EP2379446A1

EP2379446A1 - Method for producing syngas

Info

Publication number: EP2379446A1
Application number: EP09764738A
Authority: EP
Inventors: Ralph Schellen; Evin Hizaler Hoffmann; Leslaw Mleczko; Stephan Schubert; Rushikesh Apte
Original assignee: Bayer Technology Services GmbH
Current assignee: Bayer Intellectual Property GmbH
Priority date: 2008-12-20
Filing date: 2009-12-04
Publication date: 2011-10-26
Also published as: US20110240925A1; WO2010069485A8; DE102008064277A1; US8758647B2; WO2010069485A1

Abstract

The present invention relates to a method for the endothermic, catalytic gas phase oxidation of hydrocarbons using water vapor and carbon dioxide into hydrogen and carbon monoxide (syngas), wherein the conversion is performed in 5 to 30 sequentially connected reaction zones under adiabatic conditions.

Description

Process for the preparation of sudnesgas

The present invention relates to a process for the endothermic, catalytic gas phase oxidation of hydrocarbons with water vapor and carbon dioxide to hydrogen and carbon monoxide (synthesis gas), wherein the reaction in 5 to 30 successive reaction zones is carried out under adiabatic conditions.

Synthesis gas consists essentially of carbon monoxide and hydrogen, but it can also contain carbon dioxide.

The essential for the production of synthesis gas from hydrocarbons partial reactions are shown in the following formulas (I to HI). The formulas refer to the conversion of methane as a hydrocarbon. For homologues of the hydrocarbon methane are correspondingly stoichiometrically corrected formulas, but which are also well known.

CH ₄ + H ₂ O <→ CO + 3 ^■ H ₂ (I)

CO + H ₂ O <→ CO ₂ + H ₂ (H)

CH ₄ + 2 ^• H ₂ O <→ CO ₂ + 4 ^• H ₂ (YY)

The reactions of formulas (I) and (HI) are highly endothermic and represent the major reactions associated with syngas production. The reaction according to formula (BT) is the reaction formula known to those skilled in the art by the term "water-gas-shift reaction" and is exothermic All three reactions according to formulas (I to JH) are equilibrium-limited.

The synthesis gas obtained from such reactions forms an essential starting material for the further reaction, for example, to tailor-made long-chain hydrocarbons according to the Fischer-Tropsch process.

The controlled supply of heat in processes for the production of synthesis gas is important, since the position of the equilibria of the abovementioned reactions according to the formulas (I to JH) is strongly dependent on the temperature of the reaction zone and thus the yields and / or selectivities with respect to hydrogen and / or carbon monoxide can be controlled thereby.

An uncontrolled drop in temperature due to the endothermic reactions according to the formulas (I) and / or (III) can thus promote the formation of more or less large amounts of carbon dioxide, which is necessary for the further use of the synthesis gas, for example for the above Fischer-Tropsch Procedure is disadvantageous. In other regions of disadvantageous temperatures, less hydrogen may be formed, which, if desired as an alternative to the production of synthesis gas, may also be detrimental. Generally, therefore, the reactions according to formulas (I to UT) must be carried out under very controlled temperature conditions in order to obtain advantageous yields and / or selectivities with respect to the desired reaction products. This is especially true if the product is to be synthesis gas.

It is therefore advantageous to keep the temperature of the reaction zones controlled in the course of the process at a level which allows a rapid conversion while minimizing the side reactions.

The abovementioned reactions according to the formulas (I to IH) do not exhaustively represent the possible reactions in a reaction zone in which synthesis gas is to be formed according to the present invention. A fairly comprehensive overview of the large number of possible reaction mechanisms involved in this process is given by A.M. De Groote and G.F. Froment in "Reactor Modeling and Simulations in Synthesis Gas Production", published in Reviews in Chemical Engineering (1995) 11: 145-183.

The process variants disclosed herein relate exclusively to reactions carried out in a fired furnace, in which shell and tube reactors are located, in which the reactions are carried out. Consequently, these are not adiabatic procedures. However, the design as a fired furnace with tube bundles is according to the method according to A.M. De Groote and G.F. Froment required.

