RU2032648C1 - Process for preparing 1,3-butadiene and catalyst for preparation thereof - Google Patents

Abstract

FIELD: petrochemical industry. SUBSTANCE: the product is 1,3-butadiene and a catalyst for the preparation thereof with reactive oxygen. The catalyst comprises (wt %): 8.4-24.30 molybdenum oxide; 0.16-1.43 alkali metal oxide; and a carrier, the balance. The carrier consists of MgO, Al₂O₃ and/or alumomagnesium. The MgO/Al₂O₃ weight ratio is 0.38-0.79 and the surface area is 25-184,3 m.sg/g. The catalyst further comprises 0.58-0.85 wt % 25-184,3 m²/g, butane oxidative dehydrogenation conditions: 550-650 C; 7-210 kPa and other conditions that provide for at least 40 wt % butadiene selectivity. EFFECT: improved process and improved properties of the catalyst. 5 cl, 10 tbl

Description

 The invention relates to the oxidation of aliphatic hydrocarbons, such as monoolefins and alkanes, in order to obtain more highly unsaturated aliphatic hydrocarbons.

 Unsaturated aliphatic hydrocarbons, such as monoolefins and diolefins, are used as monomers or comonomers in the preparation of polyolefin-based plastics.

A known method of oxidative dehydrogenation of paraffin hydrocarbons, such as butane, in the presence of a catalyst containing hydroxy compounds of molybdenum and magnesium, and a carrier based on alumina, and may also contain vanadium and / or silicon. The process is carried out at 500-650 ^C, a pressure of 7-210 kPa, and the other conditions for a reaction mixture containing 1,3-butadiene with a selectivity of at least 40 M If butane is reacted with the catalyst, the products include butenes and butadienes. However, the selectivity and volumetric yield of butadiene remains below the desired level. In addition, the initial content of hydrocarbon and oxygen is undesirable for safety reasons. Finally, a magnesium oxide-based carrier does not have sufficient toughness and abrasion resistance, which are necessary for use in fluidized bed reactors or moving bed reactors.

In one embodiment, the invention relates to a method for producing unsaturated aliphatic hydrocarbons, which comprises reacting an aliphatic hydrocarbon having at least three carbon atoms with a catalyst of the invention, which will be described below. Under the reaction conditions of the proposed method, more unsaturated aliphatic hydrocarbons, such as diolefins, are formed with a selectivity of at least about 40 M%
Advantageously, aliphatic hydrocarbons can be directly oxidized by the method of the invention in order to obtain more highly unsaturated aliphatic hydrocarbons. The proposed method showed high selectivity and high productivity of more highly unsaturated aliphatic hydrocarbons, in particular diolefins. The proposed method also showed low selectivity and low yield of undesirable deep oxidation products, such as carbon monoxide and carbon dioxide. In addition, using the proposed method, butadiene can be obtained directly from butane with a high degree of selectivity and productivity and at the same time with a low degree of selectivity for products of deep oxidation. For the purposes of the invention, “productivity” is defined as grams of the desired product obtained per gram of catalyst in 1 hour.

 Unsaturated aliphatic hydrocarbons, such as monoolefins or diolefins, are used as monomers or comonomers in the preparation of polyolefins. Butadiene can also be used as an intermediate in the preparation of styrene.

 In another embodiment, the invention relates to a solid heterogeneous catalytic composition containing an active oxygen atom, and the specified catalyst can be used in the specified method for producing unsaturated aliphatic hydrocarbons. This catalyst consists mainly of magnesium oxide, molybdenum oxide and an activator containing an alkali metal in a concentration of about 0.1-5 wt. by weight of the combined oxides of magnesium and molybdenum. The catalyst may contain vanadium oxide.

In yet another embodiment, the invention relates to a catalytic composition of molybdate containing reactive oxygen. The composition comprises a carrier comprising magnesium oxide and at least one aluminum oxide selected from the group consisting of Al ₂ O ₃ and magnesium aluminate (MgAl ₂ O ₄ ). The specified carrier has a mass ratio of MgO / Al ₂ O ₃ in the range of 0.30-4.0 and a surface area of at least about 25 m ² / year This catalyst composition also contains a catalyst component consisting mainly of molybdenum oxide, magnesium oxide and an activating amount of an alkali metal promoter. The catalyst may contain vanadium oxide. In a preferred embodiment of the invention, the index of abrasion resistance, defined below, is less than 5 wt. for 1 hour

 The catalyst composition of the invention is used in the process for oxidizing aliphatic hydrocarbons to more unsaturated aliphatic hydrocarbons. In a preferred embodiment, the catalyst composition has an increased abrasion resistance compared to the prototype catalysts. Thus, the catalyst of the invention can be used in fluidized bed reactors and moving bed reactors, such as an upstream reactor.

 Aliphatic hydrocarbons that can be used in the proposed method include alkanes and olefins that have three or more carbon atoms.

 Alternatively, alkanes can be characterized as paraffin hydrocarbons. These compounds are known to those skilled in the art as saturated hydrocarbons. As indicated, alkanes contain at least three carbon atoms and may have a straight or branched chain structure. Usually alkanes contain up to 20 carbon atoms. Examples of suitable alkanes include n-butane, n-pentane, n-hexane, n-heptane, n-octane, n-nonane, n-decane, n-dodecane and more saturated homologs, as well as isobutane, isopentane, neopentane and, in addition, branched hexanes, heptanes, octanes, nonanes, decans, dodecans, and more branched homologs. Some alicyclic hydrocarbons are suitable reagents and therefore are also used in the proposed method. Examples of aliphatic hydrocarbons include cyclobutane, cyclopentane, cyclohexane, cycloheptane, cyclooctane, methylcyclopentane, methylcyclohexane, and other alkyl substituted cycloalkanes. Preferably, the alkane is normal or linear.

 Olefins, in addition, can be characterized as aliphatic hydrocarbons containing at least one unsaturated double bond. As indicated above, olefins also contain at least three carbon atoms, and usually up to 20 carbon atoms. The location of the double bond is not critical; therefore, it may be at the end or inside the carbon chain. Preferably, however, the olefin has a normal or linear rather than branched structure. For example, 1-butene is preferred over isobutylene. Examples of suitable olefins include: 1-butene, 2-butene, 1-pentene, 2-pentene, 3-pentene, 1-hexene, 2-hexene, 3-hexene and, in addition, 1-heptene, 1-octene, 1-ionic, 1-decene and their isomers, where unsaturation occurs at any other position of the carbon chain.

 Olefins containing more than one double bond, such as 1,3-hexadiene and isoprene, which are converted to more unsaturated hydrocarbons, can also be used in the proposed method. Some aliphatic olefins, such as cyclohexene and vinylcyclohexane, are also suitable starting materials and therefore can also be included in the invention. A preferred olefin is a monoolefin. More preferred is 1- or 2-butene. Alkines, however, are not suitable reagents in the proposed method.

 The above examples of aliphatic hydrocarbons represent the most typical of hydrocarbons suitable for the proposed method, however, these examples should not limit other possible options. Other aliphatic hydrocarbons known to those skilled in the art may be suitable for the proposed process.

Preferred alkanes are normal paraffins, which can be represented by the following general formula:
CH ₃ - (CH ₂ ) _n —CH ₃ , where n is an integer from 1 to 8. More preferably, n is an integer from 2 to 6. Most preferably, if n is 2 and the alkane is n-butane.

Optionally, the aliphatic hydrocarbon reagent may be diluted with a non-reactive gas such as nitrogen, helium, argon, methane, carbon dioxide, or steam. Since the type of diluent is chosen mainly for economic reasons, nitrogen can be considered the preferred diluent. The amount of diluent (if used) can vary widely, depending on the structure of the reactor and the reactivity of the solid oxidizing agent. The hydrocarbon content of the hydrocarbon-diluent mixture is usually 1-100 M, the preferred hydrocarbon content of the mixture is 10-100 M%, and the most preferred 40-100 M
The catalyst composition of the invention is a solid heterogeneous oxide, at least a portion of the oxygen of which is active. This means that the active oxygen form is present in the catalyst and that this active oxygen form is capable of oxidizing an aliphatic hydrocarbon. Thus, in one embodiment of the invention, the catalyst is a solid oxidizing agent. After active oxygen is removed from the reactor, the catalyst is depleted. In addition, the catalyst may still form carbon residues on its surface for some time. The depleted and poisoned catalyst can be reduced using a source of gaseous oxygen. Thus, in order to carry out the catalytic process of the invention, oxygen must be added to the aliphatic hydrocarbon.

Oxygen is usually supplied from a gas source equipped with a continuous supply of oxygen. In this method, any oxygen source, such as pure elemental oxygen, air or nitrous oxide, can be used. The preferred source of oxygen is air (gas). Optionally, gaseous elemental oxygen may be diluted with a non-reactive gas such as nitrogen, helium, argon or carbon dioxide. A preferred diluent is nitrogen. If a non-reactive diluent is used, it is preferred that the oxygen content of the mixture does not exceed about 50 M%. More preferably, the oxygen content of the mixture is in the range of 0.5-30 M%. The most preferred oxygen content of the mixture is 1-20 M%.
The amount of oxygen used in the proposed catalytic method should be sufficient to completely oxidize the solid heterogeneous catalyst and sufficient to remove carbon residues from the surface of the catalyst. Preferably, if the regeneration of the catalyst is carried out separately from the oxidation of an aliphatic hydrocarbon.

