KR20050122239A - Method for the production of at least one partial oxidation and/or ammoxidation product of a hydrocarbon - Google Patents

Abstract

The invention relates to a method for the production of at least one partial oxidation and/or ammoxidation product of a hydrocarbon, wherein at least one saturated hydrocarbon (K) that is heterogeneously catalyzed is partially dehydrogenated and the resulting product gas mixture (A) containing the partially dehydrogenated hydrocarbon is used as such or in modified form for the heterogeneously catalyzed, partial oxidation and/or ammoxidation of the partially dehydrogenated hydrocarbon contained in the product gas mixture (A), wherein at least one mechanical separation operation is conducted between the product gas mixture (A) and the heterogeneously catalyzed, partial oxidation and/or ammoxidation.

Description

METHODS FOR THE PRODUCTION OF AT LEAST ONE PARTIAL OXIDATION AND / OR AMMOXIDATION PRODUCT OF A HYDROCARBON

The present invention applies one or more saturated hydrocarbons H to a heterogeneously catalyzed dehydrogenation in the gas phase to form a product gas mixture A comprising one or more partially dehydrogenated hydrocarbons H, wherein the saturated gas H Or leaving components other than partially dehydrogenated hydrocarbons H, or partially or completely removing them, to obtain product gas mixture A ', and as product component gas mixture B, product gas mixture A and / or product gas mixture A'. One or more partial oxidations of hydrocarbons and / or by applying to one or more heterogeneous catalytic partial oxidations and / or ammonia oxidations of one or more partially dehydrogenated hydrocarbons H present in product gas mixture A and / or product gas mixture A '. ) Relates to a process for producing ammonia oxidation products The.

In this specification, heterogeneous catalytic dehydrogenation refers to dehydrogenation, which proceeds endothermically and the main by-product formed is hydrogen. This is done on a solid catalyst which reduces the activation energy required for thermal cleavage of C-H bonds. Heterogeneous catalytic dehydrogenation is different from heterogeneous catalytic oxydehydrogenation, where heterogeneous catalytic oxydehydrogenation is driven by the oxygen present and water is formed as a major byproduct. In addition, heterogeneous catalytic oxydehydrogenation proceeds exothermicly.

In this regard, complete oxidation of a hydrocarbon means that all the carbon present in the hydrocarbon is converted to oxides of carbon (CO, CO ₂ ).

Reactions with oxygen deviating from the above are partial oxidation and partial ammonia oxidation in the presence of ammonia.

The processes described at the beginning are known (for example, DE-A 10219686, DE-A 10246119, DE-A 10245585, DE-A 10219685, EP-A 731077, DE-A 3313573). , DE-A 10131297, DE-A 10211275, EP-A 117146, GB-A 2118939, US-A 4532365, US-A 3161670 and EP-A 193310), in particular acrolein from propane, Acrylic acid and / or acrylonitrile are prepared and used to prepare methacrolein, methacrylic acid and / or methacrylonitrile from isobutane. Partial ammonia oxidation is inherently different from partial oxidation by the presence of ammonia in the reaction mixture. By appropriately selecting the NH ₃ and O ₂ ratios, partial oxidation and partial ammonia oxidation can be performed in parallel, i.e., simultaneously. The addition of an inert diluent gas keeps the reaction mixture of partial oxidation and / or ammonia oxidation out of the explosive range.

In this method (particularly according to the teachings of DE-A 3313573, EP-A 117146, GB-A 2118939, US-A 4532365 and US-A 3161670), the product gas mixture A as a component of gas mixture B The self and / or product gas mixture A 'is subjected to at least one heterogeneous catalytic partial oxidation and / or ammonia oxidation of one or more partially dehydrogenated hydrocarbons H present in product gas mixture A and / or product gas mixture A'. Can be applied.

In the simplest case, product gas mixture A can be converted to product gas mixture A 'by partially or completely removing any vapor present in product gas mixture A. This can be done, for example, by cooling the product gas mixture and partially or completely condensing any vapor present therein.

It will also be appreciated that other components present in product gas mixture A can be removed to yield product gas mixture A '. For example, product gas mixture A may be passed through a generally tubular membrane that is permeable only to hydrogen present in product gas mixture A to partially or completely remove hydrogen present in product gas mixture A. CO ₂ present in the product gas mixture A can be removed by passing the product gas mixture A through an aqueous alkali solution. Another method of removal is absorption / desorption according to DE-A 10245585.

However, a disadvantage of the prior art processes is that, before one or more heterogeneous catalytic partial oxidations and / or ammonia oxidations, neither product gas mixture A, product gas mixture A 'and gas mixture B are solid particles present in these gas mixtures. Is not subjected to mechanical separation processes that can be removed from these gas mixtures. If an absorbent separation process is used, in some cases the absorbent begins to saturate into solid particles over time. For example, during stripping and / or desorption they may be introduced.

