Ammoxidation method in fluidized-bed reactor

BACKGROUND OF THE INVENTION

Field of the invention

The present invention relates to a process for carrying out gas phase catalytic ammoxidation in a fluidized bed reactor, and provides an economically advantageous process for producing starting material to be ammoxidized (hereinafter referred to as a hydrocarbon feedstock or the like) of a corresponding nitrile in an improved yield.

State of the art

Ammoxidation in which a hydrocarbon or the like is contacted with ammonia and oxygen in the presence of a catalyst to produce a nitrile in one reaction step is very useful from an industrial standpoint. It is known that when methanol is used as the starting hydrocarbon or the like, hydrocyanic acid is produced (therefore, in the present invention, starting materials to be ammoxidized are not limited to hydrocarbons and nitriles also include HCN); and that when propylene, isobutene or 2,6-dichlorotoluene are used, acrylonitrile, methacrylonitrile or 2,6-dichlorobenzonitrile is produced, respectively. These nitriles are widely used as starting materials for polymeric compounds such as resins and fibers, and for a variety of chemical compounds.

Generally, to conduct the ammoxidation reaction of such a hydrocarbon or the like, the optimum temperature range is narrow, and the amount of exothermic energy generated by the reaction is large. Therefore, a fluidized bed reactor is often used for this ammoxidation reaction because it has excellent temperature controllability, and can handle a high concentration starting gas, and thus

07120«

-2 can achieve high productivity. However, in a conventional fluidized bed reactor, a feed gas to be added to the reactor passes through a catalyst bed as a large number of bubbles, so that the contact between the feed gas and the catalyst tends to be insufficient. On the other hand, a fluidized bed catalyst is characterized in that the mixing of catalyst particles, including the mixing of the catalyst particles in an upstream to that in a downstream, is successful. However, due to this characteristic, the mixing of an upward gas stream with a downward gas stream, back-mixing tends to take place in the catalyst bed.

For the reasons described above, when using a conventional fluidized bed reactor, the yield of nitrile tends to be low, compared to a situation where a fixed bed reactor is used. To solve this problem, which is inherent in a fluidized bed reactor, many proposals have hitherto been made. These proposals can be broadly classified into the following three groups.

Methods classified in the first group are such that molded articles made of wire mesh, screens, grids, perforated plates, horizontal plates, vertical pipes or the like are placed in a catalyst bed as obstacles to prevent bubble coalescence or growth or to promote bubble distribution, or to control the state of mixing of the catalyst particles, thereby improving the contact between the starting gas and the catalyst (the methods described in, for example, Japanese Patent Publications 2533/1965, 28491/1969, description of the description

U.S. Patent 3,783,528, etc.).

531/1973 and American of the In this

38428/1983 (U.S. Pat. No. 4,082,786),

US patent methods seem to indicate the contact between a starting gas and a

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-3 catalytic converter is improved when the obstacles are stacked close. However, such obstacles are not practical since the construction for laying them is complicated. In addition, they greatly prevent the mixing of the catalyst particles, so that the characteristics of the fluidized bed, such as, for example, the excellent controllability of the temperature and the uniformity of the temperature of the catalyst bed, are not fully utilized; and, at the same time, the catalyst in the reactor is not evenly distributed in terms of space and time, making it difficult to conduct the reaction stably and continuously. On the other hand, if the obstacles are placed such that a relatively good controllability of temperature can be obtained, the contact between the reaction gas and the catalyst will not be completely improved.

Methods classified in the second group focus on controlling the distribution of the concentration of a starting gas in a reactor with the aim of increasing the consumption of the starting gas in the reaction. In these methods, a feedstock is fed through two feed openings which are arranged separately at two different locations; or, after a hydrocarbon or the like has been thoroughly mixed with ammonia and oxygen, the mixture is brought into substantial contact with a catalyst. These methods are described, for example, in Japanese Patent Publication 41369/1970 (the description of U.S. Patent 3,546,268), Japanese Patent Laid-Open Nos. 9751/1982, 258/1990 (description of U.S. Patent

4,801,731) and 157355/1991. However, these methods have not been proposed to improve the efficiency of the reaction between a reaction gas and a catalyst, which is a fundamental problem of fluidized beds, but

5 for improving the state of mixing of a starting gas. Although for some time the starting gas has • 1

- If the addition is kept in a state that satisfies the intended purpose, these methods will not solve the problem that a catalyst and a starting hydrocarbon or the like cannot be brought into close contact with each other by the growth or enlargement of bubbles created while the starting gas flow through the catalyst bed, nor will they be able to completely prevent the cause of an uneven distribution of the residence time due to the back-mixing of the gas. Thus, these methods are insufficient as a way of improving the contact between a gas and catalyst particles

Methods classified in the third group use a fluidized state that is substantially different from that in a conventional fluidized bed. These methods are such that a fluidized bed is formed by a large number of catalyst particles transported by, and accompanied by, a high velocity gas stream. The solids density of this fluidized bed is relatively low compared to a conventional fluidized bed, and the flow of the gas and of the catalyst are comparable to a piston flow. By the way, a fluidized bed in such a state was referred to as fast fluidization by Joseph Yerushalmi et al. (Industrial and Engineering Chemistry Process Design Development, vol. 15, No 1, pp. 47-53 (1976)).

As a technique using this high speed fluidized bed, a method is known which is described in Japanese Patent Laid-Open No. 144528/1978 (description of U.S. Patent 4,102,914). This method is characterized in that the reaction is carried out by a density of the solid of about 16 to 240 kg / m³ and a gas velocity of about 1.5 to 7.5 m / s, wherein the density of the solid is low and the gas velocity is high compared to that in a conventional

-5 fluidized bed. However, although the process is advantageous because the productivity per cross-section of the reactor is high due to the high gas velocity, the following problem seems to arise: to obtain the desired nitrile in high yield requires an extremely long reaction zone; to achieve this, it is necessary to significantly increase the height of the reactor, when the reactor is in the form of a vertical cylinder, which is a common form for conventional reactors, so that the construction costs increase. To circumvent this problem, the reactor is formed in a spiral-like shape.

However, in a process using this spiral reactor, since centrifugal forces act on the catalyst particles flowing through the spiral, gas and catalyst particles are distributed unevenly, and contact between them is poor. Thus, the effect of improving the contact between the gas and the catalyst particles appears to be insufficient. Another problem associated with this process is as follows: If the catalyst is not separated and removed, in a large amount, from the reaction product immediately after the conversion of the starting hydrocarbon or the like has reached an optimum, the reaction proceeds excessively, and a nitrile produced in a lower yield. Furthermore, in this process, a cyclone is usually used as a catalyst separator and, since the amount of the gas and of the catalyst are large, the cost of manufacturing a cyclone capable of meeting these requirements, i.e., a cyclone which is large in size, has a high collection capacity and hardly causes any wear or pulverization of the catalyst. In addition, in this method, it is generally necessary to lay a pipe useful for transporting the catalyst outside the reactor system to return the separated catalyst back into the reactor. In addition, it is necessary to use the

-6 pressure balance so that the reactant gas does not flow back into this pipe.

Thus, the method of improving the contact between a gas and catalyst particles by using high-speed fluidization has such problems that an increased cost is required for the construction of the plant, that the operation is complicated, and that the effect of improving the contact between a gas and catalyst particles cannot be sufficiently achieved depending on the type of plant of the plant. Therefore, the expected effects would not be fully obtained.