Furthermore, A.M. De Groote and G.F. Froment discloses that this results in significant radial and axial temperature profiles in the individual reaction zones. However, especially radial temperature profiles are disadvantageous, because in this way regions exist in regions of the reaction zones which are operated under non-optimal conditions for the reaction of the hydrocarbons to synthesis gas. A sufficient control of the temperature in the reaction zones is thus not guaranteed. In addition, those of A.M. De Groote and G.F. Froment disclosed reaction devices designed extremely expensive, which is also disadvantageous, since they are at least very expensive. In case of failure but above all, a repair is possible only under Stilliegung and repair of the overall device.

Since an exact temperature control is apparently not possible, it may also happen that, for example due to the exothermic reaction according to the formula (II) local excess temperatures occur in the reaction zones, which can damage the reaction apparatus. Together with the aforementioned disadvantage of the necessarily complex construction and the so accompanying inevitable closure of the entire process in case of failure, it follows that the method disclosed by AM De Groote and GF Froment is strongly disadvantageous.

EP 1 251 951 (B1) discloses a device and the possibility of carrying out chemical

Reactions in the device disclosed, wherein the device is characterized by a cascade of contacting reaction zones and heat exchanger devices, which are cohesively arranged together in the composite. The method to be carried out herein is thus characterized by the contact of the various

Reaction zones with a respective heat exchanger device in the form of a cascade. A

There is no disclosure as to the usability of the apparatus and method for producing synthesis gas.

Thus, it remains unclear how, starting from the disclosure of EP 1 251 951 (B1), such a reaction should be carried out by means of the device and the method carried out therein. In particular, no method comprising endothermic reactions is disclosed.

Further, for the sake of uniformity, it is to be understood that the method disclosed in EP 1 251 951 (B1) is carried out in a device the same as or similar to the disclosure regarding the device. The result of this is that a significant amount of heat is transferred by conduction of heat between the reaction zones and the adjacent heat exchange zones as a result of the large-area contact of the heat exchange zones with the reaction zones.

The disclosure with regard to the oscillating temperature profile can therefore only be understood as meaning that the temperature peaks ascertained here would be stronger if this contact did not exist. Another indication of this is the exponential increase in the disclosed temperature profiles between the individual temperature peaks. These indicate that there is some heat sink of appreciable but limited capacity in each reaction zone which can reduce the temperature rise in it. It can never be ruled out that some dissipation of heat (for example by radiation) takes place; However, a reduction of the possible heat removal from the reaction zone would indicate a linear or degressive temperature gradient, since no further addition of reactants is provided and thus, after an exothermic reaction, the reaction becomes slower and thus the generated heat of reaction would decrease.

Thus, EP 1 251 951 (B1) discloses multi-stage processes in cascades of reaction zones from which heat in an undefined amount is removed by heat conduction. Thus, the disclosed process is not adiabatic and disadvantageous in that an accurate Temperature control of the reaction is not possible. This is especially true for the undisclosed possibility of an endothermic reaction in the reaction zones.

An application of the process disclosed in EP 1 251 951 (B1) to the production of synthesis gas using the devices there discloses E.L.C. Seris et al. in "Scaleable, micrstructured plant for steam reforming of methane" in Chemical Engineering Journal (2008) 135S: 9-16.

Herein, a process using apparatuses according to EP 1 251 951 (B1) is disclosed in which synthesis gas is produced in nine reaction zones with heat exchange zones therebetween. Although the presented process variant is explained as being multi-stage adiabatic, it is at the same time disclosed that the reaction zones are in direct contact with the heat exchange zones, as already disclosed in EP 1 251 951 (B1). Although this leads to an advantageous spatial integration of the reaction zones with the heat exchange zones, at the same time this means that the term adiabatic reaction zone is incorrect. The reaction zones are not adiabatic because they are in direct contact with the heat exchange zones at their boundaries and thus, especially at the considerable temperature gradients between the reaction zones and the heat exchange zones, a strong heat flow takes place, which is not due to the convenient transport of the process gases. This is in the sense of a precise temperature control, which is also the subject of E.L.C. Seris et al. presented method is disadvantageous.