 Alternatively, it is possible to supply an aliphatic hydrocarbon together with a small amount of gaseous elemental oxygen. The purpose of such a feed is to remove carbon residues from the surface of the catalyst, restore to some extent the reactive oxygen of the catalyst, and remove any hydrogen generated during the reaction. The oxygen concentration in the aliphatic hydrocarbon and when supplied is dictated only by the explosiveness considerations of this mixture. Preferably, if the oxygen concentration is below the knock limit.

The solid heterogeneous catalyst used in the invention mainly contains magnesium oxide, molybdenum oxide and an alkali metal promoter. Any source can be used as magnesium oxide, however, MgO is preferred. Similarly, any suitable source can be used as molybdenum oxide, for example MoO ₃ , (NH ₄ ) ₆ Mo ₇ O ₂₄ ˙ 4H ₂ O and (NH ₄ ) ₂ ˙ MoO ₄ . Molybdenum oxide can also be obtained from the preceding molybdenum compound, such as molybdenum carbonyl, for example, Mo (CO) ₆ . The preferred source of the molybdenum mixture is (NH ₄ ) ₆ Mo ₇ O ₂₄ ˙ 4H ₂ O. The alkaline metal promoter is a metal compound (Group IA metal) having a basicity sufficient to form more unsaturated compounds in the proposed method. Small amounts of other elements may also be present in the catalyst, provided that these elements do not substantially alter the performance of the catalyst.

The manufacturing technique of a catalyst containing mixed magnesium oxides and molybdenum, as well as an alkali metal promoter, is very simple. Typically, the desired amount of molybdenum oxide or its precursor, such as ammonium heptamolybdate or molybdenum carbonyl, is dissolved in a solvent to form a solution. A preferred molybdenum compound is ammonium heptamolybdate, and a preferred solvent is water. The resulting solution was poured over magnesium oxide to form a slurry, which was then dried to remove the solvent. If the solution is aqueous, the drying is carried out in an oven at a temperature of about 70-120 ^C. The dried composition is calcined to produce a catalytically active mixture based on magnesium oxide and molybdenum oxide. Calcination is usually conducted at 550-650 ^{° C} for 1-24 hr. Preferably, the calcination is carried out at about 600 ^{° C} and at least for 2 hours. Alternatively drained composition obtained as described above can be used directly in the catalytic the method of the invention, i.e. without prior calcination. Since the molybdenum compound can be converted into molybdenum oxide at ³⁰⁰ ° C or so, and the catalytic bed was heated to a temperature higher than ³⁰⁰ ° C, then drained composition will be converted in situ into a catalytically active mixture of oxides of magnesium and molybdenum.

A mixed oxide catalyst composition may exhibit amorphous X-ray diffraction, or may exhibit diffraction peaks characteristic of magnesium molybdate. Elemental analysis of the calcined solid composition gives the following composition, wt. MoO ₃ 6-50; MgO 94-50. The preferred composition is the following composition, wt. MoO ₃ 1-30; MgO 90-70, and most preferred is the composition, wt. MoO ₃ 15-25; MgO 85-75.

In a preferred embodiment of the invention, the molybdenum catalyst composition comprises a carrier component and a catalyst component. The carrier comprises magnesium oxide and at least one alumina selected from the group consisting of alumina (Al ₂ O ₃ ) and alumina-magnesia spinel (MgAl ₂ O ₄ ). The carrier has a weight MgO / Al ₂ O ₃ in the range of 0.30-4.0 and a surface area of at least about 25 m ² / g. The catalytic component mainly consists of molybdenum oxide, magnesium oxide and an activating amount of an alkali metal promoter. Optionally, the catalyst may comprise vanadium oxide.

Alumina imparts to the catalyst particles mainly hardness and abrasion resistance so that they can be used in fluidized bed reactors and with a moving bed. For this purpose, any source of alumina can be used, for example, α-β and γ forms of alumina, such as Boehm alumina, aqueous colloidal alumina, stoichiometric Al (OH) ₃ and aluminum alkoxides, as will be shown below. Magnesium aluminate and magnesium aluminate hydroxide are also suitable sources of alumina. In the proposed method, magnesium oxide plays a dual role: firstly, it acts as a carrier for the catalytic component, which neutralizes the acidity of aluminum oxide and other residual acid sites. It is highly desirable that the catalyst be the main one, since basicity enhances the desorption of olefin products. Any suitable source may be used as magnesium oxide, but MgO is preferred. Molybdenum oxide makes a significant contribution to the activity of the catalyst, especially if it is used in combination with magnesium oxide in the form of magnesium molybdate. Preferably, the molybdenum is in oxidation state +6. As the molybdenum oxide, any suitable source can be used, for example MoO ₃ , (NH ₄ ) ₆ Mo ₇ O ₂₄ ˙ 4H ₂ O and (NH ₄ ) ₂ ˙ MoO ₄ . Molybdenum oxide can also be obtained from a molybdenum precursor, such as molybdenum carbonyl, for example Mo (CO) ₆ . A preferred source of molybdenum oxide is ammonium heptamolybdate represented by the formula (NH ₄ ) ₆ Mo ₇ O ₂₄ ˙ 4H ₂ O. The alkali metal promoter is designed to increase the basicity of the catalyst and thereby increase the selectivity of the more highly unsaturated compounds of the invention. A small amount of other elements may be present in the catalyst, provided that these elements do not significantly affect the performance of the catalyst.

As noted, the carrier component may comprise a spinel phase (MgAl ₂ O ₄ ). The content of spinel in the carrier can range from 0-100 wt.

Typically, catalyst fabrication begins with a mixture of magnesium oxide and alumina components in order to form a support for other catalytic components. Any method can be used to combine these components; however, three methods are preferred. The first method consists in impregnating a preformed spinel preform (MgAl ₂ O ₄ ) with a solution containing a soluble magnesium salt such as magnesium nitrate, magnesium chloride, magnesium sulfate, magnesium acetate or the like, provided that the salt can be converted to magnesium oxide during calcination; and then the impregnated spinel is calcined. The mass ratio of magnesium oxide to magnesium aluminate can be expressed as the mass ratio of magnesium oxide to alumina. This ratio is critical to the performance of the catalyst composition and is discussed in detail below. Calcite nirovaniya temperature usually ranges between ^about 400-1200 C, preferably 450-900 ^C, and most preferably in the range of 500-700 ^C. The calcination is conducted for a period of time sufficient to form a fused composition, which plays a role carrier for catalytic components, and at least about 0.5 hours

The second method consists in impregnating a preformed alumina preform with a solution containing a soluble magnesium salt. Typical examples of pre-prepared alumina are anhydrous or hydrated solid alumina, such as α-, β- and γ-alumina and Böhm alumina. The mass ratio of magnesium oxide to alumina (MgO / Al ₂ O ₃ ) is a critical parameter and is described in more detail below. The calcination temperature usually varies between ^about 400-1200 C, preferably 450-900 ^C, and most preferably in the range of 500-700 ^C. The calcination is conducted for a period of time sufficient to form a fused and hardened composite which plays the role of the carrier for catalytic components. Typically, calcination is carried out for at least about 0.5 hours. During calcination, a portion of the alumina and magnesium oxide can be chemically mixed to form the spinel phase (MgAl ₂ O ₄ ), which is a uniform mixture between the alumina and magnesium oxide domains.

A third method for producing a carrier is that colloidal alumina is added to magnesium oxide and the resulting mixture is dried under conditions sufficient to obtain a carrier of magnesium oxide and alumina. Colloidal alumina is an acidified aqueous suspension of hydrated alumina, where the surface area of the particles is so large that their particles do not precipitate by gravity. The amount of said colloidal suspension of alumina added to the magnesium oxide must be such that the final weight ratio of magnesium oxide to alumina is within certain limits, which will be determined below. The pH of the mixture of colloidal alumina and magnesium oxide is about 9. Then, the mixture is dried using one of the standard methods, for example, holding and evaporation, spray drying, drying in a stream of hot air, drum drying, etc. One of the preferred methods is keeping and evaporating the mixture over a hot stove or equivalent heating agent in order to obtain a thicker gel and, as a result, a solid mass, which is then ground and sieved to obtain particles of the desired size. The aging and evaporation temperature may be any temperature compatible with the solvent system. Since the preferred solvent is water, said temperature is in the range 30-100 ° ^C. A preferred temperature is the temperature of 50-90 ^{° C,} and most preferred ^about 60-80 C. The time required for curing depends upon the amount of the gel and it has be sufficient to obtain a solid refractory mass.

 In industrial production conditions, the above-described mixture of magnesium oxide and colloidal alumina is preferably spray dried. For this purpose, any spray drying device that is commonly used in fluidized bed reactors can be used. An example is the Niro Atomizer S-12, 5-R / N atomizer. The specified device has a means for controlling the temperature at the inlet and outlet. Typically, the milled particles obtained by spray drying are spherical with a diameter of about 10-250 microns and good flow rates.