SUMMARY OF THE INVENTION The present invention has been accomplished to solve the aforementioned problems in the conventional techniques of ammoxidizing a starting hydrocarbon or the like in a fluidized bed reactor. Thus, an object of the present invention is to provide a method of economically producing a desired nitrile in an increased yield by improving the contact between the starting gas and the catalyst particles without adding any special equipment to a conventional fluidized bed reactor.

The present invention has been accomplished on the basis of the following: high gas velocity and relatively low solid concentration are not essential requirements for achieving close contact between a gas and catalyst particles, which is a characteristic feature of a fluidized bed for high velocity fluidization, and the mode of contact between a gas and catalyst particles can be extremely good even at a low gas velocity, by good control of the solid concentration in the reaction zone.

Namely, an ammoxidation method according to the present invention, which is carried out in a fluidized bed reactor, is characterized in that, when a too η o 71 2 η · * ι

Amoxidizing starting material is ammoxidized by means of a catalytic fluidized bed reaction in the vapor phase, the reaction is carried out in a fluidized bed reactor to which an oxygen-containing gas is supplied through feed openings provided at the bottom thereof, and an ammoxidizing starting material is supplied through supply openings provided above the supply openings for the oxygen-containing gas, the distance between the feed openings for the oxygen-containing gas and that for the starting material being between 30 and 250% of the height of a fluidized solid in a static state such that a fluidized bed is created such that the density of the fluidized solid (ie the density of the catalyst bed) at the feed openings for the starting material to be ammoxidized in the range from 50 to 300 kg / m³ and the gas velocity at the supply openings for these is 1 m / s or less.

According to the present invention, the efficiency of the contact between a starting material and catalyst particles is extremely high, and as a result, an extremely improved result (the yield of the desired product) from the reaction can be obtained.

BRIEF DESCRIPTION OF THE DRAWINGS

In the drawings is

Fig. 1 a cross-sectional view of an example of a reactor used to perform a generalized ammoxidation in accordance with the present invention, wherein 1: a reactor, 2: a catalyst, 3: an oxygen gas feed pipe, 4, 7: a gas dispersant, 5.8: openings for the gas dispersant, 6: feed tube for a mixture of starting hydrocarbon or the like and ammonia, 9, 10: a cyclone separator, 11, 12: a catalyst return tube, and 13: a discharge pipe; and

Fig. 2. a sectional view of a reactor used to perform the ammoxidation of

7120^

-8methanol according to the method of the present invention, wherein 6A: feed pipe for a gas mixture (methanol + ammonia), 7A: a gas dispersant, 8A: openings for the gas dispersant, 6B a steam supply line, 7B: a gas dispersant and 8B: openings for the gas dispersant; the other parts are indicated by like symbols as used in Fig. 1

DETAILED DESCRIPTION OF THE INVENTION

The ammoxidation method of the present invention, which is carried out in a fluidized bed reactor, is a method as defined in the appended claims. The method can be any targeted as long as it fulfills the required conditions or composition as defined in the claims.

<Reaction in a Fluidized Bed>

A reactor as shown in Fig. 1 can be mentioned as a preferred embodiment of the present invention. As described above, FIG. 1 illustrating only one embodiment of the present invention. It is therefore not necessary to mention that the present invention is not limited by the details specifically indicated in FIG. 1.

A catalyst 2 is placed in a reactor 1. An oxygen-containing gas is introduced into reactor 1 through a pipe 3, and is supplied to the lower part of the catalyst bed through openings, i.e. supply openings 5 ​​for an oxygen-containing gas, formed in a gas dispersant or a gas distributor 4. A starting hydrocarbon or the like and ammonia are introduced into the reactor 1 as a mixture thereof (or separately) through a pipe 6, and fed (in this embodiment and preferably downwardly) to the catalyst bed through openings 8, i.e. feed openings for a starting hydrocarbon or the like, formed in a gas dispersing device 7. The feed openings 8 for the

<img file="NL1007120C2_D0001.tif" />

-9 starting hydrocarbon or the like are provided above the oxygen-containing gas supply openings, the distance between the oxygen-containing gas supply openings and that of the starting hydrocarbon or the like being

30 up to 250%, preferably from 50 to 200%, of the height of the catalyst bed in a static state, i.e. the height of the stationary catalyst bed. At the top of reactor 1 is a multiple cyclone composed of cyclones 9 and 10 connected in series. The catalyst particles in the gas stream are separated therefrom by this multiple cyclone, and recycled through a pipe 11 or 12 to the lower or middle portion of the catalyst bed. As for the catalyst particles recovered by the multicyclone, it is preferable to return the particles having a smaller particle diameter to the lower part of the catalyst bed. The gas from which the catalyst particles have been removed is passed through the last cyclone and expelled from the reactor through a pipe 13

0 1. The oxygen-containing gas supplied to the bottom of reactor 1 is generally air. However, oxygen-enriched air, or oxygen diluted with an inert gas, can also be used in place of air.

As shown above, the oxygen-containing gas supply ports are means for supplying the oxygen-containing gas to the fluidized catalyst bed, and the location of or, when the reactor is vertical, the level or height of the oxygen-containing gas supply ports may typically determine the bottom of the fluidized catalyst bed.

The starting hydrocarbon or the like for use in the present invention is methanol, propylene, isobutene, methylbenzene, especially toluene or 2,6-dichlorotoluene, or any other aliphatic hydrocarbon. Hydrocyanic acid is produced from methanol, acrylonitrile is produced

<img file="NL1007120C2_D0002.tif" />

-10 from propylene; and methacrylonitrile is produced from isobutene. All of these compounds are industrially useful chemicals; they are in great demand as starting materials for various derivatives.

It is also possible to supply the above-described starting hydrocarbon to the reactor 1 mixed with ammonia and / or an inert gas or steam. Regarding the supply of steam, reference can be made to the applicable description in &lt;Ammoxidation of Methanol&gt;

as shown below.

In the reactor 1 described above, the area between the feed openings 8 for the starting hydrocarbon or the like and the inlet of the cyclone 9 located at the top of the reactor 1 is a main reaction zone. The density of the solid in this reaction zone is from 1 to 300 kg / m³. More specifically, it is important that the density of the fluidized solid at the feed openings for the starting hydrocarbon or the like is in the range of 50 to 300 kg / m³, preferably 100 to 250 kg / m³. In order to obtain such a solid density, the catalyst particles are generally placed in a high speed fluidization state using a high speed gas. However, in the present invention, an oxygen-containing gas is supplied to the bottom of the catalyst bed as previously mentioned, and hydrocarbon feed ports 8 to be supplied above the oxygen-containing gas are provided above the oxygen-containing gas supply ports 5, provided that the distance between the supply ports 5 for the oxygen-containing gas and the feed openings 8 for the hydrocarbon are from 30 to 250% of the height of the stationary catalyst bed. This allows the degree of contact between the gas and the catalyst to be kept excellent even when the gas velocity is as low or less than 1 m / s.

The density of the fluidized solid, i.e., the density of the catalyst bed, can be determined

-11 by measuring the difference between static pressures at two different points on the fluidized bed in the direction of its height, and dividing the difference by the distance between the two points.

Therefore, the density of the fluidized solid at the feed openings for the starting hydrocarbon or the like can be estimated from the distribution of the density in the direction of the height of the fluidized bed.

