which they are from a heat transfer medium 118 be cooled. Further, the reaction tubes 114A . 114B both the first and the second reactor stage 3 . 8th filled with catalyst material and are in the operation of reaction gas 11 flows through, the reaction tubes 114A the first reactor stage 3 are in flow communication with another gas distribution space and gas collection space than the reaction tubes 114B the second reactor stage 8th ,

The reaction tubes 114A the first reactor stage 3 are with the first catalyst bed 4 filled and flows through from top to bottom. They are at their ends with a first upper and a first lower tubesheet 115A . 116A attached gas-tight. In every reaction tube 114A the first reactor stage 3 is a dosing tube 21 coaxially disposed in the first catalyst bed 4 is at least partially embedded. The dosing pipes 21 and this first catalyst bed 4 For example, as in the 2 and 3 be formed represented.

The reaction tubes 114B the second reactor stage 8th are with the second catalyst bed 9 filled and flows through from bottom to top. They are at their ends with a second upper tube sheet 115B spaced above the first upper tube sheet 115A is arranged, and in a second lower tube sheet 116B at a distance below the first lower tube sheet 116A is arranged, gas-tight fastened. The reaction tubes 114B the second reactor stage 8th are exclusively with the second catalyst bed 9 filled and contain no metering tubes.

The second upper tube sheet 115B is from an upper reactor hood 119 and the second lower tube sheet 116B from a lower reactor hood 120 spans.

The space between the first upper and the second upper tubesheet 115A . 115B forms the (first) gas inlet space 119A and the space between the first lower and the second lower tubesheet 116A . 116B the (first) gas collection room 120A for the reaction tubes 114A the first reactor stage 3 ,

The lower reactor hood 120 forms the (second) gas inlet space 119B and the upper reactor hood 119 the (second) gas collection room 120B for the reaction tubes 114B the second reactor stage 8th ,

This in 4 illustrated embodiment of a reactor system according to the invention 101 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 CO ₂ and a gas stream of H ₂ , are in a mixer, not shown here to a feed gas 10 merged. This is before entering the first reactor stage 3 in running, stationary operation in countercurrent in a preheater 102 through the second reactor stage 8th upcoming product gas 43 to the light-off temperature of the first catalyst bed 4 heated. When starting up the reactor system 101 however, there is still no hot product gas 43 , The feed gas 10 is therefore during the start-up by a preferably electrically powered starting preheater 102.1 heated. The operation of this starting preheater 102.1 is possible in any way - for example, with steam, if there is a steam network. The diversion of the feed gas stream 10 done by appropriate opening or closing shut-off valves 150 in the pipeline.

Subsequently, the preheated feed gas flows 10 in the (first) gas distribution room 119A the first reactor stage 3 of the combined reactor 100 and from there into the reaction tubes 114A with the dosing pipes 21 , After exiting these reaction tubes 114A becomes the reaction gas 11 in the first gas collection room 120A directed.

From there it becomes a heat exchanger 151 where it gives off a first part of its heat.

Subsequently, this is in the reaction gas 11 - in the case of methanation - contained water 41 partly in a condenser 105 condensed and discharged, so that the reaction gas 11 only contains a water content of about 20 vol .-%. The cooled reaction gas 11 is then in the heat exchanger 151 reheated before it enters the (second) gas distribution room 119B the second reactor stage 8th of the combined reactor 100 is directed. From there, the reaction gas flows 11 into the reaction tubes 114B the second reactor stage 8th where the residual conversion takes place towards the desired product gas composition. The reaction gas 11 then emerges from these reaction tubes 114B in the second gas collection room 120B in the upper reactor hood 119 from where it is used as product gas 43 from the combined reactor 100 is led out.

In the subsequent, above-mentioned preheater 102 heats the product gas 43 in the first reactor stage 3 entering feed gas 10 and then into a second condenser 152 led, where in the second reactor stage 8th still formed water 153 is condensed and derived. 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 is achievable with other embodiments, such as those in 1 or subsequently in 7 are shown.

In the illustrated combination reactor 100 it 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 means of a pump 154 supplied feed water 155 in the condenser 105 used between the first and second reactor stage 3.8 as a coolant before it the steam drum 156 of the coolant circuit 157 of the boiling water reactor 100 is supplied.

The reactor system 101 is also equipped with a measuring and control system, not shown here, as is customary for systems of this type. In particular, the composition of the reaction gas 11 analyzed between first and second reactor stage 3.8. This controls whether the sales of the first reactor stage 3 is sufficiently high, so that the residual sales in the second reactor stage 8th does not lead to unacceptably high temperatures. Temperature measuring points are also useful in all process streams and in the second reactor stage 8th , The process streams include the feed gas components 10a . 10b before and after the mixture, the reaction gas 11 , the heat carrier 118 , condensed water 41 . 43 , Feed water 155 for the boiling water cycle 157 as well as any other auxiliary currents.