The powder obtained after drying by spraying is calcined in order to obtain a composite carrier consisting mainly of magnesium and aluminum oxides, optionally of the spinel phase of magnesium aluminate. Calcination is carried out under conditions sufficient to fuse the alumina and magnesium oxide into the cured mass. Generally, calcining is conducted at ^about 400-1200 C. More preferably carried calcining at 350-900 C and ^the most preferred temperature is 500-700 ° ^C. Usually the period of calcination depends on the amount of calcined material and is at least about 0.5 hr .

The carrier component of the invention may have any mass ratio of magnesium oxide to aluminum oxide, provided that the resulting carrier has sufficient hardness and basicity. It should be noted that although the spinel phase is present as a separate MgAl ₂ O ₄ composition, the mass ratio of MgO / Al ₂ O ₃ can be calculated. Basically, the mass ratio of MgO / Al ₂ O ₃ is in the range of 0.1-9.0; with a mass ratio of 0.3-4.0 being preferred, and 0.3-2.0 being more preferred. The most preferred ratio is 0.38-0.80. If the mass ratio is below the indicated lower preferred limit, then this means that the content of magnesium oxide is too low, and therefore the catalyst may have too much acidity. If the mass ratio is above the specified upper preferred limit, then this means that the content of magnesium oxide is too high, and therefore the catalyst may have too low toughness and abrasion resistance.

In addition, the carrier component of the invention is characterized by its surface area. Typically, the surface area is at least about 25 m ² / g. Its preferred value is at least about 35 m ² / g, and more preferred at least 50 m ² / g. Even more preferably, if the surface area is in the range of 50-250 m ² / g, and most preferably 80-170 m ² / g. Specialists are well aware that too small surface area corresponds to low catalytic activity, while a large surface area usually corresponds to high catalytic activity. The catalytic composition of the invention has a large surface area and high catalytic activity.

After receiving the carrier, the catalyst elements of molybdenum oxide, an alkali metal promoter, and optionally vanadium oxide are applied thereto. If the ratio of MgO / Al ₂ O ₃ is within the acceptable values indicated above, then the additional addition of magnesium oxide is not necessary. Typically, the required amount of molybdenum oxide or a molybdenum precursor compound, such as ammonium heptamolybdate or ammonium carbonyl, is dissolved in a solvent to form a solution. A preferred molybdenum compound is ammonium heptamolybdate, and a preferred solvent is water. The solution is then reacted with a carrier composite manufactured as described above, and the resulting suspension is dried to remove the solvent. When said solution is aqueous, the drying is carried out in an oven at 70-120 ^C. The dried slurry is calcined to form a catalytically active composition containing an aluminum oxide, magnesium oxide and molybdenum oxide. Generally, calcination is carried out at 300-900 ^{° C} for 0.5-24 hours. The preferred calcination temperature is 500-800 ^C and more preferred 550-650 ° ^C. Alternatively, the suspension was drained as described above, it is possible, without subjecting pre calcination, directly used in the catalytic process of the invention. Since the molybdenum precursor can be converted into molybdenum oxide at ³⁰⁰ ° C or so, and the catalytic bed was heated to a temperature exceeding ³⁰⁰ ° C, then drained composition will be converted in situ into a catalytically active mixture of oxides of magnesium and molybdenum.

A mixed oxide catalyst composition typically exhibits one or more diffraction peaks during X-ray diffraction, relating to the following components: magnesium oxide, magnesium molybdate, magnesia spinel, and aluminum oxide. Elemental analysis of the calcined solid composition reveals the following composition, wt. MoO ₃ 3-50; MgO 90-10 balanced with alumina. The preferred composition is the following composition, wt. MoO ₃ 10-30; MgO 60-20, and more preferred MoO ₃ 12-25; MgO 40-25.

 An activating amount of at least one alkali metal promoter must be added to the supported catalyst described above. The promoter serves to increase the selectivity and productivity of unsaturated products, for example diolefins, in the proposed method. Said promoter is typically lithium, sodium, potassium, rubidium, cesium or fractions having sufficient basicity to increase selectivity for the more highly unsaturated compounds of the invention. Suitable promoter compounds for this purpose are alkaline oxides, hydroxides and carbonates. Also suitable are compounds which, when heated, decompose into oxides, for example, such as alkali metal acetates and oxalates. Alkali metal salts can also be used, however, alkali metal halides and silicates are undesirable due to their low basicity. Preferred promoters are: oxide, hydroxide, carbonate, acetate or alkali metal oxalate. Of these, potassium or cesium oxide or hydroxide are more preferred. The most preferred promoter is potassium oxide or hydroxide.

 The amount of alkali metal promoter is a critical parameter for catalyst performance. Usually acceptable is any amount of the indicated promoter, sufficient to increase the selectivity and volumetric yield of unsaturated products, such as diolefins, in the proposed method. Typically, the amount of alkali metal promoter ranges from 0.1-5 wt. in relation to the total mass of oxides of magnesium and molybdenum. A preferred amount of said promoter, such as alkali metal hydroxide, is 0.2-2 wt. by weight of oxides of magnesium and molybdenum, and a more preferred amount is 0.5-1.5 wt. If the amount of the indicated promoter is below the lower limit of the preferred values, then the selectivity for diolefins will be lower, since in this case the selectivity for products of deep oxidation increases. If the number of promoter exceeds the upper limit of the preferred values, then in this case there is also a decrease in selectivity for diolefins.

 Typically, the amount of alkali metal promoter calculated on the basis of alkali metal hydroxide is 0.05-5 wt. from the total mass of oxides of aluminum, magnesium and molybdenum. The preferred amount of the specified promoter, calculated on the basis of alkali metal hydroxide, is 0.1-2 wt. from the total mass of oxides of aluminum, magnesium and molybdenum, and a more preferred amount of 0.3-1.5 wt. If the amount of the indicated promoter is below the lower limit of the preferred values, then there is a lower selectivity for diolefins, and the selectivity for the products of deep oxidation increases. If the amount of said promoter exceeds the upper limit of the preferred values, then in this case there is also a decrease in selectivity and productivity for diolefins.

The alkali metal promoter can be added to the molybdate-based catalyst using standard methods. For example, a promoter can be applied using a well-known impregnation technique, described for example by Charles N. Satterfield in Heterogeneous Catalysis in Practice, McGraw-Hill Book Company, New York, 1980, pp. 82-83, which is introduced into the present description by reference. In this technique, a molybdenum-impregnated carrier is immersed in a solution of an alkali metal promoter, for example a methanol solution of an oxide or an alkali metal hydroxide. Then, excess solution, the impregnated alkali medium is drained, and the residue was subjected to drying in an oven to remove residual solvent, and calcined at 550-650 ° ^C. Alternatively said carrier is impregnated with an alkaline metal promoter via humidifying equipment so that the pores were filled with a solution of alkali oxides metal or alkali metal hydroxide, but the solution should not be in excess. The impregnated carrier thus obtained is also dried in an oven to remove solvent. In yet another alternative embodiment, the molybdenum compound may be impregnated with the same solution as the alkali metal compound.

Optionally, the molybdate catalyst of the invention may contain an activator that enhances the activity of the catalyst at any given temperature. Preferably, if the activator does not significantly reduce the selectivity to diolefins and monoolefins. It is also preferred that the activator allows the reaction to be carried out at a lower temperature while maintaining high selectivity and high diolefin productivity. Suitable activators for incorporation into the catalyst may be vanadium oxides, preferably V ₂ O ₅ . Any amount of vanadium oxide may be added to the catalyst provided that the activity of the catalyst increases and the selectivity for alkanes, including mono- and diolefins, does not decrease too much. Typically, the concentration of activator (if used) ranges from 0.05-10 wt. in relation to the total weight of the catalyst. A preferred activator concentration is a concentration of 0.1-5 wt. and more preferred 0.15-1.5 wt. Said activator may be introduced into a carrier and a suspension of molybdenum oxide prior to calcination, or it may be applied to calcined aluminum-magnesium-molybdenum oxides using the impregnation technique described above.

 The preferred industrial reactor used in the invention is a mobile bed reactor, for example with an upward catalyst bed. In such reactors, the catalytic particles are subjected to constant collisions with other catalytic particles and with the walls of the reactor. These forces gradually reduce the size of the catalyst to small dust particles that are lost in the reaction products; thus, the catalyst life is very limited. Therefore, the catalyst must be manufactured in a form that could withstand shock forces and erosion. A butane oxidation catalyst containing magnesium molybdate supported on magnesia does not have sufficient abrasion resistance required for industrial production. In contrast, an alkali metal activated catalyst of the invention based on magnesium molybdate and supported on the preferred composite material described above has sufficient abrasion resistance required for use in industrial production.

 Obviously, the wear of a catalyst containing magnesium molybdate deposited on magnesia is directly related to the high sintering temperature of magnesium oxide and strong intermolecular bonds. The sintering temperature is so much higher than the normal calcination temperature and normal working temperatures of the proposed method that the particles are not able to fuse and bind to one another. One method of hardening a catalyst containing magnesium molybdate is to introduce a carrier component into the catalyst. Preferably, said component has a large surface area and high abrasion resistance and is a composite material comprising magnesium oxide and alumina and optionally an alumina-magnesia spinel phase, as described above. It is obvious that alumina gives the catalyst hardness, while magnesium oxide and / or magnesium aluminate reduce the natural acidity of alumina. However, the above theory should not limit the disclosure of the invention. Suitable abrasion test methods are given in the examples below.