In long-term indoor experiments, surprisingly, this means that very fine particles of solid catalyst used for heterogeneous catalytic dehydrogenation in the process described at the beginning can be transferred to subsequent heterogeneous catalytic partial oxidation and / or ammonia oxidation, and in part In the case it has been shown to be disadvantageous in that they settle in the catalyst bed used.

The latter is disadvantageous in that heterogeneous catalytic partial oxidation and / or ammonia oxidation is generally carried out in the presence of excess oxygen relative to the reaction stoichiometry.

In the presence of oxygen, in general, suitable catalysts for heterogeneous catalytic dehydrogenation to the complete combustion to the hydrocarbon CO ₂ and H ₂ O (for example, US-A reference 4788371 call) and H ₂ and O ₂ to generate the H ₂ O Also promotes the hydrogen-oxygen gas reaction. All of these may result in undesirable reactant consumption in heterogeneous catalytic partial oxidation and / or ammonia oxidation (which, at the same time, undesirably mean additional heat formation) in an undesirable manner, or It is disadvantageous in that it introduces harbor risk which can be difficult.

It is an object of the present invention to provide a more advantageous procedure which is particularly suitable for safe continuous processes over a relatively long period of time (the on-stream time of the catalyst for heterogeneous catalytic partial oxidation and / or ammonia oxidation is generally Years).

The inventors have found that this object is subject to the application of at least one saturated hydrocarbon H to heterogeneous catalytic dehydrogenation in the gas phase to form a product gas mixture A comprising at least one partially dehydrogenated hydrocarbon H, wherein the product gas mixture A is not saturated hydrocarbon H Leaving product components other than partially dehydrogenated hydrocarbons H, or partially or completely removing them, to obtain product gas mixture A ', product gas mixture A and / or product gas mixture A' as a component of gas mixture B One or more heterogeneous catalytic partial oxidations and / or ammonia oxidations of one or more partially dehydrogenated hydrocarbons H present in gas mixture A and / or product gas mixture A ', wherein the one or more heterogeneous catalytic partial oxidations and / or A) prior to ammonia oxidation, product gas mixture A, product gas horn One or more partial oxidation and / or ammonia oxidation products of hydrocarbons comprising subjecting water A 'and / or gas mixture B to one or more mechanical separation processes capable of removing solid particles present in these gas mixtures It was found that this is achieved by the method of preparing.

The process according to the invention, in particular for the process for heterogeneous catalytic partial oxidation and / or ammonia oxidation, uses product gas mixture A as it is (and is used) and generally condenses the vapor partially or completely therein. It is advantageous when the product gas mixture A 'obtained from the product gas mixture A is used.

Suitable gas purification apparatuses using a mechanical separation process and according to the present invention are, for example, chamber separators, collision separators and centrifuges using mass forces. However, it is also possible to use acoustic separators in the method according to the invention. Aerocyclone is preferred. In a simple manner, however, the mechanical separation process used according to the invention may be filtration. Useful filter layers include filter fabrics, porous filter compositions, paper webs, or oil-wet metal filters. Electrostatic separators can also be used in accordance with the present invention. In the simplest manner, the gas mixture can flow through an inert stationary bed in which very fine solid particles present in the gas mixture are separated before the gas mixture reaches a catalyst for heterogeneous catalytic partial oxidation and / or ammonia oxidation. have. As used herein, the term "mechanical separation process" includes a spray device in which the gas is exposed countercurrently or countercurrently to droplets (eg, high boiling organic liquids or water) capable of absorbing solid particles present in the gas. something to do. The spray liquid is exchanged after several recycles to prevent saturation by solid particles.

It will also be appreciated that different continuous mechanical separation processes can be used according to the invention.

In general, heterogeneous catalytic dehydrogenation of saturated hydrocarbons H is according to the invention (particularly in the case of propane and / or isobutane) DE-A 3313973, WO 01/96270, DE-A 10131297 or DE-A It is performed as described in 10211275.

Since the heterogeneous catalytic dehydrogenation proceeds with decreasing volume, the conversion can be increased by reducing the partial pressure of the product. This can be achieved in a simple manner, for example by dehydrogenation under reduced pressure and / or mixing in a substantially inert diluent gas (e.g. vapor) which generally constitutes an inert gas for the dehydrogenation reaction. It is. A further advantage of dilution with steam in general is that it reduces the carbonization of the catalyst used because the steam reacts with the carbon formed by the principle of coal gasification. The vapor may also be used as diluent gas downstream of one or more oxidation and / or ammonia oxidation zones. However, the vapor can also be simply partially or completely removed from the product mixture A of dehydrogenation (eg by condensation), which results in one or more partial oxidations and / or oxidation of the product mixture A 'obtainable here. When used for ammonia oxidation opens up the possibility of increasing the proportion of diluent gas N ₂ . Further diluents suitable for heterogeneous catalytic dehydrogenation are, for example, CO, methane, ethane, CO ₂ , nitrogen and inert gases such as He, Ne and Ar. All the diluents mentioned can be used alone or in the form of very different mixtures. It is advantageous if the mentioned diluents are generally also suitable diluents for one or more partial oxidations and / or ammonia oxidations. Inert gases are, in particular, diluents which behave inactively in certain reactions (ie, less than 5 mol%, preferably less than 3 mol%, even more preferably less than 1 mol% are chemically changed). In principle, catalysts useful for heterogeneous catalyst dehydrogenation are all dehydrogenation catalysts disclosed by the prior art. These are usually divided into two groups, that is, those of acid nature (eg chromium oxide and / or aluminum oxide) and generally one or more, generally relatively precious metals (eg platinum, palladium) deposited on an acid support. , Tin, gold, and silver).