5 shows a variant 200 of in 4 shown combined reactor 100 , The difference is that the reaction tubes 214A . 214B the first and second reaction stage 3.8 are not uniformly distributed over the reactor cross-section, but are arranged in separate areas. So are the reaction tubes 214A the first reactor stage 3 in an annular tube bundle 213A arranged, which is a central tube bundle 213B with the reaction tubes 214B the second reactor stage 8th encloses.

The reaction tubes 214A the first reactor stage 3 included as in the embodiment according to 4 in each case a coaxially arranged metering tube 21 at least partially into the first catalyst bed 4 is embedded. The reaction tubes 214B the second reactor stage 8th are exclusively with the second catalyst bed 9 filled and contain no metering tubes.

The upper reactor hood 219 forms the (first) gas distribution room 219A the first reactor stage 3 and has a gas inlet nozzle 250A for the feed gas 10 on. The lower reactor hood 220 forms the (first) gas collection room 220A the first reactor stage 3 and has a gas outlet nozzle 251A for the reaction gas 11 on.

The (second) gas distribution room 219B the second reactor stage 8th is through a distributor hood 252 formed within the lower reactor hood 220 ie within the first gas collection room 220A , arranged and to a second gas inlet nozzle 250B connected, extending through the lower reactor hood 220 extends through and out of this. The second gas collection room 220B is through a collecting hood 253 formed within the upper reactor hood 219 ie within the first gas distribution room 219A , arranged and to a second gas outlet nozzle 251B connected, extending through the upper reactor hood 219 extends through and out of this.

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

To compensate for manufacturing tolerances and thermal expansions, the second gas inlet nozzle 250B and the second gas outlet 251B one expansion compensator each 254 on.

The operation is as follows:

The feed gas 10 will be in the first gas distribution room 219A initiated and enters from there into the reaction tubes 214A the first reactor stage 3 on.

From the reaction tubes 214A the first reactor stage 3 the reaction gas occurs 11 in the first gas collection room 220A and is by means of the first gas outlet nozzle 251A from the combined reactor 200 discharged.

Subsequently, the reaction gas 11 cooled so far in a condenser not shown here that in the reaction gas 11 - In the case of methanation - contained water is condensed and a part of the water is discharged. Thereafter, the reaction gas 11 in a heating zone also not shown here again to the light-off temperature of the second catalyst bed 9 the second reaction stage 8th heated.

The heated reaction gas 11 enters through the second gas inlet nozzle 250B back to the combined reactor 200 and is the second gas distribution room 219B (the distributor hood 252 ).

From the distributor hood 252 from the reaction gas occurs 11 into the reaction tubes 214B the second reactor stage 8th on.

From these reaction tubes 214B the reaction gas occurs 11 in the second gas collection room 220B (the collecting hood 253 ) and is by means of the second gas outlet nozzle 251B as a product gas 43 from the combined reactor 200 discharged.

It is passed to the gas feed unit and possibly before a further condenser to condense the still in the reaction gas 11 after the second reactor stage 8th or in the now product gas 43 contained water.

By the in 4 and 5 presented embodiments of the second reactor stage 8th than in the combined reactor 100 . 200 integrated tube bundle reactor, it is possible to control the overall reaction more effectively and size of the entire reactor system 101 to optimize. The particular advantage of these reactor systems 101 is that the cooling overheating in the second reactor stage 8th can be excluded and the turnover is kept at a stable and high level, while a compact design is achieved.

In 6 is an exemplary embodiment of the first reactor stage 3 shown. Here is the tube bundle reactor 12 a cooler 350 immediately connected, so that both together form a constructive unit 300 form.

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

The upper, gas inlet-side ends of the reaction tubes 14 are in an upper tube sheet 15 attached gas-tight, that of an upper reactor hood or gas inlet hood 19 with a gas inlet nozzle 351 is overstretched.

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 reaction tubes 14 opposite side of the radiator 350 is attached.

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

In the heat transfer room 18A become the reaction tubes 14 from a heat carrier 18 flows around the heat transfer chamber 18A from a lower annular channel 352 fed 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 bottom to top - ie transverse to the reaction tubes 14 and in countercurrent to the reaction gas 11 - through the tube bundle 13 guided and by an upper ring channel 354 again from the heat transfer room 18A is derived. The ring channels 352 . 354 enclose the reactor jacket 17 on the outside and stand with the heat transfer space 18A through shell openings 355 in Flow connection. The heat carrier 18 is circulated via a heat exchanger and a heat transfer pump, both of which are not shown here. Preference is given to the heat transfer medium 18 a well-known liquid salt, but other heat transfer can be used, such as thermal oil or ionic liquids.

