JP5780069B2 - Method for producing conjugated diene - Google Patents

Description

  The present invention relates to a method for producing a conjugated diene, and more particularly to a method for producing a conjugated diene such as butadiene by a catalytic oxidative dehydrogenation reaction of a monoolefin having 4 or more carbon atoms such as n-butene.

A method for producing a conjugated diene such as butadiene by subjecting a monoolefin such as n-butene to an oxidative dehydrogenation reaction in the presence of a catalyst is conventionally known.
This reaction proceeds, for example, according to the following reaction formula, and water is by-produced.
C ₄ H ₈ + 1 / 2O ₂ → C ₄ H ₆ + H ₂ O
The production of butadiene by the catalytic oxidative dehydrogenation reaction of n-butene is industrially made from a C ₄ fraction (C ₄ hydrocarbon mixture, hereinafter sometimes referred to as “BB”) by-produced by naphtha cracking. In the butadiene extraction and separation process, a mixture containing 1-butene, 2-butene, butane and the like obtained by separating butadiene in an extractive distillation column (hereinafter, this mixture may be referred to as “BBSS”). ) Has been proposed for producing butadiene from butene contained therein.

  Specifically, butene is vapor-phase contact oxidatively dehydrogenated to produce butadiene, the product gas containing butadiene is cooled to remove high-boiling by-products from the product gas, and then aldehydes in the product gas. There is known a method of compressing a product gas through a step of removing butadiene and recovering butadiene from the compressed product gas, and butadiene from the product gas through a step of removing aldehydes in the product gas. It has been proposed to eliminate the problem of adhesion of carbon- or tar-like substances by-produced by pressure absorption for recovering etc. to the inside of the compressor and the discharge part (for example, Patent Document 1 or Patent Document 2). On the other hand, Patent Document 3 describes a composite oxide catalyst containing at least one of molybdenum, iron, nickel or cobalt and silica as a catalyst for producing butadiene by the oxidative dehydrogenation reaction of butene. There is no description of the manufacturing method.

JP 60-115532 A JP 60-126235 A JP 2003-220335 A

  According to the study by the present inventors, when butadiene is produced from butene by the method of Patent Documents 1 to 3, the product gas obtained by the oxidative dehydrogenation reaction includes a raw material gas (for example, BB, BBSS). Branched monoolefins such as isobutene and linear n-butene are converted into high-boiling by-products in the oxidative dehydrogenation reaction, and they accumulate on the catalyst and receive a thermal history during a long-term operation. It was found that the components adhered and coking of the catalyst proceeded.

Further, the coked catalyst is supplied with the raw material gas stopped during the operation of the reactor, and a gas containing molecular oxygen is supplied to the reactor to remove the carbon component of the catalyst to regenerate the catalyst (hereinafter, After starting the supply of the source gas again, the oxidative dehydrogenation reaction can be continued, but the catalyst layer in the reactor before and after the start of the oxidative dehydrogenation reaction can be continued. It has been found that when decoking is performed with a large change in porosity, the catalyst cracks because the pressure inside the catalyst pores suddenly increases due to carbon dioxide generated by the combustion of the carbon component.

Once the catalyst is cracked, it cannot be regenerated, and the cracked catalyst reduces the porosity of the reactor and further increases the differential pressure across the catalyst layer of the reactor. For this reason, when the catalyst is cracked, the reactor must be opened and the catalyst replaced, and productivity is significantly reduced.
The present invention has been made in view of the above problems, and in a method for producing a conjugated diene such as butadiene by a catalytic oxidative dehydrogenation reaction of a monoolefin such as n-butene, the cracking of the catalyst is prevented and the reaction is performed. An object of the present invention is to provide an industrially advantageous method for producing butadiene that can be stably operated without opening the vessel and replacing the catalyst.

As a result of intensive studies to solve the above problems, the present inventors have determined that the porosity of the catalyst layer before and after the start of the oxidative dehydrogenation reaction is related to cracking due to the caulking of the catalyst. By controlling the porosity of the catalyst layer in the range of 86 to 100% compared to the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction, cracking of the catalyst due to coking can be prevented, and further, the activity of the catalyst can be restored by decoking It has been found that butadiene can be stably produced without requiring a long production stop period associated with the opening of the reactor.