Starting from the prior art, it would therefore be advantageous to provide a process for the production of synthesis gas, which can be carried out in simple reaction devices and allows a precise simple temperature control of the endothermic process, so that it high conversions at the highest possible purity of the product while maintaining the desired Yields and / or selectivities allowed. Such simple reaction devices would be easily scaled up and are inexpensive and robust in all sizes.

For the endothermic, catalytic gas-phase reaction of hydrocarbons with steam and carbon dioxide to synthesis gas, as has just been shown, no suitable methods have yet been shown which permit this.

It is therefore an object to provide a process for the endothermic, catalytic gas phase reaction of hydrocarbons with steam and carbon dioxide to synthesis gas, which is feasible under precise temperature control in simple reaction devices and thereby high conversions at high purities of the product allowed. It has surprisingly been found that a process for the production of synthesis gas from hydrocarbons, carbon dioxide and water vapor in an endothermic, heterogeneously catalytic gas phase reaction, characterized in that it comprises 5 to 30 successive reaction zones with adiabatic conditions, is capable of solving this problem.

Synthesis gas, in the context of the present invention, denotes a process gas which essentially comprises the substances carbon monoxide and hydrogen. The synthesis gas may also comprise proportions of carbon dioxide, water vapor and hydrocarbons.

Hydrocarbons in the context of the present invention refer to substances present as process gas consisting of carbon, hydrogen and optionally oxygen. Essentially, however, such hydrocarbons consist of carbon and hydrogen.

Preferred hydrocarbons which are used as feedstock in the process according to the invention are those selected from the list consisting of alkanes, alkenes and alkynes.

Particularly preferred hydrocarbons are alkanes. Preferred alkanes are those comprising at most six carbon atoms, more preferably methane, ethane, propane and butane, most preferably methane.

In the context of the present invention, steam refers to a process gas which essentially comprises water in the gaseous state.

The term essentially refers in the context of the present invention, a mass fraction and / or a molar fraction of at least 80%.

The hydrocarbons used in the process according to the invention, the steam, the constituents of the synthesis gas and the synthesis gas as such, will also be referred to collectively below as process gases.

It follows that the entire process according to the invention is carried out in the gas phase. If substances used in the process, such as the hydrocarbons, at room temperature (23 ° C) and ambient pressure (1013 hPa) are not gaseous, it is assumed in the following that such substances before or during their use in the process according to the invention by increasing the temperature and / or reducing the pressure into the gas phase.

In addition to the essential components of the process gases, these may also include secondary components. Non-exhaustive examples of minor components that may be included in the process gases include argon, nitrogen and / or carbon dioxide. According to the invention, the implementation of the process under adiabatic conditions means that the reaction zone from the outside is essentially neither actively supplied with heat nor heat withdrawn. It is well known that complete isolation against heat input or removal is possible only by complete evacuation to the exclusion of the possibility of heat transfer by radiation. Therefore, in the context of the present invention, adiabat means that no heat supply or removal measures are taken.

In an alternative embodiment of the method according to the invention, however, e.g. by isolation by means of well-known isolation means, e.g. Polystyrene insulating materials, or by sufficiently large distances to heat sinks or heat sources, wherein the insulating agent is air, a heat transfer can be reduced.

An advantage of the adiabatic driving method according to the invention of the 5 to 30 reaction zones connected in series with respect to a non-adiabatic driving mode is that no means for supplying heat must be provided in the reaction zones, which entails a considerable simplification of the construction. This results in particular simplifications in the manufacture of the reactor and in the scalability of the process and an increase in reaction conversions.

Another advantage of the method according to the invention is the possibility of very accurate temperature control by the close staggering of adiabatic reaction zones. It can thus be set and controlled in each reaction zone advantageous in the reaction progress temperature.