In other words, according to the invention, WO 01/96270, EP-A 731077, DE-A 10211275, DE-A 10131297, WO 99/46039, US-A 4788371, EP-A 0 705 136 WO 99/29420, US-A 4220091, US-A 5430220, US-A 5877369, EP-A 0 117 146, DE-A 199 37 196, DE-A 199 37 105 and All dehydrogenation catalysts recommended in DE-A 199 37 107 can be used. In particular, any catalyst according to Examples 1, 2, 3 and 4 of DE-A 199 37 107 can be used.

These include from 10 to 99.9% by weight of zirconium dioxide, from 0 to 60% by weight of aluminum oxide, silicon dioxide and / or titanium dioxide, and at least one element of the first or second main group of the periodic table of elements, an element of the third transition metal group, eighth A dehydrogenation catalyst containing from 0.1 to 10% by weight of elements of the main group, lanthanides and / or tin, with the addition of 100% by weight.

In order to carry out heterogeneous catalytic dehydrogenation according to the invention, useful reactor types and process variables are in principle all of those disclosed in the prior art. Descriptions of these process variables are contained in all prior art documents cited, for example, with respect to dehydrogenation catalysts.

A relatively comprehensive description of suitable dehydrogenation processes according to the invention is also described in the Catalytica Studies Division, Oxidative Dehydrogenation and Alternative Dehydrogenation Process, Study Number 4192 OD, 1993, 430 Ferguson Drive, Mountain View, California, 94043-. 5272, USA.

Part of the heterogeneous catalytic dehydrogenation of saturated hydrocarbons is that it proceeds endothermically. This means that the heat (energy) required to obtain the required reaction temperature must be supplied to the starting reaction gas mixture before heterogeneous catalyst dehydrogenation and / or during dehydrogenation.

In addition, as a result of the high reaction temperatures required, the formation of small amounts of high boiling point, high molecular weight organic compounds (including carbon) and deposition on the catalyst surface to inactivate it results in heterogeneous catalytic dehydrogenation of saturated hydrocarbons such as propane and isobutane. Is typical. In order to minimize this adverse accompanying phenomenon, the reaction gas mixture which contains saturated hydrocarbon H to be dehydrogenated and which will pass at elevated temperatures over the catalyst surface for heterogeneous catalytic dehydrogenation can be diluted with steam. Carbon deposition is partially or completely removed in this way under the conditions given by the principle of coal gasification.

Another possibility to remove the deposited carbon compounds is to occasionally burn the deposited hydrocarbons by flowing oxygen-containing gas into the dehydrogenation catalyst at elevated temperatures. However, substantial inhibition of the formation of carbon deposits is also possible by adding molecular hydrogen to saturated hydrocarbons H (eg, propane or isobutane) which are dehydrogenated under heterogeneous catalysis before being transferred over the dehydrogenation catalyst at elevated temperatures.

It will also be appreciated that steam and molecular hydrogen may be added as a mixture to saturated hydrocarbons H which are dehydrogenated under heterogeneous catalysis. The addition of molecular hydrogen to heterogeneous catalytic dehydrogenation of propane also reduces the undesirable formation of allene (propadiene), propene and acetylene as by-products.

Suitable reactor types for heterogeneous catalyst dehydrogenation according to the invention are fixed bed tube or tube bundle reactors. This means placing the dehydrogenation catalyst in one reaction tube or bundle of reaction tubes as a stationary phase. The reaction tube is heated by a gas which is combusted in the space around the reaction tube, for example a hydrocarbon such as methane. It is advantageous to apply this direct form of catalyst tube heating first only to about 20-30% of the stationary phase and to heat the remaining stationary phase length to the reaction temperature required by the radiant heat released during combustion. In this way, approximately isothermal reaction control can be achieved. Suitable reaction tube inner diameters are about 10 to 15 cm. Typical dehydrogenation tube bundle reactors include 300 to 1000 reaction tubes. The temperature inside the reaction tube varies in the range of 300 to 700 ° C, preferably 400 to 700 ° C. Advantageously, the starting reaction gas mixture is fed to a tubular reactor preheated to the reaction temperature. It is possible for the product gas mixture to exit the reaction tube at a temperature of 50-100 ° C. lower. However, this outlet temperature may be higher or the same level. In the above procedure, it is suitable to use an acidic dehydrogenation catalyst based on chromium oxide and / or aluminum oxide. Frequently, no diluent gas will be used and the starting material used for the starting reaction gas will be essentially saturated hydrocarbon H (eg, propane or crude propane) alone. Dehydrogenation catalysts are also generally used without dilution.