 Another important property of the catalyst of the invention is its particle size. The prototypes usually consider the use of a catalyst with small spheroidal particles, which is designed for reactors with a fixed bed and a moving fluidized bed. These particles typically have a diameter of 20-200 microns, and preferably 80-120 microns. The inventor found that spheroidal (i.e., having a shape approaching a spherical) particle size of 200-1700 microns give the best performance in reactors with a moving layer. Preferably, the particles have a size of 500-1200 microns, and more preferably 600-1000 microns. Larger particles of the invention exhibit less "caking" and therefore give a smoother, lower flow rate. Thus, larger particles give a smaller pressure drop in the sections of the dense-bed reactor and a less intense collision of particles with the walls and have the ability to better differentiate the gas residence time and catalyst residence time in the reactor.

 The proposed method can be carried out in any suitable reactor, for example, in a batch reactor, in a continuous reactor with a fixed catalyst bed, in a slurry reactor, in a fluidized bed reactor and in a reactor with an ascending compacted catalyst bed. A continuous reactor, such as a fixed-bed continuous reactor or an ascending packed bed reactor of the type described below, is preferred.

 The reactor with an upwardly compacted catalyst bed contains a vertical vessel with a relatively low diameter to length ratio. The catalyst is continuously loaded into the lower part of the reactor with an upward layer. In addition, an aliphatic hydrocarbon stream in the form of a vapor phase or a liquid phase is simultaneously supplied to the lower part of the reactor. Preferably, if the alkane is supplied in the form of a vapor phase, pre-mixed with an inert gas diluent, and optionally with a small amount of oxygen. The feed stream moves up through the reactor, reacting with the catalyst. When interacting with the catalyst, the feedstock is converted into a mixture of products, for example, monoolefins, diolefins of more highly unsaturated olefins, cracked products, deep oxidation products, such as carbon monoxide and carbon dioxide, and deep cracked products, such as benzene and furan, if butane was the feedstock. The product stream present in the bottom-bed reactor is separated by standard methods, such as distillation, in order to obtain the desired product, usually a diolefin. Unreacted alkanes are returned to the reactor for further oxidation.

 The technology used in reactors with a rising catalyst bed is suitable for the invention, since this eliminates the danger of using a mixture of alkane and / or olefin and elemental oxygen and increases the selectivity for diolefins, especially at high temperatures used in this method. In contrast, when using raw materials containing alkane and oxygen at high temperature and high molar oxygen / alkane ratios, there is a tendency to form deep oxidation products such as carbon monoxide and carbon dioxide. In addition, the danger of a reaction increases at a rapidly increasing rate.

 The operation of a reactor with an ascending catalyst bed can be modeled using the alternating pulse method. So, for example, the impulse of a stream of hydrocarbon-containing feed is passed through a catalyst bed, where this feed is oxidized, forming the target products of olefins. Then, an inert gas flow pulse is passed through the catalyst bed in order to clean the layer of residual alkanes and alkenes. After purification, a pulse of an oxygen-containing feed stream is passed through the catalyst bed in order to regenerate the catalyst. Finally, a second pulse of inert gas is passed through the catalyst bed in order to clean the bed of oxygen, after which the cycle is repeated. This procedure is used in embodiments of the invention.

Aliphatic hydrocarbon interacts with the catalyst at any operating temperature, which helps to activate the oxidation process of the invention and increase the yield of target unsaturated products. Such temperature is typically the temperature of ^about 400-700 C. The preferred temperature is 500-650 ^{° C,} and more preferably a temperature ^of 530-600 C. If the temperature is below the lower limit of the preferred range, the conversion reagent may be reduced. If the temperature exceeds the upper value of the specified range, then there is a decrease in the selectivity and productivity of diolefin products.

 Similarly, an aliphatic hydrocarbon interacts with the catalyst at any operating pressure, which helps to activate the process of the invention and increase the yield of target unsaturated products. Typically, the partial pressure of the reagent is adjusted so that the specified reagent is in a gaseous state at a given operating temperature. Preferably, the partial pressure of the aliphatic hydrocarbon is within the range of from about below atmospheric pressure to about 100 psi. inch (690 kPa). More preferably, the partial pressure is in the range of 1-30 psi (7-207 kPa). Finally, the most preferred partial pressure is between 3-15 psi (21-104 kPa).

If the method of the invention is carried out in a reactor for the continuous process described below, the flow rate of the reactants may vary. Basically, in the proposed method, an aliphatic hydrocarbon is fed into the reactor at any operating flow rate, which helps to activate the oxidation process of the invention and increase the yield and selectivity for unsaturated products. The flow rate is expressed as the volumetric hourly gas flow rate (GHSV) and is given in units of the volume of the gaseous feedstock containing aliphatic hydrocarbon to the full volume of the reactor in 1 h or simply h ^-1 . Typically, these values range from 100-20000 h ^-1 . It should be noted that the space velocity controls the residence time of the reactants in the reactor. For example, in a bed reactor, the preferred residence time of the gas in the reactor is less than 10 s, more preferably less than 5 s and most preferably less than 1 s.

In the case of using a reactor with an upward catalyst bed after interaction of the catalyst with an aliphatic hydrocarbon, the spent catalyst is removed from the upper part of the reactor and transferred to the second reactor for regeneration. Regeneration is carried out using the reaction of interaction with oxygen. Typically, a preheated oxygen source similar to that described above is supplied to the bottom of the second reactor. The spent catalyst interacts with the specified oxygen source at any temperature, pressure and flow rate of the oxygen source, which are sufficient for regeneration of the catalyst. The parameters of this process should, however, be adjustable, so as to exclude the occurrence of an uncontrolled reaction or excessive heating. Preferably, the temperature ranges ^of 500-700 C and more preferably in the range of 550-650 ^C. The preferred pressure ranges from subatmospheric pressure and up to 100 pounds / square inch (690 kPa), and more preferably 2 -50 psi (14-345 kPa). The flow rate of the oxygen source depends on the heat transfer properties of a particular reactor. For example, if the flow rates are too high, an excessive temperature rise can occur, resulting in an uncontrolled reaction.

 In the interaction of an aliphatic hydrocarbon with a catalyst of the invention, the oxidation of an aliphatic hydrocarbon results from the loss of at least two hydrogen atoms from the hydrocarbon reactant to form water as a by-product. The resulting organic products are mainly unsaturated hydrocarbons, such as monoolefins and diolefins. These unsaturated products usually contain the same number of carbon atoms as the reagent aliphatic hydrocarbon. Therefore, these products are not cracked products, since the latter contain fewer carbon atoms than the original hydrocarbon. Generally unsaturated products have a higher degree of unsaturation than reagent hydrocarbon. For example, alkanes such as butane can lose two hydrogen atoms to form monoolefins such as 1-butene, trans-2-butene, and cis-2-butene. In turn, monoolefins, such as the above butenes, can lose two hydrogen atoms to form 1,3-butadiene.

Preferred diolefin products may be represented by the general formula:
CH ₂ = CH — CH = CH— (CH ₂ ) _m —H, where m is an integer from 0 to 6. Preferably, m is from 0 to 2. A more preferred value is m 0, and a more preferred unsaturated product is 1, 3-butadiene. Isomers of the above formula may also be formed, where unsaturation may occur at any other position of the carbon chain. Preferably, if unsaturation is formed by conjugated double bonds. Even more unsaturated variants of the general formula can be formed by further oxidation, which gives more than two ethylene double bonds. However, alkanes do not form in significant quantities.

 Along with alkenes, the product stream may contain various by-products. For example, if the saturated alkane is n-butane, small amounts of cracking products such as propylene and ethylene, as well as heavy fractions such as benzene and furan, and in addition, deep oxidation products such as monoxide and dioxide, may form. carbon. However, it was unexpectedly discovered that these by-products, especially the deep oxidation products, are formed in much smaller amounts with the proposed method of the invention.