Yet another advantage of the process of the invention is that, unlike the prior art processes discussed above, the addition of carbon dioxide produces the desired synthesis gas, i. an increased proportion of carbon monoxide is generated. In the prior art processes, predominantly hydrogen is produced which is also a component of the synthesis gas. However, by the supply of carbon dioxide, the ratio of hydrogen to carbon monoxide can be controlled as desired.

The catalysts used in the process according to the invention are usually catalysts which consist of a material which, in addition to its catalytic activity for the reaction according to the formers (I to JE), is characterized by sufficient chemical resistance under the conditions of the process and by a high specific surface area , Catalyst materials characterized by such chemical resistance under the conditions of the process are, for example, catalysts comprising nickel or nickel compounds.

These catalysts can be applied to support materials. Such support materials usually include alumina, calcia, magnesia, silica and / or titania. Carrier materials of magnesium spinels are preferred.

Specific surface area in the context of the present invention refers to the area of the catalyst material that can be reached by the process gases, based on the mass of catalyst material used.

A high specific surface area is a specific surface area of at least 1 m ² / g, preferably of at least 10 m ² / g.

The catalysts of the invention are each in the reaction zones and can be used in all known forms, e.g. Fixed bed, moving bed, present.

Preferably, the appearance is fixed bed.

The fixed bed arrangement comprises a catalyst bed in the true sense, d. H. loose, supported or unsupported catalyst in any form and in the form of suitable packings. The term catalyst bed as used herein also encompasses contiguous areas of suitable packages on a support material or structured catalyst supports. These would be e.g. to be coated ceramic honeycomb carrier with comparatively high geometric surfaces or corrugated layers of metal wire mesh on which, for example, catalyst granules is immobilized. As a special form of packing, in the context of the present invention, the presence of the catalyst in monolithic form is considered. Such monolithic forms may also be foams of a carrier material on which the aforementioned catalyst materials have been applied.

If a fixed-bed arrangement of the catalyst is used, the catalyst is preferably present in beds of particles having mean particle sizes of 1 to 10 mm, preferably 2 to 8 mm, particularly preferably 3 to 7 mm.

Also preferably, the catalyst is in a fixed bed arrangement in monolithic form. Particularly preferred in a fixed bed arrangement is a monolithic catalyst comprising nickel compounds supported on magnesium spinels. When a catalyst in monolithic form is used in the reaction zones, in a preferred further development of the invention the monolithic catalyst is provided with channels through which the process gases flow. Usually, the channels have a diameter of 0.1 to 3 mm, preferably a diameter of 0.2 to 2 mm, more preferably from 0.5 to 1.5 mm.

If a fluidized bed arrangement of the catalyst is used, the catalyst is preferably present in loose beds of particles, as have already been described in connection with the fixed bed arrangement.

Bulks of such particles are advantageous because the particles have a high specific surface area and, because of their size, the mass transport limitation of the reaction by diffusion can be kept low. Thus, a high turnover rate can be achieved. At the same time, however, the particles are not yet so small that disproportionately high pressure losses occur when the fixed bed flows through. The ranges of the particle sizes given in the preferred embodiment of the process, comprising a reaction in a fixed bed, are thus an optimum between the achievable conversion from the reactions according to the formulas (I to ET) and the pressure loss produced when carrying out the process. Pressure loss is coupled in a direct manner with the necessary energy in the form of compressor performance, so that a disproportionate increase in the same would result in an inefficient operation of the method.

In a preferred embodiment of the process according to the invention, the conversion takes place in 7 to 20, more preferably 10 to 15 reaction zones connected in series.

A preferred further embodiment of the method is characterized in that the process gas emerging from at least one reaction zone is subsequently passed through at least one heat exchange zone downstream of said reaction zone.

In a particularly preferred further embodiment of the process, after each reaction zone is at least one, preferably exactly one heat exchange zone, through which the process gas leaving the reaction zone is passed.

The reaction zones can either be arranged in a reactor or arranged divided into several reactors. The arrangement of the reaction zones in a reactor leads to a reduction in the number of apparatuses used. The individual reaction zones and heat exchange zones can also be arranged together in a reactor or in any combination of reaction zones with heat exchange zones in several reactors.