On an industrial scale, multiple (eg three) tubular reactors can be operated simultaneously. In this case, two of these reactors, according to the invention, may optionally be in the course of dehydrogenation and the catalyst in the third reactor is regenerated without damaging one or more partial zones.

This procedure is suitable, for example, for the Linde propane dehydrogenation process disclosed in the literature. However, it is important to use such a tube bundle reactor in the present invention.

This process has also been developed by Philips Petroleum Co. (see, eg, US-A 4 902 849, US-A 4 996 387 and US-A 5 389 342). Steam active reforming (STAR) process. The dehydrogenation catalyst used in the STAR process is advantageously a platinum containing promoter on zinc (magnesium) spinel as a support (see eg US Pat. No. 5 073 662). In contrast to the BASF-lind propane dehydrogenation process, propane that is dehydrogenated is diluted with steam in the STAR process. Typical molar ratios of steam to propane are 4 to 6. The starting reactor pressure is frequently 3 to 8 atm and the reaction temperature is advantageously selected at 480 to 620 ° C. Typical liquid hourly space velocity (LHSV) of the total reaction gas mixture is from 0.5 to 10 h ⁻¹ .

Heterogeneous catalytic dehydrogenation according to the invention can also be carried out in a mobile phase. For example, the mobile catalyst bed can be housed in a radial flow reactor. In this reactor, the catalyst moves slowly from top to bottom and the reaction gas mixture flows radially. This procedure is applied, for example, in the UOP-Oleflex dehydrogenation process. Since the reactors of this process operate substantially adiabatic, it is advantageous to operate a plurality of reactors connected in series as cells (typically up to four). This prevents the temperature difference of the reaction gas mixture at the reactor inlet and reactor outlet from becoming too large (in the adiabatic mode of the process, the starting reaction gas mixture functions as a heat carrier whose lowering of the reaction temperature depends on its heat content. ), An attractive overall conversion rate can be obtained.

When the catalyst phase exits the mobile phase reactor, it is fed to the regeneration process and then reused. The dehydrogenation catalyst used in this process can be, for example, a spherical dehydrogenation catalyst consisting essentially of platinum on a spherical aluminum oxide support. In the UOP variant, hydrogen is added to the dehydrogenated saturated hydrocarbon H (eg propane) to prevent premature catalyst aging. The working pressure is typically 2 to 5 atm. The hydrogen to propane (molar) ratio is advantageously between 0.1 and 1. The reaction temperature is preferably 550-650 ° C. and the contact time of the catalyst and reaction gas mixture is selected from about 2-6 h ⁻¹ .

In the fixed bed process described, the catalyst shape may likewise be spherical and cylindrical (hollow or filled) or may have different shapes.

Proceedings De Witt, Petrochem. A further process parameter of heterogeneous catalytic dehydrogenation of saturated hydrocarbons H as described in Review, Houston, Texas, 1992a, N1] is the possibility of heterogeneous catalytic dehydrogenation in the fluidized bed without dilution of hydrocarbons.

According to the invention, for example, it is possible to operate two fluidized beds simultaneously, one of which can sometimes be regenerated without adversely affecting the overall process. The active composition used is chromium oxide on aluminum oxide. The working pressure is typically 1 to 2 atm and the dehydrogenation temperature is generally 550 to 600 ° C. Preheating the dehydrogenation catalyst to the reaction temperature introduces the heat necessary for dehydrogenation into the reaction system. The dehydrogenation process described above is known in the literature as the Namprogetti-Yarsintez process.

Unlike the aforementioned process, heterogeneous catalytic dehydrogenation can also be realized with substantial exclusion of oxygen by a method developed by ABB Lummus Crest (Proceedings De Witt, Petrochem. Review, Houston, Texas, 1992, P1).

What is common to the heterogeneous catalytic dehydrogenation process of saturated hydrocarbons H by the substantial exclusion of oxygen described so far is that they are at least 30 mol% (generally 60 mol% or more) (based on one pass of the reaction zone) (E.g. propane) conversion rate. Advantageous in the present invention is that it is sufficient to achieve a conversion rate of saturated hydrocarbon H (eg, propane) of at least 5 mol% and at most 30 mol% or at most 25 mol%. This means that the heterogeneous catalyst dehydrogenation may be operated at a conversion rate of 10 to 20 mole% (conversion rate based on one pass of the reaction zone). Among other factors, it functions as a substantially inert diluent gas in subsequent one or more partial oxidations and / or partial ammonia oxidations, and then the ratio of the residual amount recycled into the dehydrogenation zone and / or one or more partial zones with little loss later. Based on the converted saturated hydrocarbon H (eg, propane).