For the purposes of the invention, “conversion” is defined as the percentage (in moles) of reactive aliphatic hydrocarbon lost from the feed stream from the reaction. The conversion can vary widely depending on the reagents, the form of the catalyst, and reaction conditions, such as temperature, pressure, flow rate, and residence time of the catalyst in the reactor. Within the preferred temperature range, with increasing temperature, the conversion usually increases. Within the preferred range of the hourly gas volumetric flow rate, with an increase in the volumetric flow rate, conversion usually decreases. Typically, the conversion of an aliphatic hydrocarbon is at least about 10 M%. Preferably, if the conversion is at least about 20 M%, more preferably at least about 40 M% and most preferably at least about 50 M%
Similarly, for the purposes of the invention, “selectivity” is defined as the percentage (in moles) of converted carbon that forms a particular product. Typically, selectivity varies widely with reagents, catalyst form, and reaction conditions. Typically, the proposed method achieves high selectivity to diolefins. Within the preferred temperature range, with increasing temperature, selectivity usually decreases. Within the preferred range of volumetric flow rate, with increasing volumetric flow rate, selectivity for alkanes usually increases. Preferably, the total selectivity to all alkenes is at least about 50 M%, more preferably at least 60 M%, even more preferably at least 70 M%, and most preferably at least about 80 M%. Typically, the selectivity for diolefins is at least about 40 M%. Preferably, the selectivity for diolefins is at least about 50 M%, more preferably at least about 60 M%, most preferably at least about 70 M%
The concept of both high conversion and high selectivity can be defined as a way out. For the purposes of the invention, the term “yield” means the amount of product resulting from conversion and selectivity with a single pass through the reactor. For example, in the implementation of the proposed method with a conversion of 0.65 or 65 M% and selectivity for diolefins 0.75 or 75 M%, the yield of diolefins is 0.49 or 49 M%. Typically, in the proposed method, the yield of diolefins is at least about 8 M% Preferably, the yield of the diolefin is at least about 18 M%, more preferably at least about 28 M%, and most preferably at least 35 M%. As a rule, when the butane is oxidized, the yield of the total C ₄ olefins is at least about 25 M% Preferably, if with oxide enii butane total yield of C ₄ olefins is at least about 30 M%, more preferably at least about 35 M% and most preferably at least about 40 M%
The speed at which the target product is formed in the proposed method can be determined by the concept of volumetric yield. In the invention, the concept of "volumetric yield" is defined as the yield of the target product (expressed in M% per hour (hour ^-1 )) and is a numerical expression of the conversion product at a single pass with the corresponding selectivity, hourly gas volumetric flow rate and concentration of aliphatic hydrocarbon in the feedstock, where conversion, selectivity and concentration are expressed in decimal fractions. Preferably, the volumetric yield of the diolefin is at least about 30% in 1 hour, more preferably at least about 120% in 1 hour, and more preferably at least about 200% in 1 hour.

 Another measurement of the speed at which the target product is produced is "productivity", defined as grams of the formed target product per gram of catalyst in one hour (g / g cat.-h). In the proposed method, the preferred productivity of butadiene is at least about 0.2 g / g cat-h, the more preferred productivity is at least about 0.4 g / g cat-h, and the most preferred productivity is at least about 0.5 g / g cat. The preferred productivity of total C-olefins in the proposed method is at least about 0.3 g / g cat.-h; more preferred productivity is at least about 0.4 g / g cat.-h; and the most preferred productivity is at least about 0.9 g / g cat.

Catalyst abrasion testing requires a large number of catalyst samples. It would be desirable to develop a simple abrasion strength test procedure for small catalyst samples. Such a procedure is a crush strength test, since an increased crush strength is an indicator of increased abrasion resistance. The crushing strength can be determined using any standard installation that is commonly used for this purpose, for example, Instron Model 1125. Usually in this case, a catalyst is used that is first sieved to select 8-mesh granules (2.36 mm), and then these granules were calcined at 600 ^{° C} for 2 hours, after which the crush strength. The crushing strength of the preferred catalysts of the invention is at least about 5 pounds (2270 g), preferably at least about 10 pounds (4540 g); and more preferably at least about 15 pounds (6810 g).

The actual abrasion resistance of a preferred catalyst of the invention can be determined using a standard installation that is commonly used for this purpose. Suitable test instruments include a vertical stainless steel pipe with a diameter of 1/2 inch (13 mm) and a length of about 30 inches (760 mm), which is connected via an I-valve to a vertical column of a similar diameter and a length of about 52 inches (1320 mm). A stainless steel cyclone with a diameter of 3 inches (76 mm) is connected to the specified vertical column at the inlet of the cyclone and to the specified vertical pipe at the outlet of the cyclone, thereby forming a circulation circuit in which gas and solid phases are separated, after which the solids return into a vertical pipe. Typically, the test material is loaded into the system and subjected to fluidization by introducing gas into the I-valve at a flow rate of 0.11-1.01 s ^-1 . The powder trapped in the gas moves in an upward flowing gas stream and, after separation by gravity, returns to the cyclone regeneration system. Typically, the total time for testing samples takes about 15 hours. Then, the number of small particles formed per unit time is determined and the attrition rate is calculated in comparison with the initial coarse-grained fraction. The preferred rate of abrasion of the catalyst of the invention is approximately less than 5 wt. for 1 h (wt.˙h ^-1 ); and more preferably, approximately less than 1 wt.h ^-1 .

 The invention is illustrated in detail with specific examples, which, however, do not limit the possible options for its implementation. All percentages are given in M% carbon unless otherwise specified.

PRI me R 1 (a-b). The manufacture of the catalyst. a) Magnesium oxide powder (17 g, Magox Premium Grade MgO) was dispersed in deionized water (96 g) using a high shear dispersant. To the resulting suspension was gradually added 20 wt. colloidal alumina (200 g, Nyacol) in order to obtain a viscous mixture containing, by weight. MgO 30; Al ₂ O ₃ 70. Then another 52 g of water was added to obtain controlled rheological properties. The resulting mixture was kept and gelled on a hot plate, stirring at 70 ^about C for 2 hours to obtain a white solid. This material was crushed and then heated for 4 hours at ⁶⁰⁰ ° C and calcined another 4 hours and at ⁶⁰⁰ ° C, resulting in a composite of the spinel-based carrier having a weight ratio of MgO / Al ₂ O ₃ 0.43 and surface area 169 m ² / g. The X-ray diffraction spectrum of the support gives a diffraction pattern of crystalline MgO, as well as spinel and alumina. The media was tested on an Instron Model 1125 instrument and the average crush strength was determined, which was 15 pounds (6810) g.

b) A spinel composite carrier was made according to the procedure of Example 1a, except that 44 g of magnesium oxide powder, 280 g of colloidal alumina and 249 g of deionized water were used. The resulting composite support had a MoO / Al ₂ O ₃ mass ratio of 0.78 and a surface area of 162 m ² / g. The X-ray spectrum of the carrier showed a reflective pattern, which may correspond to crystalline MgO, as well as spinel and alumina. The spinel composite carrier (20 g) was impregnated to the initial stage of humidity with 20.44 g of a solution containing, by weight. ammonium heptamolybdate 35.25; cesium hydroxide 1.85. The impregnated solution was dried for 2 hours at 125 ^{° C} was heated at ⁶⁰⁰ ° C for 4 hours and calcined at ⁶⁰⁰ ° C for a further 3 hours. The catalyst obtained (EI) contained, by weight. MoO ₃ 22.4; Cs ₂ O 1.20, and had a surface area of 122 m ² / g.

PRI me R 2. Oxidation of butane. The catalyst (EI) prepared as described in Example IB was used to oxidize butane as follows: approximately 15 cm ^{3 of} catalyst were loaded into a chamber (18 mm outer diameter x 7.6 cm length) of a Vycor reactor. The temperature of the reactor was determined using a stainless steel thermocouple pocket (1/8 inch (3.2 mm) outer diameter) embedded in a catalyst sample. The stream of feedstock containing butane 10-20 vol. and helium 90-80 vol. passed through the catalyst for approximately 10-30 seconds. Then, the feed stream was stopped and through the catalyst for 1 min and a purification stream containing pure helium was passed at the same speed. Then, the purification stream was stopped through the catalyst for 1 min. and an oxygen stream (20 vol.) in helium was passed at the same speed, after which another helium stream was passed for cleaning for 1 min. This cycle was repeated and the combined products were collected for analysis and the Saran ^® polyvinylidene chloride plastic chamber was analyzed using a Carle gas chromatograph designed to analyze C ₁ -C ₅ alkanes, alkenes and alkadienes, as well as constant gases such as N ₂ , O ₂ , CO, CO ₂ , H ₂ and heavy products, including furan, benzene and C ₆ compounds. Isobutane was mixed with the feedstock or with products as a standard. "Unknown" were obtained from the difference between the carbon balance and 100%. The results and process conditions are presented in table 1.

 As can be seen from Table 1, that a cesium activated catalyst containing magnesium molybdate and deposited on an aluminum-magnesium spinel composite catalyzes the oxidation of butane to butadiene and butene with high selectivity.

PRI me R 3. The manufacture of the catalyst. An aluminum-magnesia spinel composite carrier was made as described in Example 1b. The spinel composite (20 g) was impregnated to an initial moisture content with 19.49 g of a solution containing, by weight. cesium hydroxide 1.66; ammonium vanadate 1.39; ammonium heptamolybdate 27.22. The impregnated carrier was dried at ¹²⁵ ° C for 2 hours, heated to 600 ^{° C} for 4 hours and calcined at ⁶⁰⁰ ° C for a further 3 hours. The resulting catalyst (EZ) contained, by weight. MoO ₃ 17.44; V ₂ O ₅ 0.85; Cs ₂ O 1.43, and had a surface area after calcination of 122 m ² / g.

 PRI me R 4 (a-b). Oxidation of butane. The catalyst obtained in accordance with the description in example 3 (EZ) was used in the oxidation of butane, carried out in accordance with the description in example 2. The process conditions and results are presented in table 1. It can be seen from Table 1 that the cesium activated catalyst containing molybdenum of magnesium and vanadium oxide and supported on the alumino-magnesium spinel composite catalyzed the oxidation of butane to butadiene and butene with a high degree of selectivity.