The optimum position of the feed of the starting hydrocarbon or the like varies within the range described above depending on the physical properties of the catalyst, such as particle diameter, density, etc., the catalytic activity and the gas velocity. When the feed location of the starting hydrocarbon or the like is less than 30% of the height of the stationary catalyst bed, the density of the solid may be excessively high. Therefore, it may not be easy to keep the density at the feed openings for the starting hydrocarbon or the like at about 50 to 300 kg / m³ unless the gas speed is made excessively high. On the other hand, when the feed hydrocarbon feed location or the like is greater than 250% of the stationary catalyst bed height, the density of the solid at the feed hydrocarbon feed openings or the like may be about 50 to 300 kg / m³ amounts. In this case, however, the length of the reaction zone is not sufficiently long, so that it may be impossible to make the conversion of the hydrocarbon or the like sufficiently high.

&lt;Catalyst&gt;

Preferably, the catalyst for use in the present invention is a very active one. However, there is no specific limitation on the catalyst, and any known ammoxidation catalyst can be used.

-12 In the present invention, catalysts such as, for example, those containing iron and antimony, and at least one element selected from the group consisting of boron, magnesium, phosphorus, potassium, titanium, vanadium, chromium, manganese, cobalt, nickel, copper can be used , zinc, molybdenum, tin, tellurium, tungsten, and bismuth. And preferably, a catalyst represented by the following empirical formula (A) can be used:

Fe_aSb_bQ_cR_dO_e (A) wherein Q represents at least one element selected from the group consisting of V, Co, Ni, Cu, Mo, W and Bi; R represents at least one element selected from the group consisting of B, P, K, Zn, and Te; and the indices a, b, c, d, and e represent an atomic composition, wherein if a = 1, b, c and d 0.05 &lt;b &lt;10; 0.001 &lt;c &lt;10; and 0 &lt;d &lt;5, and e represents the number of oxygen atoms corresponding to the oxide produced when the above components are combined.

As the ammoxidation catalyst of this type, many catalysts have conventionally been proposed. For example, a metal oxide comprising molybdenum and bismuth, and a metal oxide comprising antimony and tin or iron or uranium, such as those described in Japanese Patent Publications 5870/1961, 13460/1962 and 14075/1962, Japanese Laid-Open Patent Publications 58100/1974 (de description of U.S. Patent 3,911,089) and 10200/1976, Japanese Patent Publication 33888/1976 (the description of British Patent 1,319,190) and 18014/1978 (U.S. Patent 3,988,359 and 4,083,804), and Japanese Patent Laid-Open No. 257125/1989 (U.S. Patent 5,158,787) are known for their effectiveness in the ammoxidation of methanol, propylene or isobutylene.

These catalysts can be easily obtained according to the details described below with respect to &lt;Ammoxidation of Methanol&gt;.

\, U71 2Q- *

-13 The fluidized solid contained in the fluidized bed in the present invention is composed of the particles of such an ammoxidation catalyst and, if necessary, other solid particles such as small-area silica or alumina particles, or catalyst particles deactivated by high temperature calcination.

In the present invention, the term density of the fluidized solid means the density of the solid including the particles described above which have an extremely low chemical and catalytic activity on the reaction compared to that of the catalyst, or in the have no such activity at all.

In the practice of the present invention, the range of the particle size of the catalyst particles is in the range of about 10 to 500 microns, and the average particle diameter is from 30 to 200 microns, preferably from 40 to 100 microns. The density of the

0 mass of the catalyst is in the range of 0.5 to 2 g / cm³, preferably from 0.7 to 1.5 g / cm³. The height of the stationary catalyst bed is in the range of 0.1 to 10 m, preferably 1 to 5 m. The starting gas to be fed to the reaction column is such that the molar ratio of oxygen / hydrocarbon or the like is in the range from 0.5 to 5, preferably from 1 to 3, and that the molar ratio of ammonia / hydrocarbon or the like is in the range from 0.5 to 3, preferably from 0.7 to 1.5. The gas velocity is in the range of 0.1 to 1 m / s, preferably 0.3 to 1 m / s; the reaction temperature is in the range from 350 to 500 ° C; the contact time is in the range of 0.1 to 30 seconds, preferably 0.5 to 20 seconds; and the reaction pressure is in the range of atmospheric pressure to 2 kg / cm^zG.

&lt;Ammoxidation of Methanol&gt;

-1410

Intended compounds for ammoxidation performed by the method of the present invention are a wide variety of compounds including methanol. These target compounds are all industrially useful when taking into account the use of their amoxidized products, namely nitriles. However, the production of hydrocyanic acid by ammoxidation of methanol is especially important because the resulting nitrile has a relatively simple structure and can further undergo various reactions.

The conditions of this particular ammoxidation are the same as those of the ammoxidation described above. However, an explanation that only focuses on this specific ammoxidation is as follows.

The present invention also particularly relates to a process for carrying out catalytic ammoxidation of vapor phase methanol in a fluidized bed reactor. More specifically, the present invention is intended to provide an industrially advantageous method for ammoxidizing methanol, wherein the catalytic activity is kept high to obtain a high yield of hydrocyanic acid.

Hydrocyanic acid is produced by the decomposition of formamide, by the reaction between methane and ammonia, by the ammoxidation reaction of methane, or the like. Furthermore, hydrocyanic acid, in most cases, is obtained as a by-product when acrylonitrile is produced by the ammoxidation of propene. In recent years, however, the use of hydrocyanic acid derivatives such as acetocyanohydrin and adiponitrile has expanded, and the demand for sodium cyanide used for gold recovery has increased. This can lead to a shortage of hydrocyanic acid.

In view of these conditions, the present invention contemplates a new, industrially advantageous

1007 ' ^f’ ·

-1510 to develop a method for the production of hydrocyanic acid by the ammoxidation of methanol.

As a process for the production of hydrocyanic acid by the ammoxidation of methanol, there is known a process using a molybdenum oxide catalyst as described in Russian patent 106,226, a process using an oxide catalyst containing molybdenum, bismuth and other elements as described in Japanese Patent Publication 35400/1976, a method using an oxide catalyst containing antimony and at least one of the elements selected from the group consisting of iron, cobalt, nickel, manganese, zinc and uranium as described in Japanese Patent Publication 39839/1979, a method using of an oxide catalyst containing manganese and phosphorus as described in the description of U.S. Patent 4,457,905, a method using an antimony phosphate catalyst as described in the description of U.S. Patent 4,511,548, and the like.

Furthermore, in view of improving the iron antimony oxide catalyst described in Japanese Patent Publication 39839/1979, we have proposed an oxide catalyst containing iron, copper and antimony as described in Japanese Patent Laid-Open Publication 145617/1983, an oxide catalyst, containing iron, copper, antimony and phosphorus as described in Japanese Patent Publication 64555/1995 and a catalyst containing iron, antimony and phosphorus as essential components and iron antimonates as a crystalline phase, as described in Japanese Patent Laid-open Publication 26342/1991. In addition, we also have methods of using as those methods of preparing those methods described in Japanese Patent Publications 12434/1995 and 63629/1995 and the like. Improvements have been proposed by these methods, catalyst,

-16 on many points. However, for the industrial application of these methods, it would be preferable to further improve some points.