In order to realize the conversion mentioned above, it is advantageous to carry out heterogeneous catalyst dehydrogenation at a working pressure of 0.3 to 3 atm. It is also advantageous to dilute saturated hydrocarbons H which are dehydrogenated under heterogeneous catalysis with steam. For example, the heat capacity of water, on the one hand, offsets some of the endothermic effects of dehydrogenation, and on the other hand, dilution with steam reduces the reactant and product partial pressures, which advantageously affects the equilibrium positioning of dehydrogenation. . As already mentioned, the use of steam also has a beneficial effect on the initiation time of the noble metal-containing dehydrogenation catalyst. If necessary, molecular hydrogen may be added as an additional component. The molar ratio of molecular hydrogen to saturated hydrocarbon H (eg propane) is generally 5 or less. Thus, the molar ratio of steam to saturated hydrocarbon H at a relatively low saturated hydrocarbon H conversion can be from 0 to 30, advantageously from 0.1 to 2, advantageously from 0.5 to 1. It has also been found to be advantageous for a low propane conversion rate process in which a relatively small amount of heat is consumed in one reactor passage of the reaction gas and a relatively low reaction temperature is sufficient to achieve the conversion rate in one reactor passage.

Therefore, it may be appropriate to carry out (substantially) adiabatic dehydrogenation of relatively low propane conversion. This means that the starting reaction gas mixture will generally be first heated to 500-700 ° C. (or 550-650 ° C.) (eg by direct heating of the walls surrounding it). In general, one adiabatic passage through the catalyst bed will be sufficient to achieve the desired conversion rate and the reaction gas mixture will be cooled by about 30 ° C. to 200 ° C. (depending on the conversion rate and dilution rate). In addition, the presence of steam as a heat carrier is remarkably advantageous in terms of adiabatic methods. The lower the reaction temperature, the longer the start time of the catalyst phase used.

In principle, relatively low conversion heterogeneous catalyst dehydrogenation can be carried out in a fixed bed reactor or in a mobile bed or fluidized bed reactor, whether carried out adiabatic or isothermally.

Notably, in order to realize the process according to the invention in particular in adiabatic processes, a single shaft furnace reactor in which the reaction gas mixture flows axially and / or radially is sufficient as a fixed bed reactor.

In the simplest case this is a single closed reaction volume, eg term, vessel, having an internal diameter of 0.1 to 10 m or 0.5 to 5 m and a fixed catalyst bed applied to the support device (eg lattice). A hot, saturated hydrocarbon H-containing reaction gas is axially flowed into the reaction volume filled with the catalyst and thermally insulated in the adiabatic process. The catalyst shape may be spherical or may be annular or rod-shaped. In this case, since the reaction volume can be realized by a very inexpensive apparatus, all catalyst shapes exhibiting a particularly low pressure drop are preferred. These are especially in the form of catalysts which give rise to large void volumes or which, for example, constitute a monolithic or honeycomb configuration. In order to realize the radial flow of the saturated hydrocarbon H-containing reaction gas, for example, the reactor may consist of two concentric cylindrical lattice arranged in a shell and the catalyst phase may be arranged in an annular gap. In the thermal insulation case, the metal shell will be thermally insulated.

Catalysts useful for heterogeneous catalyst dehydrogenation, which exhibit a relatively low conversion rate in one pass, are especially the catalysts disclosed in DE-A 199 37 107, especially those disclosed by way of example.

After a prolonged operating time, the catalysts described above can, for example, initially contain 300-600 ° C., often 400-600, of air (preferably) diluted with nitrogen and / or vapor in the first regeneration step. It can simply be regenerated by passing over the catalyst at an inlet temperature of 550 ° C. The gas time space velocity of the regeneration gas may be, for example, 50 to 10000 h ^−1, and the oxygen content of the regeneration gas may be 0.5 to 20 volume%.

In a further subsequent regeneration step, the regeneration gas in which other conditions are used under the same regeneration conditions may be air. In view of the application, it is advantageous to wash the catalyst with an inert gas (eg N ₂ ) before regeneration.

Then, it is generally recommended that the other conditions be regenerated with pure molecular hydrogen or molecular hydrogen diluted with an inert gas (preferably vapor) under the same conditions (the hydrogen content must be at least 1% by volume).

Heterogeneous catalytic dehydrogenation, which exhibits a relatively low conversion rate (below 30 mol%), is, in all cases, at the same hourly space velocity (both the total reaction gas and the saturated hydrocarbon H contained therein) as a variant of high conversion (greater than 30 mol%). ) May be performed. This gas time space velocity of the reaction gas can be, for example, between 100 and 10000 h ⁻¹ , frequently between 300 and 5000 h ⁻¹ , ie in many cases about 500 to 3000 h ⁻¹ .

In particular in small scale, relatively low conversion heterogeneous catalyst dehydrogenation can be realized in a tray reactor.