PRI me R 5. Production of media. Magnesium oxide powder (600 g) was dispersed in deionized water (5733 g) using a high shear dispersant. Colloidal alumina (7000 g, 20 wt.) Was gradually added to the resulting suspension at low shear. The mixture thus obtained was spray dried using an atomizer with a nozzle having an opening of 2 mm diameter under a pressure of 40 psi (276 kPa). Input injector temperature was 300 ^{° C,} and the output 120 ^C. The resulting spray drying a white powder had excellent flow characteristics. Then the specified powder was calcined at 600 ^about C for 2 hours, resulting in a composite carrier having a surface area of 184.3 m ² / g and a mass ratio of MgO / Al ₂ O _{3 of} 0.43. The X-ray spectrum of the carrier showed a reflective pattern, which may correspond to crystalline MgO, as well as spinel and alumina. Calcination part powder for 4 hours ^at 800 ° C only slightly reduces the surface area of up to 143.2 m ² / g. Carrier obtained by calcining the powder at ⁶⁰⁰ ° C, had an average particle size of about 60 microns. Examination of particles using a scanning electron microscope showed the presence of spheroidal particles having excellent flow characteristics. In addition, the carrier was subjected to an abrasion test in the circulation circuit, as described, with the result that the abrasion resistance index was obtained equal to 0.28 wt.h ^-1 . It can be seen from the comparison that this support is clearly superior to the commercial alumina-based fluidized cracking catalyst (FCC), which is commonly used in mobile bed reactors and has an abrasion index of 0.99 wt. ˙h ^-1 .

Examples of catalysts (bd). Carriers based on alumina-magnesia spinel were obtained in accordance with the description in paragraph (a), except that these carriers contained a sufficient amount of MgO to obtain a composition with 40, 44 and 30 wt. MgO; and besides, they were not calcined. Then, these carriers were compressed at 5 kilo pounds / sq. inch (35 MPa) in the isostatic press with a refinement to 20-80 mesh (850-180 microns) and calcined at ⁶⁰⁰ ° C for 5 hours. To 50 g of each support was added a solution comprising, by weight. ammonium heptamolybdate 33.23; K ₂ CO ₃ 0.62; hydrogen peroxide 2.5 from 30 wt. peroxide solution, neutralized to pH 9 with ammonium hydroxide. The impregnated supports (E-5-bd) dried for 18 hours at 110 ^C and calcined for 3 hours at 600 ^C, to obtain catalysts. The amount of solution added to each carrier is shown in Table 2.

 PRI me R 6. Oxidation of butane. The catalysts obtained in accordance with the description in example 5 (E-5b-E-5d) were used to oxidize butane by the method described in example 2. The conditions of the method and the results are shown in table. 2. As can be seen from Table 2, potassium-activated catalysts containing magnesium molybdate and deposited on a composite of aluminum-magnesium spinel catalyzed the oxidation of butane to butadiene and butene with a high degree of selectivity.

PRI me R 7. The manufacture of the catalyst. The spinel precursor was prepared by a standard method for co-precipitating magnesium and aluminum hydroxides (at pH 9) from an aqueous solution of magnesium and aluminum salts by adding NaOH and Na ₂ CO ₃ . The spinel precursor was calcined at ⁶⁰⁰ ° C for 3 hours, yielding alyumomagnezialnuyu spinel (MgAl ₂ O _4). The resulting spinel was impregnated to incipient stage of humidity using 1.75 M aqueous solution of magnesium atsetattetragidrata (37 wt.). The impregnated spinel was calcined for 3 hours at 600 ^{° C,} to yield the spinel composite containing, besides 17.1 wt. MgO and having a mass ratio of MgO / Al ₂ O _{3 of} 0.69. Next, the composite (16.3 g) was impregnated to incipient wetness step with an aqueous solution (17.53 g) of ammonium heptamolybdate (26.5 wt.) And then calcined for 4 hours at 600 ^C. This material was then impregnated with a methanolic a solution (8.5 g) of cesium hydroxide (0.63 wt.) and again calcined for 3 hours at 600 ^about C. The resulting catalyst (E7) contained, wt. MoO ₃ 18.8; Cs ₂ O 0.25, and had a surface area of 115 m ² / g.

 PRI me R 8 (a-b). Oxidation of butane. The catalyst of example 7 (E7) was used in the oxidation of butane, carried out in accordance with the description in example 2. The results are presented in table.3.

 As can be seen from Table 3, a cesium activated catalyst containing magnesium molybdate and deposited on a spinel composite catalyzed the oxidation of butane to butadiene and butene with a high degree of selectivity.

 Comparative example 1 (a-b). In accordance with the description in Example 7, a composition was prepared with the only exception that the catalyst was not impregnated with cesium hydroxide. The resulting composition was used in the oxidation of butane, carried out in accordance with the description in example 2. The results are presented in table. 3. From a comparison of Example 8 and Comparative Example 1, it can be seen that the catalyst of the invention containing a small amount of cesium promoter has a higher selectivity for butadiene and butene and a lower selectivity for deep oxidation products than a similar catalyst that does not contain a cesium promoter.

PRI me R 9 (ae). The manufacture of the catalyst. An aqueous solution (1184 g) containing 71.5 wt. Magnesium nitrate hexahydrate was added to porous alumina beads (713.3 g; VOP, 700 μm). The resulting slurry was dried at ¹⁵⁰ ° C for 18 hours and calcined by heating to 460 ^{° C} for 4 hours, holding at 460 ^{° C} for 2 hours, another two-hour heating to 540 ^{° C} and holding at 540 ^C. 3 hours in a stream of air. Then, 1186.9 g of magnesium nitrate solution was added to the dried alumina balls, after which the drying and calcination procedures were repeated. The resulting composite carrier contained 27.7 wt. MgO and had a mass ratio of MgO / Al ₂ O _{3 of} 0.38 and a surface area of 101 m ² / g Samples (35 g) of the composite were impregnated with a solution containing ammonium heptamolybdate and potassium carbonate, as shown in Table 4. The concentration of ammonium heptamolybdate in each solution was different, while the concentration of potassium carbonate was approximately the same. The impregnated composites were calcined at ⁶⁰⁰ ° C for 4 hours, to obtain activated potassium molybdate catalysts (E9 a-e) supported on and alyumomagnezialny spinel composite characterized in Table 4.

PRI me R 10 (a-e). Oxidation of butane. The catalysts (E9 a-e) obtained in example 9 were used in the oxidation of butane in accordance with the description given in example 2, the results of which are presented in table 4. From the table. Figure 4 shows that a potassium-activated molybdate catalyst supported on an alumino-magnesium spinel composite catalyzed the oxidation of butane to butenes and butadiene with a high degree of selectivity. In addition, it was found that the selectivity to C ₄ olefins decreases if the concentration of molybdate drops below 11 wt.

PRI me R 11. The manufacture of the catalyst. Alumina balls (4.09 g; UOP-SAB16) having a nominal diameter of 1700 μm were impregnated three times to the initial moisture stage with 10 cm ^{3 of} each solution containing 1.15 g of MgO in the form of magnesium nitrate hexahydrate.

Between each addition of solution, the impregnated alumina was calcined for 2 hours at 600 ^C, yielding a composite support having a weight ratio of MgO / Al ₂ O ₃ 0.75. To a sample (6.42 g) of the support was added a solution (9.2 ml) containing ammonium heptamolybdate (11.8 g) K ₂ CO ₃ (0.228 g) and NH ₄ VO ₃ (0.60 g) per 100 ml of solution , after which the carrier was calcined and received a catalyst (E-11) containing, by weight. MoO ₃ 12.03; V ₂ O ₅ 0.58; K ₂ O 0.16, and having a surface area of 34 m ² / g.

PRI me R 12. The oxidation of butane. The catalyst (5.65 g) Example 11 (E-11) was loaded into a fixed bed reactor as described in Example 2. Then, at ⁶⁰⁰ ° C and 1200 h ^-1 GNSV passed through the reactor pulse flows: a mixture of helium and 20 vol. butane for 5 s, and then for 60 s, respectively, helium, air and helium. Then the cycle was repeated and the following results were obtained: butane conversion 45.2% selectivity 58.3% for butadiene, 20.6% for butene; 78.9% and 78.9% for total butenes; 6.4% for products of deep oxidation (CO ₂ and CO); 8.0% for cracked foods and 6.8% for unknown foods. Productivity was 0.34 g C ₄ / g cat. and 0.25 g of butadiene / g cat. hours Obviously, the productivity of C ₄ -carbon and butadiene is greater than the productivity adopted for the standard method of 0.2 g / g cat.

PRI me R 13. Obtaining a catalyst. Alumina beads (713.3 g; UOP-MS-P-27) with particles with a diameter of 500-1000 μm were used in the preparation of a composite carrier containing alumina-spinel, by the method described in example 2. The resulting composite carrier had a mass ratio of MgO / Al ₂ O ₃ 0.38 and a surface area of 106 m ² / g. A sample (686 g) of the composite was impregnated with 728 g of a solution containing, by weight. ammonium heptamolybdate 22.5; potassium carbonate 0.66. The impregnated composite was calcined for 4 hours at 600 ^{° C,} to give a catalyst composition (E-13) containing, wt. MoO ₃ 18.4; K ₂ O 0.45.