The cooler 350 has a bundle 356 cooling pipes 357 whose number may be less than the number of reaction tubes 14 , The cooling pipes 357 extend from an upper radiator tube sheet 358 vertically to a lower cooler tube sheet 359 and are of a cylindrical radiator jacket 360 enclosed. The ends of the cooling pipes 357 are with the respective radiator tubesheet 358 . 359 gas-tight welded. The cooling pipes 357 are catalyst-free and are from the reaction gas 11 flows through from top to bottom.

Between the lower tubesheet 16 of the tube bundle reactor 12 and the upper radiator tube sheet 358 is a gas transfer room 361 arranged, in which the reaction gas 11 from the reaction tubes 14 exits and out of it into the cooling pipes 357 entry.

The cooling tube bundle 356 is in one of the heat transfer circuit of the tube bundle reactor 12 independent heat transfer circuit of a heat transfer medium 362 flows through it. The heat carrier flow passes through a in the lower end of the radiator 350 arranged inlet nozzle 363 in the cooler 350 one and through one in the upper end of the radiator 350 arranged outlet nozzle 364 out of this again. Halfway up the radiator 350 is between the inlet and outlet nozzles 363 . 364 a baffle 365 arranged horizontally through the entire cooling tube bundle 356 extends so that the heat transfer stream from the inlet nozzle 363 coming and after its deflection to the outlet nozzle 364 across the entire cooling tube bundle 356 to be led.

The bottom radiator tubesheet 359 is on his the cooling tubes 357 opposite side of a lower reactor hood or a gas outlet hood 20 with a gas outlet 366 spans.

From the cooling pipes 356 the reaction gas occurs 11 in the gas outlet hood 20 and is through the gas outlet 366 from the first reactor stage 3 off and led to a capacitor, not shown here.

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

A lower reactor hood 420 forms a gas distribution room 419A into a preheater 450 preheated feed gas 10 entry.

To the gas distribution room 419A closes a tube bundle reactor 412 as the first reactor stage 3 on. The reaction tubes 414 of the tube bundle reactor 412 contain the first catalyst bed 4 , into the dosing pipes 21 at least partially embedded, much like in 2 or 3 shown. The dosing pipes 21 are, however, opposite 2 or 3 rotated through 180 ° so that their gas inlet opening in each case with the gas distribution chamber 419A is in flow communication. Via a heat transfer inlet 451 and a heat transfer outlet 452 In a heat transfer circuit, a heat transfer medium 18 here in co-current with the reaction gas 11 through the heat transfer space 418A guided. The heat carrier 18 can z. As liquid salt or boiling water or pressurized water.

Above the tube bundle reactor 412 the first reactor stage 3 is a workroom 453 arranged for a fitter. The workroom 453 is through a manhole 454 accessible and used 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.

About this workroom 453 there is a device 455 for collecting condensed constituents 456 of the reaction gas 11 - In the case of methanation so of water and to divert the condensed constituents 456 , The said device 455 is covered by at least one covered passage 457 for reaction gas 11 broken, so that the reaction gas 11 from the first reactor stage 3 in the combined reactor 400 can continue to flow up to the following components. The roof 458 Prevents condensed components from entering the workspace 453 or in the first reactor stage 3 can reach.

This area is followed by a capacitor 405 with a bunch 459 catalyst-free cooling tubes 460 to that of the reaction gas 11 be flowed through, which is cooled so far that some of the components - for example water - condenses. The condensed components 456 flow in the cooling pipes 460 down into the collecting and discharge area 455 . 456 , The capacitor 405 has an inlet nozzle 461 and an outlet 462 for the coolant 463 on, here in countercurrent to the reaction gas flow through the condenser 405 is directed. As coolant 463 For example, water can be used.

To the capacitor 405 closes a heating zone 407 on, in which the from the condenser 405 exiting reaction gas 11 to the reaction temperature of the second catalyst bed 9 in the second reactor stage 8th is heated. In the illustrated embodiment, the tubes extend 460 through the capacitor 405 and the heating zone 407 and are free of catalyst material. Via an inlet 464 and an outlet 465 becomes heat carrier 466 here in countercurrent to the reaction gas stream through the heating zone 407 guided. The heat carrier 466 can 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 occurs 11 in the second reactor stage 8th a, in the illustrated embodiment, an adiabatic reactor 442 is. It contains the second catalyst bed 9 in which the residual reaction of the reaction gas 11 he follows.

From the second reactor stage 8th the reaction gas occurs 11 in the upper reactor hood 419 one who has a gas storage room 420A forms, from which the finished reaction gas as product gas 43 exit. If the product gas 43 - For example, in a methanation - is still too moist, it can be another capacitor 467 be fed, 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 that the tube bundle reactor in 1 flows through from bottom to top as in 7 shown.