It comprises more than one catalyst phase in series in the space, promoting dehydrogenation. The catalyst phase number may be 1-20, advantageously 2-8, or 3-6. The catalyst phases are preferably arranged continuously radially or axially. In view of the application, it is suitable to use fixed bed catalyst types in this tray reactor.

In the simplest case, the fixed catalyst phase in the axial furnace reactor is arranged axially or in the annular gap of the concentric cylindrical lattice. However, it is also possible to arrange the annular gaps in pieces on top of each other, and after the gas passes through one fragment, radially transfer the gas to the next fragment above or below the fragment.

Suitably, the reaction gas mixture is passed over a heat exchanger rib heated by, for example, hot gas, or from a pipe heated by hot combustion gas, while going from one catalyst bed to the next. By doing so, the reaction gas mixture will be intermediately heated in the tray reactor.

When the tray reactor is operated adiabatic, in particular when using the catalysts described in DE-A 199 37 107, in particular the catalysts of typical embodiments, the dehydrogenation reactor preheated to a temperature of 450 to 550 ° C. It is sufficient for the desired conversion (30 mol% or less) to be delivered at and maintained within this temperature range in the tray reactor. This means that total dehydrogenation can be realized at very low temperatures, which is particularly advantageous for the start time of the catalyst stationary bed between two regenerations.

It is even more advantageous to carry out the above-described intermediate heating directly (self heating method). To this end, a limited amount of molecular oxygen is added to the reaction gas mixture before the reaction gas mixture passes through the first catalyst bed and / or between subsequent catalyst beds. Depending on the dehydrogenation catalyst used, the hydrocarbons contained in the reaction gas mixture, any coke or coke compounds already deposited on the catalyst surface and / or are formed during heterogeneous catalytic dehydrogenation and / or added to the reaction gas mixture Limited combustion of hydrogen occurs (this also means catalysts which in particular (optionally) promote the combustion of hydrogen (and / or hydrocarbons) in view of application (examples of useful catalysts are described in US Pat. Nos., US-A 5430209, US-A 5530171, US-A 5527979, and US-A 5563314; for example, such catalyst phases may alternately be accommodated in a tray reactor with a phase containing a dehydrogenation catalyst. May be advantageous for introducing the catalyst phase into a tray reactor filled with The heat of reaction released allows the heterogeneous catalyzed propane dehydrogenation to operate substantially isothermally in a substantially autothermal manner. As the selected residence time of the reaction gas in the catalyst bed increases, dehydrogenation is possible, or at a substantially constant temperature, which in particular extends the onset time between the two regenerations.

In general, the oxygen supply described above should be carried out such that the oxygen content of the reaction gas mixture is from 0.5 to 30% by volume, based on the amount of saturated hydrocarbon H contained in the reaction gas mixture. Useful oxygen sources include pure molecular oxygen and inert gases such as CO, CO ₂ , N ₂ or inert gases, especially oxygen diluted with air. The resulting combustion gases generally exhibit an additional dilution effect and thus support heterogeneous catalytic dehydrogenation.

The isothermality of heterogeneous catalytic dehydrogenation can be further improved by introducing (but not necessarily) advantageously evacuated closed (eg tubular) guts before filling the space between the catalyst beds in the tray reactor. This gut may be placed on each catalyst. This gut contains suitable solids or liquids that evaporate or melt above any temperature to consume heat and condense again to release heat when the temperature drops below this value.

Another possible way of heating the starting reaction gas mixture for heterogeneous catalytic dehydrogenation to the required reaction temperature is to burn some of the saturated hydrocarbons H and / or H ₂ contained therein by molecular oxygen and (e.g. On certain combustion catalysts, for example by simply passing over and / or therein), heating to the desired reaction temperature by the heat of combustion released in this way. The resulting combustion products (eg CO ₂ and H ₂ O), and any N ₂ accompanying the molecular oxygen required for combustion, advantageously constitute an inert diluent gas.

The hydrogen combustion described above can be realized in particular on a small scale as described in DE-A 10211275. This continuously feeds the reaction gas containing the saturated hydrocarbon H to be dehydrogenated into the reaction zone, delivers the reaction gas in the reaction zone over one or more fixed catalyst beds, over which molecular hydrogen and at least in part by catalysis dehydrogenation Dehydrogenated hydrocarbon H is formed), at least one molecular oxygen-containing gas is added to the reaction gas before and / or after entering the reaction zone, and the molecular hydrogen contained in the reaction gas within the reaction zone is replaced by molecular oxygen. Partial oxidation to provide vapor and recover product gas comprising molecular hydrogen, steam, partially dehydrogenated hydrocarbon H and saturated hydrocarbon (dehydrogenated) from the reaction zone, the product gas being removed from the reaction zone being Divide into two parts of the composition and return one of the two parts to the dehydrogenation reaction zone as circulating gas. A process for continuous partial dehydrogenation of saturated hydrocarbons H in the gaseous phase under heterogeneous catalysis, which involves circulating and also using another portion as the product gas mixture A according to the invention.