PRI me R 14. The oxidation of butane. The catalyst obtained in accordance with the description of example 13 was used in a laboratory reactor with an upward catalyst bed. The feed tank for the fluidized bed catalyst (with a capacity of 3500 ml) is connected to its lower part by means of a long tube, called a "catalytic tube". The specified vertical pipe is connected via valve "J" with a vertical reactor chamber with a length of 2 m and a diameter of 6 mm, which is connected to the first of three cyclone separators connected in series with each other. The catalyst recovered in cyclones returned to the top of the tank to feed the catalyst. Gaseous feed stream containing, vol. nitrogen 10; butane 6.9 and balanced helium were preheated in a heating chamber having only 4 mm in diameter, which minimizes the residence time and, consequently, thermal cracking of butane. The preheating chamber was connected to the system at the junction of the “catalytic pipe” and valve “J”, and the gaseous feed stream was fed at a speed of 500 cm ³ / min. Then the catalyst was heated to 650 ^° C in the feed chamber, fed into the reactor chamber at a speed of 200 cm ³ / min. The reactor chamber temperature was about 580 ° ^C. The gas residence time in the reactor was 1.6 s. As a result of the described procedure, the following results were obtained: butane conversion 23.3% selectivity, butadiene 61.9; butenes 22.1; deep oxidation products (CO ₂ and CO) 5.1; C _2-3 -cracking products 7.7; methane 1.0; unknown 2.1. Productivity was about 0.43 g C ₄ / g cat-h and 0.32 g butadiene g cat-h From the above results it is seen that the productivity of C ₄ -carbon and butadiene exceeds the value of productivity adopted for the standard method (0.2 g / g cat.-h).

PRI me R 15. Oxidation of butane. The catalyst of example 13 was used in the reactor with an upward catalyst bed described in example 14, together with a gaseous feed stream containing 67 vol. butane and balanced nitrogen and helium. The gas feed rate was 1000 cm ³ / min, and the gas residence time in the reactor was about 0.82 s. The catalyst was fed at a speed of about 455 cm ³ / min, while the residence time of the catalyst in the reactor was about 3.1 s. The average temperature in the riser was 584 ^C. As a result, the following data were obtained: conversion of butane 17.1; selectivity, butadiene 39.1; butanes 2,4; deep oxidation products (CO and CO ₂ ) 2.5; cracked products, including C ₂ -C ₃ products 10.6; methane 1.6; isobutane 0.5; unknown products 1.8. The productivity of the combined butadiene and butene products was 0.94 g C ₄ / g cat-h, and the productivity of butadiene was 0.49 g butadiene / g cat-h From the above results it is seen that the productivity of C-carbons and butadiene exceeds the value of productivity adopted for the standard method (0.2 g / g cat.-h).

 PRI me R s 16-20 (a-b).

 a) The manufacture of catalysts activated by cesium 16-20. A series of cesium doped molybdate catalysts was prepared in accordance with the procedure described below, the amount (g) of cesium hydroxide monohydrate per 40 g of methanol solution being varied as follows, g: catalyst 1 0.16; catalyst 2 0.23; catalyst 3 0.29; catalyst 4 0.53; and catalyst 5 0.76.

(NH ₄ ) ₆ Mo ₇ O ₂₄ ˙ 4H ₂ O (72 g, 0.0583 M) was dissolved in 175 ml of water. The resulting solution was mixed with magnesium hydroxide 300 g of Alfa Research Chemcal Materials) to obtain a suspension. The suspension was dried at ⁷⁰ ° C for 18 hours and at 130 ^{° C} for 2 hours to obtain a solid mixture. The resulting solid was crushed and passed through a sieve to obtain particles having a size of 20-120 mesh. (850-122 μm). The sieved particles were calcined by heating slowly over 5 hours to 600 ^{° C,} and then kept at this temperature for 2 hours. The calcined solid was cooled to room temperature to provide the magnesium molybdate comprising, by weight. MoO ₃ 22; MgO 78. A methanol solution was prepared by adding an appropriate amount of cesium hydroxide as described above to methanol, whereby 40 g of a methanol solution was obtained. Magnesium molybdate (20.6 g) was immersed for 15 min in 40 g of methanol solution. The resulting mixture was filtered and the filtered solid was dried in air at room temperature (about 22 ^C) for 4 hours and then dried in an oven at ¹²⁰ ° C for 16 h. The dried solid was calcined by heating over 5 hours to a temperature 600 ^о С followed by keeping at this temperature for another 2 hours, as a result of which a cesium activated catalyst containing magnesium molybdate was obtained. The percentage of cesium in the catalysts 16-20 are shown in table.5.

b) Oxidation of butane using catalysts 16-20. The catalysts 16-20 obtained above were tested by using them in the oxidation of butane in accordance with the following procedure. The catalyst (15 cm ³ ) was loaded into the glass chamber of the Vycor reactor (7.6 cm long x x18 mm external diameter). In its upper part, the specified glass chamber is connected to the inlet pipe for the flow of raw materials, and in its lower part to the outlet pipe for collecting products. The temperature in the reactor, measured from the pocket of a thermocouple made of stainless steel with ext. diam. 3.2 mm (1/8 ") introduced into the catalyst bed is understood to 550 ° ^C. A feed stream comprising butane about 10;. 50 nitrogen, helium 40 is passed through the catalyst bed at a rate of 84 ml / min (GHSV 336 h ^-1 ) for 30 s. Then the feed stream was stopped and a helium stream was passed through the catalyst bed for purification at a rate of 84 ml / min for 1 min. Then, the purification stream was stopped and passed through the catalyst bed at the same speed for 1 min oxygen flow containing 10 vol% oxygen in helium, then the oxygen flow was stopped and the catalyst bed was passed for 1 minute at the same speed of the helium flow for cleaning. After that the complete cycle was repeated for 30 min, and the combined products were collected in a plastic tank containing polyvinylidene chloride (Saran ^®). The resulting products were analyzed by gas chromatography using a multicolumn Carle gas chromatograph equipped with columns with a diameter of 1/8 "(3.2 mm) having the following contents: (1) 2.70% carbovax (Carbowax 1540) on Porasil Porasil (21" (530 mm), 80/100 mesh. (180/150 microns)); T _max 175 ^{° C} (2) 3,0% Carbowax 1540 on Porasil (4 '(1220 mm), 60/80 mesh (250/180 micron).), T _max ¹⁵⁰ ° C; (3) 27.5% 2 (XE) A on Chromosorb PAW (17 ', (5180 mm) 45/60 mesh (328/250 micrometers), T _max ¹⁵⁰ ° C. (4) Porapak Q 9' (2745 mm), 50/80 mesh (292/180 micrometers), T _max 250 ^{° C,} (5) molecular sieve 13X (9 '(2745 mm)., 45/60 (328/60 microns) mesh ^of 300 T _max. C (6) molecular sieve 5A (3 '(914 mm), 80/100 mesh (180/150 micron).), T _max ^{300 °} C (7) 28% DC 200/500 on Chromosorb PAW 3,5 '(1068 mm), 60/80 mesh. (250/180 micron)), T _max 175 ° ^C. The results are shown in Table. 5. From these data it can be seen that in the presence of cesium-doped catalyst containing magnesium molybdate, butane oxidized mainly to butadiene and butenes.

Comparison Example 2. A composition was prepared by the method described above for the manufacture of catalysts 16-20, except that the cesium doping step was not carried out. The resulting composition contained, wt. MoO ₃ 22; MgO 78. As already indicated, this composition did not contain cesium. Then the specified composition was used as a catalyst in the oxidation of butane, by the method described in examples 16-20 (b). The results are presented in table. 6.

A comparison of the results of comparative example 2 (GHSV 336 h ^-1 and the results of examples 16-20 shows that the presence of cesium in the catalyst significantly improves the selectivity of heavy products such as furan and benzene, and especially to deep oxidation products such as carbon monoxide, However, a comparison of the results of comparative example 2 (GHSV 336 h ^-1 ) with the results of examples 16-20 showed that the presence of cesium in the catalyst reduces the conversion of butane.

 PRI me R s 21-25. a) Production of alkali activated catalysts 21-25. A series of alkali-doped catalysts containing magnesium molybdate was prepared according to the procedure described below, the amount of alkali metal hydroxide per 40 g of methanol solution being varied as follows: Catalyst 6, KOH 0.049 g; catalyst 7, KOH 0.098 g; catalyst 8, KOH 0.133 g; catalyst 9, KOH 0.320 g; and catalyst 10, NaOH 0.073 g.