That is, the gist of the present invention is that a raw material gas containing a monoolefin having 4 or more carbon atoms and a molecular oxygen-containing gas are mixed and supplied to a fixed bed reactor having a catalyst layer to perform an oxidative dehydrogenation reaction. Thus, when the corresponding conjugated diene is produced and the product gas containing the conjugated diene is obtained, the catalyst layer is provided with 2 to 5 catalyst layers having different dilution ratios , and an oxidation represented by the following formula (1) Production of a conjugated diene characterized in that the porosity ε of the catalyst layer during the dehydrogenation reaction is controlled in the range of 86 to 100% compared to the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction. Lies in the way.

  However,

(In the above formula, Δp is the pressure difference [kPa] between the inlet and outlet of the reactor, L is the catalyst layer length [m], g _c is the gravitational acceleration [kg / m · s ² ], and μ is supplied to the reactor. The mixed gas viscosity [kg / m / s], u is the linear velocity [m / s] of the mixed gas supplied to the reactor, d _p is the catalyst particle diameter [m], G
Is the mass rate [kg / m ² · s] of the mixed gas supplied to the reactor. )

  According to the present invention, the stable operation of the plant can be continuously performed by maintaining the shape of the catalyst in the reactor.

It is a systematic diagram which shows embodiment of the manufacturing method of the conjugated diene of this invention.

Hereinafter, the present invention will be described in detail, but the present invention is not limited to the following description, and various modifications can be made within the scope of the gist of the present invention.
The method for producing a conjugated diene of the present invention comprises butene (n-butene such as 1-butene and / or 2-butene).
(Butene, isobutene), pentene, methylbutene, dimethylbutene, etc., can be effectively applied to the production of the corresponding conjugated dienes by catalytic oxidative dehydrogenation of monoolefins having 4 or more carbon atoms, preferably 4 to 6 carbon atoms. it can. Among these, it is most suitably used for the production of butadiene from butene and further from n-butene (1-butene and / or 2-butene).

  The raw material gas of the present invention contains a monoolefin having 4 or more carbon atoms, but it is not necessary to use an isolated monoolefin having 4 or more carbon atoms as the raw material gas, and any mixture can be used as necessary. Can be used. For example, when obtaining butadiene, high-purity n-butene (1-butene and / or 2-butene) can be used as a raw material gas, but the C4 fraction (BB) produced as a by-product in the above-described naphtha decomposition is used. ) Dehydrogenation or oxidative dehydration of fractions (BBSS) and n-butane mainly composed of n-butene (1-butene and / or 2-butene) obtained by separating butadiene and i-butene (isobutene) from Butene fraction produced by elementary reaction, fraction obtained by separating 1-butene from BBSS (raf-3), 1-butene obtained by dimerization of ethylene, cis-2-butene, trans-2-butene Alternatively, a gas containing a mixture thereof, or a hydrocarbon having 4 carbon atoms obtained when fluid heavy oil fraction is subjected to fluid catalytic cracking can also be used. In addition, the main component as used herein is usually 40% or more, preferably 60% or more, more preferably 70% or more with respect to the raw material gas.

  In addition, the source gas of the present invention may contain an arbitrary impurity as long as the effects of the present invention are not impaired. In the case of producing butadiene from n-butene (1-butene and / or 2-butene), specific examples of impurities that may be included include branched monoolefins such as isobutene; propane, n-butane, i- And saturated hydrocarbons such as butane and pentane; olefins such as propylene and pentene; dienes such as 1,2-butadiene; and acetylenes such as methylacetylene, vinylacetylene and ethylacetylene. The amount of this impurity is usually 40% or less, preferably 30% or less, more preferably 10% or less, and particularly preferably 1% or less. If this amount is too large, the concentration of 1-butene or 21-butene as the main raw material will decrease and the reaction will slow down, or the yield of butadiene as the target product will tend to decrease.