When hydrocyanic acid is produced by the ammoxidation of methanol, in a conventional process it may be necessary to make the ratio of ammonia or oxygen to methanol in a feed gas high to maintain the catalytic activity. However, even when such a reaction condition is applied, the problems of economic efficiency and safety remain unresolved. Generally, the following method is employed to recover hydrocyanic acid from a reaction product: a reaction product with a high temperature is cooled, then absorbed by a solvent such as water and the absorbed liquid is subjected to distillation to remove the hydrocyanic acid therefrom divorce. During this operation hydrocyanic acid can polymerize when ammonia is present in the system. This is not only a loss of hydrocyanic acid, but also the cause of blockage of the used device. It is therefore necessary to separate ammonia and treat the cooling phase, neutralization. However, when ammonia is supplied in excess to methanol, the amount of ammonia used per amount of HCN produced increases, the amount of acid needed to neutralize the unreacted ammonia increases, and costs are incurred for the construction of a plant to remove the 30 salts are formed during the neutralization treatment. This usual method is therefore advantageous from an economic point of view.

Furthermore, when the ratio of oxygen to methanol is high, a large amount of nitrogen, which is not required in the reaction, will be supplied to the reaction system, because air is generally used, for example, by removing it with an acid in reaction product

-17 as the source of oxygen, with the result that the volumetric efficiency of the reactor is lowered. In addition, when the concentration of oxygen in the reaction product is increased, the reaction product can enter the explosion area, so that an explosion hazard is present.

For the reasons described above, it is desirable that the ratio of ammonia or oxygen to methanol in a feed gas be low. However, when this ratio is low, hydrocyanic acid is produced in a reduced yield, and the catalyst may deteriorate in catalytic activity over time.

The present invention aims to overcome the aforementioned shortcomings of the prior art.

Therefore, an object of the present invention is to provide an economically advantageous process in which the reaction can be effectively conducted even under such a reaction condition that the ratio of ammonia or oxygen to methanol in the feed gas is low, and the desired product, Knowing hydrocyanic acid, can be obtained in a high yield with a high selectivity and stable over time.

We have studied industrial methods of hydrocyanic acid production by conducting vapor phase catalytic ammoxidations of methanol in a fluidized bed reactor and as a result found the following: when methanol, ammonia and steam are fed to the reaction system from specific positions, respectively, an oxide catalyst containing specific elements such as iron, antimony, phosphorus and copper can improve the efficiency of the use of ammonia and oxygen; the oxidation reduction stability of the catalyst has remarkably increased, so that the catalyst hardly undergoes deterioration; and hydrocyanic acid can be obtained in high yield even when the reaction is carried out under conditions such that the ratio of ammonia or

1007120”

-Ιδιο oxygen until methanol is low. The present invention has been accomplished based on these findings.

Accordingly, the present invention relates to a process for the ammoxidation of methanol, wherein when the catalytic ammoxidation of methanol in the vapor phase is conducted in a fluidized bed reactor under the aforementioned general conditions for ammoxidation to obtain hydrocyanic acid, the reaction is carried out using, as a catalyst, a metal oxide containing [I] at least one element selected from the group consisting of iron, antimony, phosphorus and copper, and [II] at least one element selected from the group consisting of vanadium, manganese, molybdenum, tungsten and bismuth, by supplying, to the catalyst bed, an oxygen-containing gas through supply openings located in the lower portion of the reactor, and methanol and ammonia as a steam-containing gas through supply ports positioned above the supply ports for the oxygen-containing gas, or by supplying all or part of the steam to the catalyst bed through supply ports located below the supply ports for methanol and ammonia, with the mole ratio of steam to methanol controlled at 0.1 to 3.

An example of the present invention focused on the ammoxidation of methanol will now be discussed. The description as given below of this one ammoxidation refers, whenever applicable, to the ammoxidation as set forth above in general terms.

In the process of the present invention, it is important to conduct the ammoxidation reaction of methanol by using a metal oxide catalyst containing specific elements such as iron, antimony, phosphorus and copper, and by feeding methanol, ammonia and steam to a specific reaction zone.

&lt;Catalyst (bis)&gt;

-1910

The catalyst for use in the method of the present invention is a metal oxide composition containing [I] at least one element selected from the group consisting of iron, antimony, phosphorus and copper, and [II] at least one element selected from the group consisting of from vanadium, manganese, molybdenum, tungsten and bismuth. The atomic ratio of the element [I]: the element [II] is preferably in the range of (20 to 100) :( 0.5 to 5). This catalyst gives excellent performance when steam is present in the reaction system and when the density of the catalyst bed is low, so that hydrocyanic acid can be produced with a high production efficiency. A preferred catalyst is an oxide composition comprising iron, antimony, phosphorus and vanadium. The vanadium content in this oxide composition is at least 0.6, preferably from 0.6 to 3, for 10 iron, when expressed in atomic ratio.

More specifically, preferred catalysts are oxide compositions represented by the following empirical formula (B):

Fe_a, Sb_b, P_c.V_d, Mo_e, Cu_£, W_g,X_h, Y_i, Z_j,O_k(SiO1), (B) wherein Fe, Sb, P, V, Mo, Cu and W represent iron, antimony, phosphorus, vanadium, molybdenum, copper and tungsten, respectively; X represents at least one element selected from the group consisting of Mg, Zn, La, Ce, Al, Cr, Mn, Co, Ni, Bi, U and Sn, preferably at least one element selected from the group consisting of Mg, Zn , Al, Mn, Co, and Ni; Y represents at least one element selected from the group consisting of B and Te; Z represents at least one element selected from the group consisting of Li, Na, K, Rb, Cs, Ca and

<td>Ba;</td><td>the indexes</td><td>a ', b', c ', d', e '</td><td>, f ', g', h ', i'</td><td>, j '</td><td>, k '</td><td>and</td>

<td> 1 '</td><td colspan="3">represent an atomic relationship, in which,</td><td colspan="2">when</td><td>a'</td>

<td>is</td><td>10, b 'from</td><td>12 to 30 is, at</td><td>preference of 15</td><td>until</td><td> 27,</td><td>c '</td>

<td>from</td><td>1 to 30</td><td>preferably</td><td>from 3 to 20,</td><td>Bee</td><td colspan="2">higher</td>

preferably from 5 to 15, provided that b '/ C is greater than 1.5, d' is from 0.6 to 3, preferably from 0.8 to 2.8, at

Higher preference from 1 to 2.5, e 'is from 0 to 0.3, f' is from 0 to 5, g 'is from 0 to 3, h' is from 0 to 6, i 'is from 0 to 5 , j 'from 0 to 3, k' is a number corresponding to the oxide produced when the above components are combined, and 1 'is from 0 to 200.

The catalyst represented by the above empirical formula (B) contains as its main components iron, antimony, phosphorus and vanadium. It is not clear how these components form compounds in the catalyst that give effects on activity and physical properties. However, when the composition of the above catalyst exceeds the range defined by the above empirical formula, the selectivity to hydrocyanic acid, or the properties of the catalyst are weakened, with the result that it sometimes becomes difficult to achieve the target. It is therefore to be considered that the elements constituting the catalyst are closely interrelated to show the effects. More particularly, it is preferred to have iron antimonate present in the catalyst to form a crystalline phase. The vanadium component is believed to be dissolved in this crystalline phase in a solid state. The presence of iron antimonate is effective in increasing the yield of hydrocyanic acid, in preventing yield reduction in operation over time, and in imparting appropriate properties to the catalyst.