If reaction zones and heat exchange zones are present in a reactor, then in an alternative embodiment of the invention there is a heat insulation zone between them, in order to be able to obtain the adiabatic operation of the reaction zone.

In addition, each of the series-connected reaction zones can be replaced or supplemented independently of one another by one or more reaction zones connected in parallel. The use of reaction zones connected in parallel allows in particular their replacement or supplementation during ongoing continuous operation of the process.

Parallel and successive reaction zones may in particular also be combined with one another. However, the process according to the invention particularly preferably has exclusively reaction zones connected in series.

The reactors preferably used in the process according to the invention can consist of simple containers with one or more reaction zones, as e.g. in Ullmann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, VoI B4, page 95-104, page 210-216), wherein in each case between the individual reaction zones and / or heat exchange zones heat insulation zones can be additionally provided.

In an alternative embodiment of the method, therefore, there is a heat insulation zone at least between a reaction zone and a heat exchange zone. Preferably, there is a heat isolation zone around each reaction zone.

The catalysts or the fixed beds thereof are mounted in a manner known per se on or between gas-permeable walls comprising the reaction zone of the reactor. Particularly in the case of thin fixed beds, technical devices for uniform gas distribution can be provided in the flow direction in front of the catalyst beds. These can be perforated plates or other internals that cause a uniform entry of the process gas into the fixed bed by generating a small but uniform pressure loss.

In a preferred embodiment of the method, the inlet temperature of the process gas entering the first reaction zone is from 700 to 1000 ^° C., preferably from 800 to 950 ^° C., particularly preferably from 850 to 900 ^° C. In a further preferred embodiment of the method, the absolute pressure at the inlet of the first reaction zone is between 10 and 40 bar, preferably between 20 and 35 bar, particularly preferably between 25 and 30 bar.

In yet another preferred embodiment of the method, the residence time of the process gas in all reaction zones together is between 0.05 and 20 s, preferably between 0.1 and 5 s, particularly preferably between 0.5 and 3 s.

The hydrocarbon, the carbon dioxide and the water vapor are preferably fed only before the first reaction zone. This has the advantage that the entire process gas is available for absorbing heat of reaction in all reaction zones. In addition, by such a procedure, the space-time yield can be increased, or the necessary catalyst mass can be reduced. However, it is also possible before one or more of the reaction zones following the first reaction zone to meter in hydrocarbon, carbon dioxide and / or water vapor as needed in the process gas. In addition, the temperature and the conversion can be controlled via the supply of these process gases between the reaction zones. Among other and / or alternative, the process gases may also be preheated.

In a preferred embodiment of the method according to the invention, the process gas is heated after at least one of the reaction zones used, more preferably after each reaction zone. For this purpose, the process gas is passed after exiting a reaction zone through one or more of the above-mentioned heat exchange zones, which are located behind the respective reaction zones. These may be used as heat exchange zones in the form of heat exchangers known to those skilled in the art, e.g. Tube bundle, plate, Ringnut-, spiral, finned tube, micro-heat exchanger be executed. The heat exchangers are preferably microstructured heat exchangers.

In the context of the present invention, microstructured means that the heat exchanger for the purpose of heat transfer comprises fluid-carrying channels, which are characterized in that they have a hydraulic diameter between 50 μm and 5 mm. The hydraulic diameter is calculated as four times the flow cross-sectional area of the fluid-conducting channel divided by the circumference of the channel.

In a particular embodiment of the method, the heating of the process gas takes place in the heat exchange zones by a condensation of a heat transfer medium. Within this particular embodiment, it is preferable to carry out condensation, preferably partial condensation, in the heat exchangers containing the heat exchange zones on the side of the heating medium.

Partial condensation referred to in the context of the present invention, a condensation in which a gas / liquid mixture of a substance is used as a heating medium and in which there is still a GasTFlüssigkeitsgemisch this substance after heat transfer in the heat exchanger.

The execution of a condensation is particularly advantageous because in this way the achievable heat transfer coefficient to the process gases from the heating medium is particularly high and thus efficient heating can be achieved.