The molecular oxygen-containing gas of the present invention is usually a gas containing 10% or more of molecular oxygen, preferably 15% or more, more preferably 20% or more, and specifically preferably air. . In addition, from the viewpoint of increasing the cost necessary for industrially preparing the molecular oxygen-containing gas, the upper limit of the molecular oxygen content is usually 50% or less, preferably 30% or less, More preferably, it is 25% or less. Moreover, the molecular oxygen-containing gas may contain an arbitrary impurity as long as the effects of the present invention are not impaired. Specific examples of impurities that may be included include nitrogen, argon, neon, helium, CO, CO ₂ , and water. In the case of nitrogen, the amount of this impurity is usually 90% or less, preferably 85% or less, more preferably 80% or less. In the case of components other than nitrogen, it is usually 10% or less, preferably 1% or less. When this amount is too large, it tends to be difficult to supply oxygen necessary for the reaction.

  In the present invention, when supplying the raw material gas to the reactor, the raw material gas and the molecular oxygen-containing gas are mixed to obtain a mixed gas, and the mixed gas is supplied to the reactor. , And water (steam) may be supplied to the reactor. Nitrogen gas adjusts the concentration of combustible gas and oxygen in the same way as nitrogen gas, because the concentration of combustible gas and oxygen is adjusted so that the mixed gas does not form squeal. For reasons and to suppress coking of the catalyst, it is preferable that water (water vapor) and nitrogen gas are further mixed with the mixed gas and supplied to the reactor.

  Further, the method of supplying the source gas, the molecular oxygen-containing gas, the nitrogen gas, and water (water vapor) is not particularly limited, and may be supplied by separate pipes, but the formation of the blast is surely avoided. Therefore, before obtaining the mixed gas, nitrogen gas is supplied to the source gas in advance, or nitrogen gas is supplied to the molecular oxygen-containing gas, and in this state, the source gas and the molecular oxygen-containing gas are supplied. Is preferably mixed to obtain a mixed gas.

  When the raw material gas supplied to the reactor is mixed with the molecular oxygen-containing gas, it becomes a mixture of oxygen and combustible gas, so that each gas (raw material gas, air, and necessary) is not included in the explosion range. The composition control at the inlet of the reactor is performed while monitoring the flow rate with a flow meter installed in a pipe that supplies nitrogen gas and water (steam) according to the conditions. For example, n-butene (1-butene and / or When butadiene is produced from 2-butene), it is adjusted to the range of the mixed gas composition described later. The explosive range referred to here is a range having a composition in which oxygen and a combustible gas of combustible gas ignite in the presence of some ignition source. It is known that if the concentration of combustible gas is lower than a certain value, it will not ignite even if an ignition source is present, and this concentration is called the lower explosion limit. It is also known that if the concentration of combustible gas is higher than a certain value, it will not ignite even if an ignition source is present, and this concentration is called the upper limit of explosion. Each value depends on the oxygen concentration. In general, the lower the oxygen concentration, the closer the values are, and the two match when the oxygen concentration is ready. The oxygen concentration at this time is called a critical oxygen concentration. If the oxygen concentration is lower than this, the mixed gas will not be ignited regardless of the concentration of the combustible gas.

  When starting the reaction, first adjust the amount of molecular oxygen-containing gas, nitrogen, and water vapor supplied to the reactor so that the oxygen concentration at the inlet of the reactor is below the critical oxygen concentration, then the combustible gas ( The supply of flammable gas (mainly raw material gas) and molecular oxygen-containing gas such as air is increased so that the concentration of flammable gas is higher than the upper explosion limit. It is good to go. When the supply amount of the combustible gas (mainly source gas) and the molecular oxygen-containing gas is increased, the supply amount of the mixed gas may be made constant by decreasing the supply amount of nitrogen and / or water vapor. By doing so, the residence time of the mixed gas in the piping and the reactor can be kept constant, and the pressure fluctuation can be suppressed.