In the ammoxidation reaction of propylene, it is known that the addition of the vanadium, molybdenum or tungsten component to a catalyst containing iron and antimony is effective in increasing the reaction rate and in increasing the resistance to deterioration by reduction. In particular, the vanadium component has an excellent effect on a catalyst containing iron, antimony and phosphorus. Actually, when a specific amount

100 7'i 2 04

-21 of the vanadium component was added to the catalyst, the resulting catalyst exhibited such a remarkable effect that hydrocyanic acid was obtained in high yield with high selectivity and also stable over time, even when the ratio of oxygen to methanol in a starting gas increased feed to the reaction system was made low, that is, the concentration of starting methanol was made high. The molybdenum and tungsten components do not have such a remarkable effect. This is a fact that should not be expected, given the general knowledge in the art.

The addition of a copper component and the X component is effective in preventing the formation of bulges on the surface of the catalyst, which tend to occur especially when the antimony content is high, to improve strength of the catalyst, and for controlling the rate of reaction and catalytic properties. The addition of the Y component helps to improve selectivity; and the addition of the Z component contributes to the control of the reaction rate and the by-product.

The catalyst for use in the present invention can be used readily, without the use of a carrier. However, it is preferable to use the catalyst with a silica support as a support. The amount of carrier can be freely changed within a range of 10 to 90% by weight relative to the total weight of the catalyst.

The above-described catalyst of formula (B) can be prepared by any known method. For example, a method as described in Japanese Patent Publication 12434/1995 or 63629/1995, in Japanese Laid-Open Patent Publication 26342/1991, or the like may be used.

-22 As described above with respect to the catalysts of formula (A), the following also applies to the production of the catalysts of formula (A) where applicable.

A starting material for each component of the catalyst can be selected from the various types of compounds such as metal oxides, oxides, hydroxides, chlorides and nitrates of the component. Furthermore, those compounds which can be converted into oxides by a chemical treatment or a calcination treatment can also be used.

As the starting material for the iron component, an iron oxide such as ferrous oxide, ferric oxide or tri iron tetraoxide, an iron salt of an inorganic acid such as ferric chloride, ferric chloride, ferric nitrate, iron carbonate or a product obtained by oxidizing metallic iron with nitric acid or an organic salt of iron such as iron oxalate or iron citrate.

As the starting material for the antimony component, antimony trioxide, antimony tetroxide, antimony pentoxide, antimony acid, polyantimonic acid, sodium antimonate, potassium antimonate, antimony trichloride, antimony pentachloride or the like can be used. Furthermore, a product can also be used, which can be obtained by oxidizing metallic antimony with nitric acid.

As the starting material for the phosphorus component, it is preferred to use phosphorus pentoxide, orthophosphoric acid, ammonium dihydrogen phosphate, diammonium hydrogen phosphate, triammonium phosphate or the like.

As the starting material for the vanadium component, vanadium pentoxide, ammonium metavanadate, vanadyl oxalate, vanadyl sulfate or the like can be used.

As the starting material for the copper component, cupride, cuprous oxide, copper nitrate or the like can be used.

10071 20-,

As a starting material for the molybdenum component, molybdenum trioxide, molybdenumic acid, ammonium paramolybdate, ammonium metamolybdate, a molybdenum halogen or the like can be used. As the starting material for the tungsten component, tungsten trioxide, ammonium para-tungstate, ammonium meta-tungstate, tungsten acid or the like can be used.

As the starting material for the tin component, stannous oxide, stannous oxide, or the like can be used.

As the starting material for the X, Y and Z components, oxides, hydroxides, sulfates, carbonates, organic acid salts or the like of these elements can be used, respectively.

It is preferable to use silica sol as the material for the silica support. However, silica hydrogel, fumed silica or the like can also be used as part or all of the silica sol

The fluidized bed catalyst can be made in the following manner: a mixed slurry of materials containing the components of the catalyst is prepared; the pH of the prepared slurry is adjusted (the pH is adjusted to about 7 or less, preferably about 1 to 4), a heat treatment (heating to approximately 40 to 150 ° C), or the like is performed if necessary; then the slurry is spray dried, and the resulting fine particles are subjected to a calcination.

Calcination is important for imparting a predetermined activity to the catalyst. It is preferable to carry out the calcination in a temperature range from 200 to 900 ° C, preferably from 400 to 850 ° C, for 0.5 to 10 hours. There is no particular limitation on the atmosphere in which the calcination is performed, although a reducing atmosphere is not recommended, and the calcination can be carried out or in a

-24 atmosphere of an oxygen-containing oxidizing gas or in an atmosphere of an inert gas containing only nitrogen. Generally, for economic reasons, it is preferable to conduct the calcination in an atmosphere of air. A tube furnace, rotary furnace or a swirl furnace can be used for the calcination.

The size of particles can be chosen depending on the purpose. The fluidized bed catalyst is generally used with particles having a particle diameter in a range from 10 to 200 microns.

&lt;Reaction in a Fluidized Bed (bis)&gt;

A preferred embodiment of this specific method of the present invention is, for example, as shown in Fig. 2. The device as shown in Fig. 2 is the same as shown in Fig. 1 except the means for supplying methanol and ammonia (6A, 7A and 8A), and the means for introducing steam (6B, 7B and 8B); like numbers in these figures identify corresponding parts. The number 1 indicates the reactor, the number 8A 2 0 indicates the supply opening for a mixed gas of methanol + ammonia or the like, the number 8B indicates supply openings for steam, the number 4 indicates supply openings for an oxygen-containing gas and the figure 2 indicates the catalyst. This specific example 25 of the present invention regarding the ammoxidation of methanol is not limited to only the scope specifically shown in FIG. 2.

A catalyst 2 is placed in a reactor 1. An oxygen-containing gas is introduced into the reactor 1 through a pipe 3 and supplied to the lower part of the catalyst bed through openings, namely supply openings 4 for an oxygen-containing gas, formed in a gas dispersing device 5. Methanol and ammonia are mixed with steam and the resulting mixture is introduced into the reactor 1 through a pipe 6A provided above the oxygen-containing gas supply ports 4, and

-25 fed to the catalyst bed through openings, namely supply openings 8A for a mixed gas of methanol + ammonia, etc., formed in the gas disperser 7A. It is possible, if necessary, to introduce at least a part of the steam into the reactor 1 through a pipe 6B, which is provided under the supply openings 8A for methanol and ammonia in order to supply the steam in a divided manner in this way. catalyst bed through the supply openings, namely supply openings 8B for water vapor, formed in a gas disperser 7B. An alternative is to mix at least a portion of the steam with the oxygen-containing gas, and to supply the mixture to the catalyst bed by a gas dispersant 5. At the top of the reactor 1 is a multicyclone composed of cyclones 9 and 10 connected in series. The catalyst particles accompanied by the gas stream are separated therefrom by this multicyclone and recycled to the lower or middle portion of the catalyst bed through a pipe 12 or 11. The oxygen-containing gas supplied to the bottom portion of the reactor 1 is, generally, sky. However, oxygen- enriched air or oxygen diluted with an inert gas can also be used in place of air.

Again, the oxygen-containing gas supply ports are means for supplying the oxygen-containing gas to the catalyst bed, which is fluidized, and the location of or, when the reactor is vertical, the level or height of the gas supply ports typically determine the bottom of the fluidized catalyst bed.

In the process of the present invention of the ammoxidation of methanol, it is preferable to control the amount of steam to be supplied to the catalyst bed, such that the steam to methanol molar ratio will be from 0.1 to 3. Furthermore, it is preferred the method of the present invention of this

-26 perform specific embodiment by setting the molar ratio of oxygen to methanol in the feed gas to less than 1.6, preferably from 0.8 to 1.5, and the molar ratio of ammonia to methanol in the feed gas to less than 1.2, preferably 0.7 to 1.1. The concentration of methanol in the feed gas can be changed within a range of 3 to 20%.