The aforementioned variant of catalysis dehydrogenation used according to the invention can be used, in particular when the saturated hydrocarbon H to be dehydrogenated is propane and / or isobutane.

The product gas mixture A and / or the product gas mixture A 'obtained therefrom as described above can then be used in a manner known per se to effect heterogeneous catalytic partial oxidation and / or ammonia oxidation by the gas mixture B. (See prior art cited at the beginning). Prior to one or more heterogeneous catalytic partial oxidation and / or ammonia oxidation, product gas mixture A, product gas mixture A 'and / or mixture B are subjected to mechanical separation processes (removing solid particles present in these gas mixtures from these gas mixtures). Only) is essential to the present invention.

Partial oxidation useful in the process according to the invention is obtained, in particular, by partial oxidation of propene (obtained for partial dehydrogenation of propane) to acrolein and / or acrylic acid, and isobutene (partial dehydrogenation of isobutane). ) Is a partial oxidation of methacrolein and / or methacrylic acid.

Partial ammonia oxidation useful in the process according to the invention is in particular partial ammonia oxidation of propene to acrylonitrile and partial ammonia oxidation of isobutene to methacrylonitrile.

1. Dehydrogenation Catalyst

In a manner similar to that described in DE-A 10219879, a Pt / Sn alloy promoted with the elements Cs, K and La in oxide form was extruded into a ZrO ₂ .SiO ₂ mixed oxide support (3-8 mm long and 2 mm in diameter). Water; prepared according to example 3 of DE-A 10219879). Elemental stoichiometry (mass ratio) was as follows: Pt _0.3 Sn _0.6 La _3.0 Cs _0.5 K _0.2 (ZrO ₂ ) _88.3 (SiO ₂ ) _7.1 .

The alloy promoted as described above was applied by saturating the cleaved support with a salt solution of a suitable metal, followed by heat treatment (1.5 hours) at 560 ° C. in an air stream. During the heat treatment, the active components Pt and Sn, and the promoters were converted into their oxide form. In the dehydrogenation reactor described below, the active component of the catalyst precursor was reduced to a metal as described below in a hydrogen stream at 500 ° C. to obtain an active catalyst.

2. Dehydrogenation reactor (C330)

Using 650 ml (762 g) of the catalyst precursor obtained as described above, a vertical tubular reactor (tube length: 2049 mm, wall thickness: 7 mm, inner diameter: 41 mm, material: alonized inside) (i.e. Aluminum oxide-coated) steel tubes) were charged (supported from bottom to top on the catalyst substrate).

299 mm ring of steatite (outer diameter × length × inner diameter = 7 mm × 3 mm × 4 mm), followed by a 50 mm steatite sphere with a diameter of 4 to 5 mm, followed by a 25 mm steatite with a diameter of 1.5 to 2.5 mm A sphere, followed by a 500 mm catalyst precursor, followed by a 50 mm steatite sphere having a diameter of 4 to 5 mm, followed by a 1080 mm ring of steatite (outer diameter × length × inner diameter = 7 mm × 3 mm × 4 mm). The steatite used was Steatite C-220 from Ceram Tec. The reaction tubes were heated in four immediate continuous heating zones (each extending 2 × 200 cm in length) from the bottom to the top by means of a HTM Reetz heating coil furnace (reaction tube introduced). There was an air gap of about 7.25 mm wide between the tube surface and the heating coil. From bottom to top, the outer wall of the reaction tube was coated with a microporous insulating material of MPS-super G type (Mientherm, Delaware, USA) with a length of 1000 mm and a layer thickness of 100 mm (semi-insulating reaction part). . In this area, the heating only served as support heating, but in the upper area it served to directly heat the tube part.

From the top to the bottom, a thermowell was introduced with a length of 1.50 m (outer diameter = 6 mm, inner diameter = 4 mm) to be positioned in the center of the reaction tube to accommodate a number of thermal elements.

From bottom to top, a 60 cm long protective tube was introduced into the reaction tube in a corresponding manner.

The catalyst precursor was activated in a hydrogen stream as in Example 1 of DE-A 10211275.

3. Experiment apparatus

Propane or n-butane (hereinafter referred to as liquefied petroleum gas = LPG) was used as the hydrocarbon dehydrogenated under heterogeneous catalysis.

The experimental apparatus consisted of four parts: a weighing unit; Reactor unit; A circulating gas unit; And release unit.

In the metering unit, liquid starting material LPG and water (if necessary in a stream of nitrogen) were passed through an evaporator W100 and converted therein into a gas phase. Downstream of the evaporator, if necessary, air, molecular oxygen and molecular hydrogen were mixed with the gas mixture produced in the evaporator to obtain the desired fresh starting mixture.

In the reactor unit, the fresh starting mixture (optionally with a circulating gas) was heated to 400-500 ° C. in a preheater (heat exchanger) W200 and then passed to the dehydrogenation reactor C330 (from top to bottom). The heating zones exhibited temperatures of 500 ° C., 550 ° C., 550 ° C. and 550 ° C. from top to bottom. The inlet pressure into the selected dehydrogenation reactor was 2 bar.