  In addition, even if it is outside the explosion range, it may ignite if kept for a certain period of time under a certain temperature and pressure condition. This holding time is called ignition delay time. When designing around the reactor, it is necessary to design so that the residence time of the raw material piping and the product gas piping is less than the ignition delay time. Although the ignition delay time depends on temperature, pressure, and composition, it cannot be generally stated. However, the residence time of the mixed gas pipe is 1000 seconds or less, the residence time of the product gas pipe is 10 seconds or less, or the product gas is 350 seconds within 10 seconds. It is desirable to cool to below.

The typical composition of the mixed gas is shown below.
<Combination gas composition>
n-butene: 50 to 100 vol / vol% with respect to the total of C ₄ fractions
C ₄ fractions total: 5-15 vol%
O _{2: C 4} fraction summed for 40~120vol / vol%
N _{2: C 4} fraction summed for 500~1000vol / vol%
90 to 900 vol / vol% with respect to the total of H ₂ O: C ₄ fractions
The lower limit of the ratio of the raw material gas in the mixed gas of the present invention is usually 4.2 vol% or more, preferably 7.6 vol% or more, more preferably 9.0 vol% or more. As this lower limit value increases, the size of the reactor can be reduced, and the cost for construction and operation tends to decrease. On the other hand, the upper limit is usually 20.0 vol% or less, preferably 17.0 vol% or less, and more preferably 15.0 vol% or less. As this upper limit value becomes smaller, the causative substance of coking on the catalyst in the raw material gas also decreases, so that the catalyst coking is less likely to occur.

The reactor used for the oxidative dehydrogenation reaction of the present invention is not particularly limited, and specific examples include a tubular reactor, a tank reactor, or a fluidized bed reactor, preferably a fixed bed reactor, Among them, a fixed-bed multitubular reactor and a plate-type reactor are preferable, and a fixed-bed multitubular reactor is most preferable.
The reactor of the present invention is packed with an oxidative dehydrogenation reaction catalyst described later, and the catalyst is a fixed bed used for producing a corresponding conjugated diene from a monoolefin having 4 or more carbon atoms and oxygen. It is preferable that the reactor is used by being charged, and it is more preferable that n-butene is used by being charged in a fixed bed reactor used to react with oxygen to produce butadiene and water.

  The reactor used in the present invention has a catalyst layer, and this catalyst layer is filled with a certain amount of catalyst, and is formed in a layer shape from the inlet of the reactor toward the product gas outlet of the reactor. Is done. When a raw material is supplied to the catalyst layer, a conjugated diene as a product is generated. The catalyst layer may be formed of only the catalyst or may be formed of a mixture of a non-reactive solid and a catalyst, but it is possible to suppress a rapid temperature rise of the catalyst layer due to heat generation during the reaction. Therefore, it is preferable to form a mixture of a non-reactive solid and a catalyst. In this case, the dilution rate represented by the following formula (3) of the catalyst layer is not particularly limited, but is usually 99 vol% or less, preferably 90 vol% or less, more preferably 80 vol% or less. The smaller the upper limit value, the smaller the reactor can be made, and the construction cost and operation cost can be suppressed.

Dilution rate (vol%) = [(volume of solids not reactive with catalyst) / (volume of catalyst + volume of solids not reactive with catalyst)] × 100 (3)
In the present invention, the catalyst layer provided in the reactor is not particularly limited, but may be one or more layers, and preferably 2 to 5 layers. As the number of catalyst layers increases, the catalyst filling operation tends to become complicated, and as the number of catalyst layers decreases, it tends to be easier. When two or more catalyst layers are provided in the reactor, the dilution rate of each catalyst layer can be appropriately determined depending on the reaction conditions and reaction temperature, but it is preferable to provide catalyst layers having different dilution rates. Further, it is more preferable to dispose the catalyst layer so that the dilution rate becomes smaller toward the product gas outlet side of the reactor than the catalyst layer provided closest to the raw material gas inlet side of the reactor.

  The reason for this is that as the raw material gas in the mixed gas supplied to the reactor decreases as it goes toward the reactor product gas outlet, which is the flow direction of the raw material gas, the reaction of the raw material gas is reduced by reducing the dilution rate. This is because the amount can be kept constant, the conversion rate of monoolefin having 4 or more carbon atoms in the raw material gas can be increased, and further, a rapid temperature drop on the product gas outlet side of the reactor can be suppressed. .