Also in this particular embodiment of the present invention, the position of the mixed gas supply ports such as a mixture of methanol and ammonia cannot be determined solely by the shape, size and the like of the reactor. However, if the reactor is cylindrical, they may be provided above the oxygen-containing gas supply ports provided that the distance between the oxygen-containing gas supply ports and that for a mixed gas is between 30 and 250% of the height of the fluidized catalyst bed in static condition, namely the height of the stationary catalyst bed. Steam can be supplied after mixing with methanol and ammonia.

Alternatively, all or part of the steam may be supplied to the catalyst bed from a position below the mixed gas supply ports such as a mixture of methanol and ammonia. The divided supply of steam can be achieved by allowing steam to be present in an oxygen-containing gas, or by supplying steam to the catalyst bed through supply openings separately provided for supplying steam. The amount of steam that can be fed in part is from 5 to 95%, preferably from 20 to 90% of the entire amount of steam that is supplied to the reactor. Methanol and ammonia can be fed by mixing the entire amount of methanol with that of ammonia. Alternatively, they can be fed in aliquots.

In this particular example of the present invention, when the reaction is conducted using the catalyst under such a condition

Since the density of the catalyst bed is low, the degree of contact between the gas and the catalyst can be made good so that hydrocyanic acid can be obtained in high yields. The density of the catalyst bed at the position of the feed openings for methanol and ammonia is from 50 to 300 kg / m³, preferably from 100 to 250 kg / m³. Generally, a high gas velocity is used to obtain a density of the catalyst in the above-described range. However, in the process of the present invention, the degree of contact between the gas and the catalyst can be kept excellent even when the gas velocity is kept at 1 m / sec or less. Namely, it is important to decide on the position of the feed ports for methanol and ammonia, and to choose the steam split ratio so that the catalyst bed density described above can be obtained under the given reaction conditions. The density of the catalyst bed can be determined by determining the difference between static pressures measured at two different points on the fluidized bed in the direction of its height, dividing the difference by the distance between the two points. Therefore, the density of the catalyst bed at the methanol and ammonia feed ports can be estimated from the distribution of the density in the direction of the height of the fluidized bed.

By adding steam to the catalyst bed, the efficiency of the use of ammonia and oxygen is improved. In particular, the oxidation reduction stability of the catalyst is remarkably improved so that the catalyst hardly undergoes deterioration. In addition, the hydrocyanic acid yield is not reduced even when the ammonia or oxygen ratio is made low. Thus, this particular method of the present invention is advantageous from an industrial point of view.

-28 The reaction temperature for the specific ammoxidation is in the range between 350 to 500 ° C, preferably 380 to 470 ° C. The reaction can be carried out under conditions of atmospheric, superatmospheric or subatmospheric pressure. It is convenient, however, for the reaction pressure to be between a pressure near atmospheric pressure and a pressure of 2 kg / cm², Gauge.

The contact time is in the range 0.01 to 20 seconds, preferably 0.05 to 10 seconds, especially 0.1 to 6 seconds, based on the gas volume at the reaction temperature below the reaction pressure.

Examples

The following Examples serve to more specifically illustrate the composition and effects of the present invention.

The height of a stationary catalyst bed, the relative height H of the feed openings for a starting hydrocarbon or the like, the gas velocity, the yield of nitrile and the conversion of a starting hydrocarbon or the like used in this Description are defined as follows.

Catalyst bed height [m] = (amount of packed catalyst [kg] / {(mass density of catalyst [kg / m³]) x (effective diameter of reactor [m²] ) }

H [%] = {(distance between supply ports for oxygen containing gas and those for starting hydrocarbon or the like [m]) / (height of stationary catalyst bed [m])} x 100

Gas velocity [m / s] = (volume velocity of whole feed gas under reaction conditions [m³/ s] / (effective diameter of reactor [m²] )

Nitrile yield [%]

<img file="NL1007120C2_D0003.tif" />

-29 = {(weight of carbon in nitrile produced [kg] / (weight of carbon in feed hydrocarbon feed or the like [kg])} x 100

Conversion of starting hydrocarbon or the like [%] = {(weight of carbon converted into starting hydrocarbon or the like [kg]) / (weight of carbon fed into starting hydrocarbon or the like [kg])} x 100 &lt;Comparative Example Al&gt;

Hydrocyanic acid was synthesized by carrying out the ammoxidation reaction of methanol in a fluidized bed reactor using an iron-antimony-phosphorus copper vanadium-molybdenum fluidized bed catalyst supported on a silica support. A reactor with an internal diameter of about 0.2 m and a height of 5 m was used for this reaction. Air was supplied to the bottom of this reactor, and a mixed gas of methanol and ammonia was supplied from the position 0.2 m above the air supply openings. A cyclone is provided at the top of the reactor and those catalyst particles that come out of the catalyst bed and are transported or accompanied by the gas are collected by this cyclone. These catalyst particles are then returned to the location 0.15 m above the air supply openings, through a vertical pipe with an internal diameter of 0.037 m, attached to the lower part of the cyclone so that they are returned to the catalyst bed.

kg of the catalyst was packed into this reactor. At this time, the stationary catalyst bed height was 1.04 m, the relative height H of the methanol feed openings was 19%, and the catalyst bed density at height H was 510 kg / m³.

By controlling the composition of the gases to the reactor so that the oxygen / methanol molar ratio is 1.36, and the ammonia / methanol molar ratio is 1.0, and by controlling the

<img file="NL1007120C2_D0004.tif" />

-30 gas velocity at 0.5 m / s, the reaction was carried out at a temperature of 43 0 ° C under a pressure of 0.5 kg / cm²G. The results of the reaction were such that the hydrocyanic acid yield was 82.0% and the conversion of methanol was 99.9%.

&lt;Example Al&gt;

Hydrocyanic acid was synthesized by conducting the ammoxidation reaction of methanol using the same reactor and catalyst used in

Comparative Example A1. The reaction was performed in the same manner as in Comparative Example A1, except that the relative height H of the methanol feed openings was changed to 108% and the reaction temperature was changed to 440 ° C. To the height

H to 108%, the methanol feed openings were provided at the position 0.7 m above the air feed openings located at the bottom of the reactor. Furthermore, the catalyst was packed in the reactor in an amount of 25 kg to bring the height of the stationary catalyst bed to 0.65 m. At this time, the density of the catalyst bed at height H was 200 kg / m³.

The results of the reaction were such that the hydrocyanic acid yield was 92.0% and the conversion of methanol was 98.5%.

Thus, it was found that the yield of hydrocyanic acid obtained in this example was considerably higher than that obtained in Comparative Example A1. It is noted that the reason why the reaction temperature was raised to a temperature 10 ° C higher than in Comparative

Example A1 is to ensure sufficient conversion of methanol. However, a 10 ° C difference in reaction temperature only slightly affects the hydrocyanic acid yield as long as the conversion of methanol is in a specific range. It is therefore reasonable to compare Example Al with Comparative Example Al to be critical

1007120^

-Show the importance of the difference in the placement of the feed openings for the starting material and in the density of the catalyst.