(NH ₄ ) ₆ Mo ₇ O ₂₄ ˙ 4H ₂ O (72 g; 0.0583 M), was dissolved in 345 ml of water, and the resulting solution was added to magnesia (300 g; CE Basic Industries, Madox Premium magnesia to obtain a suspension. The suspension was dried for 18 hours at ⁷⁰ ° C and for 2 h at ± 30 ^{° C} to obtain a solid mixture. The resulting mixture was ground and passed through a sieve to obtain particles having a size of 20-120 mesh. (850-122 microns ). The sieved particles were calcined by heating slowly over 5 hours to 600 ^{° C} temperature and then maintained at this temperature for 2 hours. The calcined solid Cooling If it reached room temperature, magnesium molybdate containing, by weight of MoO ₃ 22; MgO 78 was obtained. A methanol solution was obtained by adding an appropriate amount of alkali metal hydroxide, as described above, to 40 g of a methanol solution. Magnesium molybdate was immersed for 15 minutes in a methanol solution . The resulting mixture was filtered and the filtered solid was dried in air at room temperature for 4 hours and then in an oven at ¹²⁰ ° C for 16 h. The dried solid was calcined by heating over 5 hours to 600 ^{° C,} and then kept at this temperature for another 2 hours, resulting in an activated alkali metal catalyst containing magnesium molybdate. The percentage of alkali metal in the catalysts 21-25 are shown in table.7.

b) Oxidation of butane using catalysts 21-25. The potassium and sodium doped catalysts 21-25 obtained above were used in the oxidation of butane by the method described in Examples 16-20 b. The results are presented in table.7. The data obtained showed that in the presence of a catalyst containing magnesium molybdate and doped with potassium and sodium ions, butane is predominantly oxidized to butadiene and butanes. In addition, from a comparison of the results of examples 21-25 and the results of comparative example 2 (GHSV 336 h ^-1 ) it is seen that when using catalysts doped with potassium and sodium (catalysts 21-25), a significantly higher selectivity for butadiene and butanes is achieved, than with unalloyed catalysts containing magnesium molybdate. At the same time, the selectivity to deep oxidation products when using alkali doped catalysts is significantly reduced.

EXAMPLES 26 and 27. a) Preparation of potassium activated catalysts 26 and 27. Two solutions were prepared containing (NH ₄ ) ₆ Mo ₇ O ₂₄ ˙H ₂ O (35 g) and potassium carbonate in 125 g aqueous solution. The first solution contained 0.63 g (catalyst 26) of potassium carbonate, and the second solution contained 1.44 g (catalyst 27) of potassium carbonate. Each solution was then added, with stirring, to magnesium oxide (100 g; CE Basic Industries, Madox Premium Grade magnesia and the resulting mixtures were dried and calcined as described in Preparation Examples Catalysts 16-20, whereby potassium-doped catalysts containing magnesium molybdate were obtained .

b) Oxidation of butane using catalysts 26 and 27. Catalysts 26 and 27 were used in the oxidation of butane (examples 26 and 27, respectively) according to the procedure described in examples 16-20. The results obtained under various conditions of the method are presented in table.8. The data in Table 8 show that in the presence of a catalyst containing magnesium molybdate and doped with potassium carbonate, butane is predominantly oxidized to butadiene and butanes. These data also show that increasing reaction temperature from 550 to 570 ^C. butane conversion increases, while selectivity to butenes and butadiene decrease. In addition, the results show that when the volumetric gas flow rate decreases from 600 to 336 h ⁻¹ , butane conversion increases, and the selectivity to butanes and butadiene decreases. A comparison of examples 26 and 27 shows that the catalyst containing a larger amount (wt.) Of potassium gives a lower conversion of butane and higher selectivity for butenes and butadiene. A comparison of examples 27 (a) and 27 (c) with comparative examples 2 (b) and 2 (a) shows that the catalyst containing potassium ions gives a significantly higher selectivity to butadiene and lower selectivity to heavy products and to products deep oxidation than a catalyst that does not contain alkali metal ions.

Example 28. a) Preparation of Potassium Doped Catalyst 28. Ammonium Vanadate NH ₄ VO ₃ (2.55 g; 0.022 M) was added to 100 ml of water and the resulting mixture was heated to 60 ^° C in order to accelerate dissolution, after then the solution temperature was raised to about 97 ^{° C} and stirring, were added (NH _{4) 6} Mo ₇ O ₂₄ ˙H ₂ O (18,5 g, 0.015 mol). The resulting solution was boiled until its volume decreased to 70 ml, after which magnesia (34 g; 0.85 M) was added and a thick paste was obtained. Then, another 50 ml of water was added to this thick paste and a uniform, siliceous mixture was obtained. The mixture was poured into a quartz crucible, dried overnight in air and then for 2 hours, dried at ¹¹⁰ ° C and calcined at ⁶⁰⁰ ° C for 2 hours to yield a calcined solid containing, wt. V ₂ O ₅ 4; MoO ₃ 30; MgO 65. Then, 8.2 ml of a methanol solution containing sodium hydroxide (1 g KOH / 100 ml of solution) was diluted with an additional amount of methanol to a total volume of 16 ml. Then, the diluted solution was added dropwise to the above calcined solid material (7.8 g). Potassium hydroxide impregnated solid was air dried for 30 minutes and then for 1 h was dried in an oven at 100 ^C. The resulting solid contained vanadium oxide, magnesium oxide, molybdenum oxide and about 1 wt. potassium calculated as potassium hydroxide.

 b) Oxidation of butane using catalyst 28. The catalyst 28 obtained above was used in the oxidation of butane, carried out in accordance with the procedure described in examples 16-20. The results are presented in table.9.

 It can be seen from the results that in the presence of a catalyst containing magnesium molybdate, vanadium oxide and potassium ions, butane is oxidized predominantly to butadiene and butenes.

Comparison Example 3. The material for the comparative example was obtained as described in example 28 except that the catalyst was not impregnated with potassium hydroxide. The resulting material contained, wt. V ₂ O ₅ 4; MoO 30; MgO 66. This material was then used in the oxidation of butane in accordance with the procedure described in examples 16-20. The results obtained are presented in table.9. Comparison of comparative example 3 with example 28 shows that the presence of a potassium promoter in the catalyst contributes to a significant increase in the selectivity for butadiene and butanes and a decrease in the selectivity for deep oxidation products.

PRI me R 29. a) Preparation of potassium activated catalyst 29. (NH ₄ ) ₆ Mo ₇ O ₂₄ ˙ 4H ₂ O (13.5 g; 0.11 M) was dissolved in 120 ml of water, and the resulting mixture was heated up to 60 ^about With to obtain a solution. Then, gradually and intensively stirring, magnesium oxide (39 g) was added to the solution over 5 min, as a result of which a uniform suspension was obtained. This slurry was dried overnight at ¹¹⁰ ° C to obtain a sintered material which was subsequently calcined in air at 600 ^{° C} for 2 hours, thereby obtaining a mixture of magnesium oxide and molybdenum oxide. This powdery oxide mixture was compacted by compression in a Carver press with a force of 20,000 pounds (9,000 kg) to give 1 tablet in 1/8 "28.6 mm. This tablet was crushed into several pieces and soaked in a methanol solution of potassium hydroxide containing 1 g KOH per 100 ml of methanol. potassium hydroxide Dosage reached 0.5 wt. potassium hydroxide impregnated catalyst containing magnesium molybdate, dried in air for 2 hours and then another 2 hours in an oven at 110 ° ^C.

 b) Oxidation of butane using catalyst 29. A potassium hydroxide-impregnated catalyst containing magnesium molybdate (catalyst 29; 7.15 g) obtained as described above was charged into a pulsed tube reactor as described in Examples 16-20. The oxidation conditions of butane and its results are presented in table 10.

The results showed that in the presence of a potassium-doped catalyst containing magnesium molybdate, butane oxidizes predominantly to butadiene and butenes. In addition, it was found that at a constant temperature and volumetric hourly gas flow rate, the volumetric flow rate of butadiene significantly increases with an increase in the percentage of butane in the feed. Under the same conditions, butane conversion and selectivity for butadiene are maintained at a high level. Similar conclusions can also be made regarding the volumetric flow rate and selectivity of the sum of C ₄ products (butadiene and butenes).

Claims (2)

1. The method of producing 1,3-butadiene by oxidative dehydrogenation of butane in the presence of a catalyst comprising magnesium and molybdenum oxides and an alumina-based support at 500 650 ^o C and 7 210 KPa and other conditions providing a reaction mixture containing 1, 3-butadiene with a selectivity of at least 40 mol. characterized in that, in order to increase the productivity of the process, they use a catalyst with reactive oxygen, additionally containing alkali metal oxide and containing a carrier with a surface area of 25.0 184.3 m ² / g, consisting of magnesium and aluminum oxides and / or magnesium aluminum spinel and having a mass ratio of MgO / Al ₂ O _{3 of} 0.38 0.79, with the following content of components, wt. Molybdenum oxide 8.40 24.30
Alkali metal oxide 0.16 1.43
Media Else
2. The method according to claim 1, characterized in that they use a catalyst, optionally containing 0.58 to 0.85 wt. vanadium pentoxide and the process is carried out at 500 650 ^o C, 10.5 210.0 KPa, a volumetric feed rate of butane 400 4000 h ^{- 1} and other conditions providing a reaction mixture containing butadiene with a selectivity of at least 50 mol. 3. A catalyst for producing 1,3-butadiene, comprising magnesium and molybdenum oxides and an alumina-based support, characterized in that, in order to increase the activity and stability of the catalyst, it further comprises an alkali metal oxide and contains a support with a surface area of 25, 0 184.3 m ² / g, consisting of magnesium oxide and aluminum and / or aluminum spinel, having a mass ratio of MgO / Al ₂ O _{3 of} 0.38 0.79, and has the ability to supply active oxygen, with the following content of components, wt. . Molybdenum oxide 8.4 24.3
Alkali metal oxide 0.16 1.43
Media Else
4. The catalyst according to claim 3, characterized in that, in order to increase the activity of the catalyst, it additionally contains 0.58 to 0.85 wt. vanadium pentoxide.

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