In addition, the packing length of the catalyst layer is the activity of the catalyst to be packed (when diluted with a non-reactive solid, the activity as a diluted catalyst), the size of the reactor, the reaction raw material gas temperature, the reaction Once the temperature and reaction conditions are determined, it can be determined by mass balance and heat balance calculations.
The non-reactive solid used in the present invention is stable under conjugated diene formation reaction conditions, and is a material that is not reactive with raw materials such as monoolefins having 4 or more carbon atoms, and products such as conjugated diene. If it is a thing, it will not specifically limit. Generally, it is sometimes called an inert ball. Specific examples include ceramic materials such as alumina and zirconia. Moreover, the shape is not specifically limited, Any of spherical shape, a column shape, a ring shape, and an indefinite shape may be sufficient. Moreover, the magnitude | size should just be a magnitude | size equivalent to the catalyst used by this invention.

Moreover, in this invention, you may have a layer formed only in a solid substance without reactivity other than said catalyst layer in a fixed bed reactor. Usually, it is preferable to provide a layer formed only of a non-reactive solid substance at both or one of the location closest to the inlet of the source gas of the reactor and the location closest to the product gas outlet. The size of these layers can be appropriately determined according to the size of the catalyst layer, the size of the reactor, and the reaction conditions.

  The oxidative dehydrogenation reaction of the present invention is an exothermic reaction, and the reaction temperature rises due to the reaction, but the upper limit of the reaction temperature is usually 450 ° C. or lower, preferably 430 ° C. or lower, more preferably 400 ° C. or lower. is there. The lower the upper limit, the more rapid catalyst activity can be suppressed. On the other hand, the lower limit of the reaction temperature is usually 250 ° C. or higher, preferably 265 ° C. or higher, more preferably 280 ° C. or higher. As the lower limit increases, the yield of the conjugated diene of the target product tends to increase. The reaction temperature can be controlled to be constant using a heat medium (for example, dibenzyltoluene or nitrite).

  In the present invention, the porosity of the catalyst layer of the reactor represented by the above formula (1) is controlled in the range of 86 to 100% as compared with the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction. It is characterized by the fact that after supplying raw material gas and starting the production of conjugated diene by oxidative dehydrogenation reaction, the supply of all gas to the reactor is stopped and the operation of the reactor is stopped. In addition, the porosity of the catalyst layer may be controlled so as not to deviate from the range of 86 to 100% compared to the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction. Usually, it is in the range of 86 to 100% of the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction, preferably 87 to 98.8%, more preferably 90 to 98. .5%.

The above formula (1) for determining the porosity of the catalyst layer is the Ergun formula, which is one of the calculation formulas for calculating the pressure loss of the packed bed filled with the particle group proposed by Ergun, which is filled with the solid catalyst. It is derived by applying to the catalyst layer of the reactor.
In the present invention, the closer the porosity of the catalyst layer of the reactor is to 100% compared to the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction, the higher the effect of suppressing the cracking of the catalyst contained in the catalyst layer. However, the number of operation operations for controlling the porosity increases, operation management becomes complicated, and productivity decreases. On the other hand, if the porosity of the catalyst layer of the reactor is less than 86% of the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction, the carbon component is burned during decoking. Due to the generated carbon dioxide, the pressure inside the catalyst pores suddenly increases, so the catalyst is cracked.

As means for controlling the porosity in the present invention, it is preferable to manipulate the supply amount of the source gas and / or the temperature in the reactor. As a preferable reason, the change in the porosity is caused by the deposition of high-boiling by-products due to branched monoolefins such as isobutene and linear n-butene contained in the raw gas (for example, BB or BBSS) on the catalyst. However, the carbon component adheres on the catalyst by receiving the heat history in the long-term operation, and the coking of the catalyst proceeds. By this coking, the porosity of the catalyst layer tends to be lower than that at the start of the oxidative dehydrogenation reaction, but in the range of 86 to 100% compared to the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction. If present, the porosity of the coked catalyst is reduced by stopping the supply of the raw material gas during the operation of the reactor and supplying the gas containing molecular oxygen to the reactor to perform the decoking of the catalyst. Can be suppressed and the porosity can be increased. After the decoking operation, the supply of the raw material gas can be started again to continue the oxidative dehydrogenation reaction. In the present invention, when a multi-tube reactor is used, one or more reaction tubes are arbitrarily selected, and the reaction tube is controlled by measuring the porosity ε of the above formula (1). Just do it.