&lt;Example A2&gt;

Hydrocyanic acid was synthesized by conducting the ammoxidation reaction of methanol using the same reactor and catalyst used in Comparative Example A1. The reaction was carried out in the same manner as in Comparative Example A1 except that the relative height H of the methanol feed openings was brought to 144% and the reaction temperature was changed to 44 ° C. To bring a height H to 144%, the methanol feed openings were provided at the position 1.5 m above the air feed openings located at the bottom of the reactor. Furthermore, the catalyst was packed in the reactor in an amount of 40 kg to make the height of the stationary catalyst bed at 1.04 m. At this time, the density of the catalyst bed at height H was 180 kg / m³.

The results of the reaction were such that the hydrocyanic acid yield was 91.5% and the conversion of methanol was 97.6%. Thus, it was found that the yield of hydrocyanic acid obtained in this Example A2 was significantly higher than that of Comparative Example A1.

cComparative Example A2&gt;

Acrylonitrile was synthesized by conducting the ammoxidation reaction of propylene using the same reactor used in Comparative Example A1 and a molybdenum bismuth nickel iron antimony potassium catalyst supported on silica. The relative height H of the supply openings for a mixed gas of propylene and ammonia was brought to 15%. To bring the height H to 15%, the supply openings for propylene and ammonia were provided at the position 0.2 m above the supply openings for air which were placed at the bottom of the reactor. Furthermore, the catalyst was packed in

7120

-32th reactor in an amount of 40 kg to bring the height of the stationary catalyst bed to 1.35 m. At that time, the density of the catalyst bed at height H was 460 kg / m³.

By controlling the composition of the gases to the reactor so that the oxygen / propylene molar ratio would be 2.1, and the ammonia / propylene molar ratio would be 1.2, and by controlling the gas velocity at 0.5 m / s, the reaction was carried out at a temperature of 430 ° C under a pressure of 0.5 kg / cm²G.

The results of the reaction were such that the yield of acrylonitrile was 78.0% and the conversion of propylene was 99.3%.

&lt;Example A3&gt;

Acrylonitrile was synthesized by conducting the ammoxidation reaction of propylene using the same reactor and catalyst used in Comparative Example A2. The reaction was carried out in the same manner as in Comparative Example A2 except that the relative height H of the methanol feed openings was adjusted to 65%. To achieve a height H of 65%, the supply openings for propylene were provided at the position 1.1 m above the supply openings for air placed on the bottom of the reactor. Furthermore, the catalyst was packed in the reactor in an amount of 50 kg to bring the height of the stationary catalyst bed to 1.69 m. At this time, the density of the catalyst bed at height H was 230 kg / m³.

The results of the reaction were such that the yield of acrylonitrile was 81.1% and the conversion of propylene was 98.5%. Thus, it was found that the yield of acrylonitrile obtained in this example was significantly higher than that of Comparative Example A2.

In examples of the present invention, focusing on the ammoxidation of methanol, the height of a stationary catalyst bed, the contact time (sec), the

Q 0 7 'i 2 0'

-33 yield of hydrocyanic acid, the conversion of methanol, and the height of the feed openings for methanol are defined as follows.

Stationary catalyst bed height [m] = (amount of packed catalyst [kg] / {(mass density of catalyst [kg / m³]) x (effective diameter of reactor [m²] ) }

Contact time (sec) = (Volume of the catalyst in the static state, existing above supply ports for methanol gas [m³]) / (volume flow rate of the entire feed gas under reaction conditions [m³/ sec])

Hydrocyanic acid yield [%] = {(weight of carbon in produced hydrocyanic acid [kg] / weight of carbon in supplied methanol [kg])} x 100

Methanol conversion [%] = {(weight of carbon to converted methanol [kg]) / (weight of carbon to fed methanol [kg])} x 100

Height H of supply openings for methanol (%) = {(distance between supply openings for oxygen-containing gas and those for methanol [m]) / (height stationary catalyst bed [m])} x

100

Catalysts used for the following reactions and methods for their preparation are as follows. [Catalyst 1]

A catalyst with the empirical formula:

Fe₁₀Sb₁₉P₆V₁Cu₂₅Mo_0ilO_73&lt;3 (SiO₂) ₆₀ was prepared in the following manner.

(I) 247.3 antimony trioxide was taken.

(II) 385 ml nitric acid (specific gravity: 1.38) and

480 ml of water were mixed and the mixture heated. 49.9 grams of electrolytic iron powder were added gradually

-34ane and dissolved in the mixture. 54.0 g copper nitrate was then added to and dissolved in this solution.

(III) 10.5 g of ammonium metavanadate and 1.6 g of ammonium paramolybdate were dissolved in 300 ml of water.

(IV) 1,612 g of silica sol (SiO₂: 20 wt%) was taken.

(IV), (I) and (III) were added sequentially with stirring to (II) in the order mentioned. The pH of this mixture was adjusted to 2 by the addition of a 15% aqueous ammonia solution. The resulting slurry was heat-treated at 98 ° C for 3 hours with stirring, then 61.8 g of phosphoric acid (85% content) was added to the slurry. The mixture was stirred thoroughly. The slurry was then spray dried by a rotary disk type spray dryer. The resulting fine spherical particles were calcined at 200 ° C for 2 hours, 500 ° C for 3 hours and at 800 ° C for 3 hours.

[Catalyst 2]

A catalyst with the empirical formula: Fe₁₀Sb₁₉P_BV₁Cu₂O_{77i 5} (SiO₂) ₆₀ was prepared in the same manner as in the case of Catalyst 1.

&lt;Catalyst 3&gt;

A catalyst with the empirical formula:

Fe₁₀Sb₂₀P₁₀V₁,₂B_O,₂O₈₃, 3 (θίθ₂) θ₀ was prepared in the same manner as in the case of Catalyst 1, except that boric anhydride was used as the starting material for component B and that a solution of this boric anhydride in water was added after the addition of antimony trioxide powder.

&lt;Catalyst 4&gt;

A catalyst with the empirical formula:

Fe₁₀Sb₂₅P₁₄V_{2 5}Mn_{2 5}O_{in 25} (SiO₂) ₈₀ was prepared in the same manner as in the case of Catalyst 1, except that manganese nitrate was used as

712 0’

-3510 starting material for the Mn component and that a solution of this manganese nitrate in water was added after the addition of antimony trioxide powder.

&lt;Comparative Example B1&gt;

43.3 kg of Catalyst 1 was packed in a fluidized bed reactor with an internal diameter of 20 cm and a height of 5 m (the height of the stationary catalyst bed: 120 cm). Air was supplied to the bottom of this reactor, and a mixed gas of methanol and ammonia was supplied from the position 30 cm above the air supply openings. By controlling the composition of the gases to the reactor so that the oxygen / methanol molar ratio would be 1.15, and the ammonia / methanol molar ratio would be 1.0, and by controlling the gas velocity at 50 cm / s, the reaction was carried out at a temperature of 430 ° C under a pressure of 0.5 kg / cm²G. At this time, the contact time was 2.0 seconds, the height H of the methanol feed openings was 25%, and the density of the catalyst bed at the height H was 460 kg / m³. The results of the reaction were as follows: one hour after the start of the reaction, the hydrocyanic acid yield was 82.1% and the conversion of methanol 99.2%; and after 20 hours after the start of the reaction, the hydrocyanic acid yield was 80.1% and the conversion of methanol 99.0%. Thus, it was found that the hydrocyanic acid yield decreased over time.