Performing a partial condensation is particularly advantageous because the release of heat by the heating medium thereby no longer results in a change in temperature of the heating medium, but only the gas-liquid equilibrium is shifted. This has the consequence that over the entire heat exchange zone, the process gas is heated to a constant temperature. This in turn safely prevents the occurrence of radial temperature profiles in the flow of process gases, thereby improving control over the reaction temperatures in the reaction zones and, in particular, preventing the formation of local overheating by radial temperature profiles.

In an alternative embodiment, instead of a condensation / partial condensation, it is also possible to provide a mixing zone upstream of the inlet of a reaction zone in order to standardize the radial temperature profiles, if appropriate during the heating, in the flow of the process gases by mixing transversely to the main flow direction.

In a preferred embodiment of the process, the reaction zones connected in series are operated at an average temperature increasing or decreasing from reaction zone to reaction zone. This means that within a sequence of reaction zones, the temperature can be both increased and decreased from reaction zone to reaction zone. This can be adjusted, for example, via the control of the heat exchange zones connected between the reaction zone. Further options for setting the average temperature are described below.

The thickness of the flow-through reaction zones can be chosen to be the same or different and results according to laws generally known in the art from the residence time described above and the process gas quantities enforced in the process. The inventively enforceable with the method mass flows of product gas (Carbon monoxide), which also results in the process gas quantities to be used, are usually between 5 and 10 t / h, preferably between 7 and 8 t / h, more preferably between 7.3 and 7.4 t / h.

The maximum outlet temperature of the process gas from the first reaction zone is usually in a range from 500 ^° C. to 850 ^° C., preferably from 650 ^° C. to 800 ^° C., more preferably from 700 ^° C. to 750 ^° C. subsequent measures with regard to their inlet temperature by the skilled person, are determined freely according to the inventive method.

The control of the temperature in the reaction zones preferably takes place by at least one of the following measures: dimensioning of the adiabatic reaction zone, control of the heat supply between the reaction zones, addition of further process gas between the reaction zones, molar ratio of the educts / excess of water vapor used and / or carbon dioxide, Addition of secondary constituents, in particular nitrogen, before and / or between the reaction zones.

The composition of the catalysts in the reaction zones according to the invention may be identical or different. In a preferred embodiment, the same catalysts are used in each reaction zone. However, it is also advantageous to use different catalysts in the individual reaction zones.

Thus, especially in the first reaction zone, when the concentration of the reaction educts is still high, a less active catalyst can be used and in the further reaction zones the activity of the catalyst can be increased from reaction zone to reaction zone. The control of the catalyst activity can also be carried out by dilution with inert materials or carrier material.

0.1 kg / h to 10 kg / h, preferably 2 kg / h to 5 kg / h, more preferably 3.5 kg / h to 4.5 kg / h of carbon monoxide as a constituent per 1 kg of catalyst with the erfϊndungsgemäßen method of the synthesis gas are produced.

The erfϊndungsgemäße method is thus characterized by high space-time yields, combined with a reduction in the size of the apparatus and a simplification of the apparatus or reactors. This surprisingly high space-time yield is made possible by the interaction of the erfϊndungsgemäßen and preferred embodiments of the new method. In particular, the interaction of staggered adiabatic reaction zones with interposed heat exchange zones and the defined residence times allows accurate Control of the process and the resulting high space-time yields, as well as a reduction of the by-products formed.

The present invention will be elucidated with reference to the drawings without, however, being limited thereto.

Fig. 1 shows reactor temperature (T) and methane conversion (U) over a number of 12 reaction zones (S) with downstream heat exchange zones (according to Example 1).

The present invention will be further illustrated by the following example, without being limited thereto.

Examples

A process gas consisting of water vapor, methane and carbon dioxide is fed to the process. The molar ratio of methane to carbon dioxide is 1: 1 and the molar ratio of methane to water vapor is 1: 2. The process is in a total of 12 fixed catalyst beds of magnesium spinel coated with nickel, wherein a proportion of 15.2 wt .-% of nickel is present on the catalyst, ie in 12 reaction zones operated.