As a method for controlling the supply amount of the raw material gas, the supply amount of the raw material gas may be merely reduced, or the reduced raw material gas supply amount may be supplemented with an inert gas such as nitrogen, CO ₂ , or water. Further, as long as the raw material gas supply amount is reduced, the inert gas supply amount is not limited. After the raw material gas supply amount is reduced, the reaction gas containing molecular oxygen is supplied to the reactor.
As the operation of the temperature in the reactor, the reaction temperature may be maintained or reduced at the stage of reducing the supply amount of the raw material gas. In the case of lowering, considering the time loss for raising the temperature again, the temperature is maintained at 150 ° C. or higher, preferably 200 ° C. or higher, more preferably 250 ° C. or higher. In the step of supplying the reaction gas containing molecular oxygen to the reactor after reducing the supply amount of the source gas, the temperature is maintained at 300 ° C. or higher, preferably 320 ° C. or higher, more preferably 350 ° C. or higher. When the temperature is low, it takes time to complete the decoking, and the productivity is lowered.

Decoking is terminated when the difference between the carbon dioxide concentrations at the inlet and outlet of the reactor reaches 100 ppm by volume or less, and after completion, the supply of a raw material gas containing a monoolefin having 4 or more carbon atoms is resumed, or the supply amount is reduced. To increase.
The residence time of the reactor of the present invention is not particularly limited, but the lower limit is usually 0.72 seconds or longer, preferably 0.80 seconds or longer, more preferably 0.90 seconds or longer. There is a merit that the higher the value, the higher the conversion rate of monoolefin in the raw material gas. On the other hand, the upper limit is 7.20 seconds or less, preferably 4.50 seconds or less, and more preferably 3.60 seconds or less. The smaller this value, the smaller the reactor.

In the fixed bed reactor of the present invention, the lower limit of the ratio of the flow rate of the mixed gas to the amount of catalyst in the reactor (hereinafter sometimes abbreviated as SV) is 500 h ⁻¹ or more, preferably , and at 800h ^-1 or more, still more preferably 1000h ^-1 or more. As this value increases, the yield of conjugated diene tends to decrease. Further, while the upper limit is at 5000h ^-1 or less, or preferably 4500H ^-1, still more preferably 4000h ^-1 or less. As this value decreases, the catalyst tends to be less likely to deteriorate even if the oxidative dehydrogenation reaction is continuously performed.

The flow rate difference between the inlet and outlet of the reactor depends on the flow rate of the raw material gas at the reactor inlet and the flow rate of the product gas at the reactor outlet, but the ratio of the outlet flow rate to the inlet flow rate is usually 100. -110 mol%, Preferably, it is 102-107 mol%, More preferably, it is 103-105 mol%. The outlet flow rate increases because the number of molecules increases stoichiometrically in the reaction in which butene is oxidatively dehydrogenated to produce butadiene and water, and in the reaction in which CO and CO ₂ are produced as a side reaction. A small increase in the outlet flow rate is not preferable because the reaction does not proceed, and an excessive increase in the outlet flow rate is not preferable because CO and CO ₂ increase due to side reactions.

  In this invention, you may have a cooling process which cools the product gas containing the conjugated diene obtained from a reactor. The cooling step is not particularly limited as long as the product gas obtained from the outlet of the reactor can be cooled, but a method of cooling by directly contacting the cooling solvent and the product gas is preferably used. Although it does not specifically limit as a cooling solvent, Preferably it is water and alkaline aqueous solution, Most preferably, it is water.