QUOTES INCLUDE IN THE DESCRIPTION

This list of the documents listed by the applicant has been generated automatically and is included solely for the better information of the reader. The list is not part of the German patent or utility model application. The DPMA assumes no liability for any errors or omissions.

Cited patent literature

• DE 2549439 A [0015]
• EP 2110425 A1 [0015]
• DE 2705141 A1 [0016]
• DE 2940334 A1 [0016]
• US 3268299 [0018]
• DE 102010040757 A1 [0019]
• DE 1645840A [0020]
• US 2009/0247653 A1 [0020]
• DE 102009059310 A1 [0020]

Claims (20)

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, all of 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 with a first reactor 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, 113, 213A) of reaction tubes (14, 114A, 214A, 414) connected to the first catalyst bed (4) are filled and are cooled by a heat transfer medium (18), and in step b) the respectively in a reaction tube (14, 114A, 214A, 414) flowing reaction gas (11) into at least two partial streams (11.1, 11.2) is divided which are introduced in different directions into the first catalyst bed (4) in the axial direction of the reaction tube (14, 114A, 214A, 414). Method according to Claim 1 characterized in that the catalytic gas phase reaction includes methanation reactions. 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. 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%. 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 flow direction of the reaction gas (11) first catalyst layer (4a) is lower than the activity the at least one further catalyst layer (4b) and in step b) a first substream (11.1) into the first catalyst layer (4a) and each further substream (11.2) past the first catalyst layer (4a) into the at least one further catalyst layer (4b) is initiated. 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). 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) of the partial streams (11.1, 11.2 A, 11.2B, 11.2C), the partial flow rates and the catalyst activation are set so that the maximum temperatures in the first and second catalyst layers (4a, 4b) are in a range of 300 ° C to 900 ° C, preferably 500 ° C up to 700 ° C, lie. 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. 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. 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. Reactor system for carrying out the method according to one of Claims 1 to 10 , with a first reactor stage, a condenser, a heating zone and a second reactor stage, all of which are arranged one behind the other in the flow direction of a reaction gas flowing through the reactor system, wherein - the first reactor stage is an isothermal reactor with a first catalyst bed, through which flows into the reactor system Reaction gas is passed through, - 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 part of the reaction gas and to remove at least part of the condensed constituents, the heating zone being adapted to heat the reaction gas from the reaction gas Condenser, a part of the condensed constituents has been removed, and - the second reactor stage is a reactor with a second catalyst bed through which the effluent from the heating reaction gas is hindurchleitbar, characterized in that the first reactor stage (3) a tube bundle reactor (12, 100 , 200, 412) with a bundle (13, 113, 213A) reaction tubes (14, 114A, 214A, 414), which are filled with the first catalyst bed (4) and in operation flows through the reaction gas (11) and of a Heat carrier (18) are cooled, and that in each reaction tube (14, 114A, 214A, 414) koa xial to this a metering tube (21) is at least partially embedded in the first catalyst bed (4), the interior (22) is catalyst-free and extending from the gas inlet end (36) of the first catalyst bed (4) a predetermined length in this and at least one gas inlet point (35) outside the catalyst bed (4) and at least one gas outlet point (37A) in the region of the first catalyst bed (4), wherein the first gas outlet point (37A) in the flow direction of the reaction gas (11) is at a predetermined distance from the gas inlet end (36) of the first catalyst bed (4) is arranged so that there flows a partial stream (11.2) of each in a reaction tube (14, 114A, 214A, 414) flowing reaction gas (11) in the bypass to the first catalyst bed (4). Reactor system after 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 the at least one further catalyst layer ( 4b) and in the flow direction of the reaction gas (11) first gas discharge point (37A) opens into the at least one further catalyst layer (4a). Reactor system after 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 contains a further catalyst layer (4b). 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). 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 carrier temperature of the reactor (12) of the first reactor stage (3). Reactor system after 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. Reactor system after 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 (118a). Reactor system according to one of Claims 11 to 17 characterized in that, for 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 (14) 21) to the particle diameter of the first catalyst bed (4) is in the range of 2 to 6. Reactor system according to one of Claims 11 to 18 , characterized in that the axial distance between the at least one Gaseinströmstelle (35) and in the flow direction of the reaction gas (11) first Gasausströmstelle (37A), the axial distance between the Gasausströmstellen (37A, 37B, 37C) and the axial distance between the last Gas Ausströmstelle (37C) and the end (39) of the metering tube (21) and the number are selected 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 . Reactor system according to one of Claims 11 to 19 , characterized in that the square ratio of the inner diameter (28) of the reaction tube (14, 114A, 214A, 414) to the outer diameter (29) of the metering tube (21) is in a range of 2 to 6.

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