The composition of the process gas used at the beginning of the first reaction zone causes the reactions according to the formulas (II and III) to be strongly forced to the left equilibrium side. This is especially true for the reaction according to formula (H).

Each after a reaction zone is a heat exchange zone in which the exiting process gas is reheated before it enters the next reaction zone.

The absolute inlet pressure of the process gas directly in front of the first reaction zone is 29 bar. The length of the fixed catalyst beds, ie the reaction zones, is always 0.1 m. The activity of the catalyst used is not variable across the reaction zones. The proportion catalyst volume per total volume of each reaction zone is always 25 vol .-%. There is no addition of process gas before the individual reaction zones. The total residence time in the plant is 0.6 seconds.

The results are shown in FIG. Here, the individual reaction zones are listed on the x-axis, so that a spatial course of developments in the process is visible. The temperature of the process gas is indicated on the left y-axis. The temperature profile across the individual reaction zones is shown as a thick, solid line. On the right y-axis the total conversion of methane is indicated. The course of the conversion over the individual reaction zones is shown as a thick dashed line.

It can be seen that the inlet temperature of the process gas before the first reaction zone is about 900 ^° C. Due to the substantially endothermic reaction to synthesis gas under adiabatic conditions, the temperature in the first reaction zone drops to about 735 ° C before the process gas is reheated in the downstream heat exchange zone. The inlet temperature before the next reaction zone is again about 900 ⁰ C. By endothermic adiabatic reaction, it decreases again to about 780 ⁰ C. The sequence of cooling by endothermic, adiabatic reaction and heating continues with successively increasing inlet temperatures before the respective reaction zones. The inlet temperature of the process gas before the last reaction zone thus changes in the course of the process to a value of about 850 ^° C. It is obtained a conversion of methane of 74.8%. The space-time yield achieved, based on the mass of catalyst used, is 3.99 kg _Kθh _enemone χid / kgκath.

Claims

claims

1. A process for the production of synthesis gas from hydrocarbons, carbon dioxide and water vapor in an endothermic, heterogeneous catalytic gas phase reaction, characterized in that it comprises 5 to 30 successive reaction zones with adiabatic conditions.

2. The method according to claim 1, characterized in that the conversion takes place in 7 to 20, preferably 10 to 15 successive reaction zones.

3. The method according to claim 1 or claim 2, characterized in that the inlet temperature of the entering into the first reaction zone process gas 700 to 1000 ⁰ C, preferably from 800 to 950 ⁰ C, more preferably from 850 to 900 ⁰ C.

4. The method according to claim 1 to 3, characterized in that the absolute pressure at the entrance of the first reaction zone between 10 and 40 bar, preferably between 20 and 35 bar, more preferably between 25 and 30 bar.

5. The method according to any one of the preceding claims, characterized in that the residence time of the process gas in all reaction zones between 0.05 and 20 s, preferably between 0.1 and 5 s, more preferably between 0.5 and 3 s.

6. The method according to any one of the preceding claims, characterized in that the catalysts are present in a fixed bed arrangement.

7. The method according to claim 6, characterized in that the catalysts are present as monoliths.

8. The method according to claim 7, characterized in that the monolith comprises channels having a diameter of 0.1 to 3 mm, preferably from 0.2 to 2 mm, particularly preferably from 0.5 to 1.5 mm.

9. The method according to claims 1 to 6, characterized in that the catalysts are present in beds of particles having average particle sizes of 1 to 10 mm, preferably 2 to 8 mm, more preferably from 3 to 7 mm.

10. The method according to any one of the preceding claims, characterized in that after at least one reaction zone is at least one heat exchange zone through which the process gas is passed.

11. The method according to claim 10, characterized in that after each reaction zone at least one, preferably a heat exchange zone is located, through which the process gas is passed.

12. The method according to any one of the preceding claims, characterized in that there is at least one heat insulation zone between a reaction zone and a heat exchange zone.

13. The method according to claim 12, characterized in that there is a heat insulation zone around each reaction zone.