  The cooling temperature of the product gas varies depending on the product gas temperature obtained from the reactor outlet, the kind of the cooling solvent, etc., but is usually 5 to 100 ° C., preferably 10 to 50 ° C., more preferably 15 to 40. Cool to ° C. The higher the temperature to be cooled, the lower the construction cost and the cost required for operation. The lower the temperature, the lower the load on the process of compressing the product gas.

Moreover, in this invention, you may have a dehydration process which removes the water | moisture content contained in the product gas discharged | emitted from a reactor. Providing a dehydration step is preferable because it can prevent equipment corrosion due to moisture in each step in the subsequent process and accumulation of impurities in the solvent used in the solvent absorption step and solvent separation step described later.
The dehydration process of the present invention is not particularly limited as long as it is a process capable of removing moisture contained in the product gas. The dehydration step may be performed anywhere as long as it is a subsequent step of the reactor, but it is preferable to perform the dehydration step after the above-described cooling step. Usually, the amount of water contained in the product gas discharged from the reactor varies depending on the type of raw material gas, the amount of molecular oxygen-containing gas, and water vapor mixed with the raw material gas. It contains ~ 35 vol%, preferably 10-30 vol% moisture. In the dehydration step, the water content in the product gas is reduced to a certain level of tatami mat, but the water content is usually 100 volppm to 2.0 vol%, preferably 200 volppm to 1.0 vol%. Moreover, as a dew point, it is 0-100 degreeC, Preferably, it is 10-80 degreeC.

A means for dehydrating water from the generated gas is not particularly limited, and a desiccant (moisture adsorbent) such as calcium oxide, calcium chloride, or molecular sieve can be used. Of these, desiccants (moisture adsorbents) such as molecular sieves are preferably used from the viewpoint of ease of regeneration and ease of handling.
When a desiccant such as molecular sieve is used in the dehydration step, high-boiling by-products contained in the product gas other than water are adsorbed and removed. The high-boiling by-products removed here are anthraquinone, fluorenone, phthalic acid, and the like.

  The water content in the product gas obtained through the dehydration step is usually 10 to 10,000 volppm, preferably 20 to 1000 volppm, and the dew point is -60 to 80 ° C, preferably -50 to 20 ° C. . As the water content in the product gas increases, the contamination of the reboiler of the solvent absorption tower and the solvent separation tower tends to increase. On the other hand, when the content decreases, the utility cost used in the dehydration process tends to increase. .

  In this invention, you may have the collection | recovery process which collect | recovers the product gas discharged | emitted from the reactor exit. From the viewpoint of reducing the energy cost required for the separation, it is preferable to recover the product gas by absorbing it in a solvent. The recovery process is not particularly limited as long as it can be recovered by absorbing the product gas in a solvent. The recovery step may be performed anywhere as long as it is a subsequent step of the reactor, but it is preferable to perform the recovery step after the above-described dehydration step.

As a specific method of absorbing the product gas in the solvent in the recovery step, for example, a method using an absorption tower is preferable. As the type of the absorption tower, a packed tower, a wet wall tower, a spray tower, a cyclone scrubber, a bubble tower, a bubble stirring tank, a plate tower (bubble bell tower, perforated plate tower), a foam separation tower, and the like can be used. A spray tower, a bubble bell tower, and a perforated plate tower are preferable.
In the case of using an absorption tower, the absorption solvent and the product gas are usually brought into countercurrent contact so that the conjugated diene in the product gas and the unreacted monoolefin having 4 or more carbon atoms and the hydrocarbon having 3 or less carbon atoms are used. The compound is absorbed into the solvent. Examples of the hydrocarbon compound having 3 or less carbon atoms include methane, acetylene, ethylene, ethane, methylacetylene, propylene, propane, and allene.

  In the recovery step, when the product gas is recovered using an absorption tower, the pressure in the absorption tower is not particularly limited, but is usually 0.1 to 2.0 MPaG, preferably 0.2 to 1.5 MPaG, Preferably, it is 0.25 to 1.0 MPaG. The larger the pressure, the better the absorption efficiency, and the smaller the pressure, the less energy required for boosting the gas when the gas is introduced into the absorption tower, and the further the reduced oxygen amount in the liquid.