&lt;Example Bl&gt;

The ammoxidation reaction of methanol was carried out in the same manner as in Comparative Example B1, except that a mixed gas of methanol and ammonia, to which steam had been added, was fed to the reactor from the position 70 cm above the air supply openings placed in the bottom part of the reactor. The molar ratio of water vapor to methanol was brought to 1.0. At this time, the contact time was 1.4 seconds, the height H of the methanol feed openings was 58% and

The density of the catalyst bed at height H was 220 kg / m³. The results of the reaction were as follows: one hour after the start of the reaction, the hydrocyanic acid yield was 84.5% and the conversion of methanol 98.6%; and 20 hours after the start of the reaction, the hydrocyanic acid yield was 84.0% and the conversion of methanol 98.3%. Thus, the following was found: the yield of hydrocyanic acid obtained in Example B1 was higher than the yield of hydrocyanic acid obtained in Comparative Example BI; the hydrocyanic acid yield did not decrease over time; and improvement in the results of the reaction and the stabilization of the catalytic activity can be obtained simultaneously.

cComparative Example B2&gt;

In a fluidized bed reactor with an internal diameter of 4 cm and a height of 1.5 m, 20 pieces of perforated plates of 2 mm thickness were placed, each with perforations with a diameter of 2 mm with an opening ratio of approx. 4 0 %, at an interval of 5 cm, the lowest plate was placed at a position 7 cm above the air supply openings provided in the lower part of the reactor. 538 g of Catalyst 2 was placed in this reactor (the height of the stationary catalyst bed: 40 cm). The ammoxidation reaction of methanol was carried out by supplying air from the bottom of the reactor, and a mixed gas of methanol and ammonia from the position 10 cm above the air supply openings, changing the composition of the gases in three steps to the following way: the molar ratio of oxygen / methanol was brought to 1.36 and that of ammonia / methanol was brought to 1.0 for the first 3 hours; the oxygen / methanol molar ratio was brought to 0.91 and that of ammonia / methanol was brought to 0.67 for the next 2 hours; the oxygen / methanol molar ratio was brought to 1.3 and that of ammonia / methanol to 1.0

712 0

-37 brought for the last 3 hours. So the reaction was run for a total of 8 hours. During this reaction, the reaction temperature was kept at 430 ° C, the pressure of the reaction was at 0.5 kg / cm²G, and the gas velocity was kept at 15 cm / s. At this time, the contact time was 1.7 seconds, the height H of the methanol feed openings was 25%, and the density of the catalyst bed at the height H was 730 kg / m³. The results of the reaction were as follows: one hour after the start of the reaction, the yield of hydrocyanic acid was 82.4% and the conversion of methanol 99.6%, and at 8 hours after the start of the reaction, the yield of hydrocyanic acid was 77, 5% and the conversion of methanol 95.5%. Thus, it was found that the yield of hydrocyanic acid decreased over time.

cExample B2&gt;

The ammoxidation reaction of methanol was carried out in the same manner as in Comparative Example B2, except that a mixed gas of methanol and ammonia was fed to the reactor from the position 3 cm above the air supply openings located in the lower part of the reactor and that air mixed with steam was supplied to the reactor for 2 hours in the middle phase of the reaction. The molar ratio of steam to methanol was brought to 0.67. At this time, the height H of the feed openings for methanol was 75% and the density of the catalyst bed at the height H was 250 kg / m³. The results of the reaction were as follows: one hour after the start of the reaction, the hydrocyanic acid yield was 85.9% and the conversion of methanol 98.1%; and at 8

0 hours after the start of the reaction, the hydrocyanic acid yield was 86.1% and the conversion of methanol 98.3%. The hydrocyanic acid yield in Example B2 had not decreased as much over time as compared to Comparative

Example B2. Thus, the catalyst undergoes hardly any reduction in catalytic activity because the catalyst is stabilized even below that

Reaction conditions in which the ratio of oxygen to methanol is low.

cComparative Example B3&gt;

44.5 kg of Catalyst 2 was packed into the same reactor 5 as that used in Comparative Example B1 (stationary catalyst bed height: 120 cm). Air was supplied from the bottom of this reactor, and a mixed gas of methanol and ammonia was supplied from the position 30 cm above the air supply openings. By controlling the composition of the gases to the reactor so that the oxygen / methanol molar ratio would be 1.36, and the ammonia / methanol molar ratio would be 1.0, and by controlling the gas velocity to 55 cm / s, the reaction was carried out at a temperature of 430 ° C under a pressure of 0.5 kg / cm²G. At this time, the contact time was 1.8 seconds, the height H of the methanol feed openings was 25%, and the density of the catalyst bed at the height H was 430 kg / m³. The results of the reaction were as follows: 3 hours after the start of the reaction, the hydrocyanic acid yield was 84.4% and the conversion of methanol 99.9%; and 300 hours after the start of the reaction, the hydrocyanic acid yield was 81.5% and the conversion of methanol 99.2%. Thus, the hydrocyanic acid yield decreased with time.

&lt;Example B3&gt;

The ammoxidation reaction of methanol was carried out in the same manner as in Comparative Example B3, except that a mixed gas of methanol, ammonia and steam was supplied from the position 140 cm above the air supply openings located in the lower part of the reactor and that the reaction temperature was changed at 435 ° C. The steam to methanol molar ratio was brought to 1.0. At this time, the contact time was 1.3 seconds, the height H of the methanol feed openings was 117%, and the

The density of the catalyst bed at height H was 190 kg / m³. The results of the reaction were as follows: 3 hours after the start of the reaction, the hydrocyanic acid yield was 88.2% and the conversion of methanol 99.7%; and on

3 00 hours after the start of the reaction, the hydrocyanic acid yield was 88.0% and the conversion of methanol 99.3%; and 900 hours after the start of the reaction, the hydrocyanic acid yield was 88.0% and the conversion of methanol 99.0%. The hydrocyanic acid yield in Example B2 was higher than that in

Comparative Example B3 and was kept at a constant value for many hours. Thus excellent results were obtained from the reaction and at the same time the stabilization of the catalytic activity was maintained. &lt;Examples B4 and B5&gt;

454 g of Catalyst 3 (Example B4) or Catalyst 4 (Example B5) were packed in the same reactor as used in Comparative Example B2 (stationary catalyst bed height: 35 cm). Air was supplied from the lower part of this reactor, while a mixed gas of methanol and ammonia was supplied from the position 30 cm above the air and water vapor supply openings from the position 5 cm above this. By controlling the composition of the gases to the reactor so that the oxygen / methanol molar ratio would be 1.36, and the ammonia / methanol molar ratio would be 1.0, and the steam / methanol molar ratio would be 1.0 and by controlling the gas velocity at 15 cm / s, the reaction was carried out at a temperature of 440 ° C under a pressure of 0.5 kg / cm²G. At this time, the contact time was 1.5 seconds, the height H of the methanol feed openings was 86%, and the density of the catalyst bed at the height H was 240 kg / m³. The results of the reaction were as follows.

<td>After</td><td>1 hour</td><td>After 20 hours</td>

<td>Yield</td><td>Conversion</td><td>Yield Conversion</td>

<td>On</td><td>On</td><td>on to</td>

<td>Hydrocyanic acid</td><td>Methanol</td><td>Hydrocyanic acid Methanol</td>

<td> (%)</td><td> (%)</td><td> (%) (%)</td>

<td>Example</td><td> 4</td><td> 86.0</td><td> 98.5</td><td> 86.2</td><td> 98.7</td>

<td>Example</td><td> 5</td><td> 85.5</td><td> 99.0</td><td> 85.2</td><td> 98.9</td>

100 7:20-¹

-4110