The product gas mixture A leaving the dehydrogenation reactor was cooled to 300 ° C. by passing through an air cooled heat exchanger. Thereafter, it was further processed in or through the induction tube completely or partially through the discharge part. If only part of the cooled product gas mixture A was delivered to the outlet portion, the remaining portion was fed to the circulating gas unit.

In the circulating gas unit, a portion of the recycled product gas mixture A is initially cooled to 150 ° C. in a heat carrier oil operated heat carrier oil, then recompressed by two bars in a compressor V440, and then passed through a heat exchanger W460, When heated to 300 ° C., it was mixed with the fresh starting gas mixture upstream of preheater W200 (then fed to the preheater).

4. Results

The experimental apparatus was kept running substantially over three months. A total of 26 reaction cycles were performed. After each reaction cycle, regeneration of the dehydrogenation catalyst was carried out in the order of purging, burning, purging and reducing according to DE-A 10028582. Within the reaction cycle (the shortest duration of the reaction cycle was 3 hours and the longest duration of the reaction cycle was 100 hours), the conditions were kept constant. All cycle conditions were within the following configurations:

Hydrocarbon: LPG;

Cycle duration: 3-100 hours;

LPG amount: 500-2000 g / h;

Fresh steam: 500-1000 g / h;

Nitrogen: 0-50 L (STP) / h;

Air: 0-375 L (STP) / h;

Hydrogen: 0-150 L (STP) / h;

Temperature of heating zone: 500 ° C., 550 ° C., 550 ° C. and 550 ° C .;

Inlet pressure: 2 bar

Circulating gas ratio: 0 or 5 (ratio of amount recycled to amount released).

Parts of the test apparatus exposed to temperatures below 300 ° C. were made of V2A steel. Parts of the test apparatus exposed to temperatures above 300 ° C. were made of 1.4841 steel.

Upon completion of the experimental series, a significant amount of flash rust was detected in all parts of the experimental apparatus flowed by the circulating gas and product gas mixture A. Its original cause was the corrosion of the guide tube where the water was condensed (obviously, caused by the prior cooling of the product gas mixture A).

Analysis of the property rust by atomic absorption spectroscopy revealed the amount of Zr, Si and Pt that could be attributed only to the dehydrogenation catalyst, in addition to the expected components Fe, Ni and Cr. In total, the analyzed rust samples showed the following element amounts:

Carbon: 1.5 g / 100 g;

Al: 2.2 g / 100 g;

Cr: 2.6 g / 100 g;

Fe: 13.1 g / 100 g;

Mg: 10.6 g / 100 g;

Ni: 1.9 g / 100 g;

Pt: 0.019 g / 100 g;

Si: 20.6 g / 100 g;

Zr: 0.17 g / 100 g.

In the analysis of randomly sampled product gas mixture A samples, Zr, Si and Pt were not detected. Clearly, the amount present is below the detection limit.

Claims (5)

One or more saturated hydrocarbons H is subjected to heterogeneously catalyzed dehydrogenation in the gas phase to form a product gas mixture A comprising one or more partially dehydrogenated hydrocarbons H, and the product gas mixture A is partially saturated and not saturated hydrocarbon H. Leaving product components other than the digested hydrocarbons H, or partially or completely removing them, to obtain product gas mixture A ', product gas mixture A and / or product gas mixture A' as a component of gas mixture B; Applies to one or more heterogeneous catalytic partial oxidations and / or ammonia oxidations of one or more partially dehydrogenated hydrocarbons H present in A and / or product gas mixture A ', wherein the one or more heterogeneous catalytic partial oxidations and / or Prior to ammonia oxidation, product gas mixture A, product gas mixture A And / or applying gas mixture B to one or more mechanical separation processes capable of removing solid particles present in these gas mixtures to produce one or more partial oxidation and / or ammonia oxidation products of hydrocarbons. The process of claim 1 wherein the saturated hydrocarbon H is propane and the heterogeneous catalytic partial oxidation of the partially dehydrogenated hydrocarbon H is partial oxidation of propene to acrolein and / or acrylic acid. The process of claim 1 wherein the saturated hydrocarbon H is isobutane and the heterogeneous catalytic partial oxidation of the partially dehydrogenated hydrocarbon H is partial oxidation of isobutene to methacrolein and / or methacrylic acid. The process of claim 1 wherein the saturated hydrocarbon H is propane and the heterogeneously catalyzed partial ammonia oxidation of the partially dehydrogenated hydrocarbon H is a partial ammonia oxidation of propene to acrylonitrile. 2. The process of claim 1 wherein the saturated hydrocarbon H is isobutane and the heterogeneously catalyzed partial ammonia oxidation of the partially dehydrogenated hydrocarbon H is partial ammonia oxidation of isobutene to methacrylonitrile.