Moreover, the temperature in an absorption tower is although it does not specifically limit, Usually, 0-50 degreeC, Preferably, it is 10-40 degreeC, More preferably, it is 20-30 degreeC. The higher this temperature is, the more advantageous is that oxygen, nitrogen, and the like are less likely to be absorbed by the solvent, and the smaller is the advantage that the absorption efficiency of hydrocarbons such as conjugated dienes is improved.
The absorption solvent for use in the recovery process of the present invention, but are not limited to, aromatic hydrocarbons, saturated hydrocarbons and C ₆ -C ₈ of C ₆ -C _10, an amide compound and the like are used. Specifically, for example, dimethylformamide (DMF), toluene, xylene, n-methyl-2-pyrrolidone (NMP) and the like can be used. Among these, C _{6 to} C ₈ aromatic hydrocarbons are preferable because toluene is difficult to dissolve inorganic gas, and toluene is particularly preferable.

Although there is no restriction | limiting in particular in the usage-amount of an absorption solvent, It is 1-100 weight times normally with respect to the flow volume of the target product supplied to a collection | recovery process, Preferably, it is 2-50 weight times. As the amount of the absorbing solvent used increases, it tends to be uneconomical, and as the amount used decreases, the recovery efficiency of the conjugated diene tends to decrease.
The solvent in which the product gas obtained in the recovery process is absorbed (hereinafter sometimes referred to as “solvent absorption liquid”) mainly contains the conjugated diene that is the target product, and the solvent absorption of the conjugated diene. The concentration in the liquid is usually 1 to 20% by weight, preferably 3 to 10% by weight. The higher the concentration of the conjugated diene in the solvent absorbent, the more the amount of conjugated diene lost due to polymerization or volatilization tends to increase. The energy cost required tends to increase.

In the present invention, since the solvent absorption liquid obtained in the recovery process also absorbs a small amount of nitrogen and oxygen, the solvent absorption liquid has a degassing step for gasifying and removing nitrogen and oxygen dissolved in the solvent absorption liquid. May be. The degassing step is not particularly limited as long as it is a step capable of gasifying and removing nitrogen and oxygen dissolved in the solvent absorption liquid.
The crude conjugated diene can be obtained through a separation step in which the crude conjugated diene is separated from the solvent absorption liquid of the conjugated diene thus obtained. The separation step is not particularly limited as long as the crude conjugated diene can be separated from the solvent absorption liquid of the conjugated diene, but the crude conjugated diene can be usually separated by distillation separation. Specifically, for example, a conjugated diene is distilled and separated by a reboiler and a condenser, and a conjugated diene fraction is extracted from the vicinity of the top of the column. The separated absorption solvent is extracted from the bottom of the column, and when it has a recovery step that uses the solvent in the previous step, it is recycled as an absorption solvent in the recovery step. Impurities may accumulate during recycling of the solvent, and a part of the solvent may be extracted and removed by a known purification method such as distillation, decantation, sedimentation, or contact treatment with an adsorbent or ion exchange resin. desirable.

  Although the pressure at the time of distillation of the distillation column used at a separation process can be set arbitrarily, it is usually preferred that the column top pressure is 0.05-2.0 MPaG. More preferably, the tower top pressure is 0.1 to 1.0 MPaG, and particularly preferably 0.15 to 0.8 MPaG. If the pressure at the top of the column is too low, a large amount of cost is required to condense the conjugated diene distilled off at a low temperature, and if it is too high, the temperature at the bottom of the distillation column increases, resulting in an increase in steam costs. End up.

The tower bottom temperature is usually 50 to 200 ° C, preferably 80 to 180 ° C, more preferably 100 to 160 ° C. If the column bottom temperature is too low, it will be difficult to distill the conjugated diene from the column top. If the temperature is too high, the solvent will be distilled off from the top of the column. The reflux ratio may be 1 to 10, and preferably 2 to 4.
As the distillation column, either a packed column or a plate column can be used, but multistage distillation is preferred. In order to separate the conjugated diene and the solvent, it is preferable that the theoretical column of the distillation column is 5 or more, particularly 10 to 20 plates. A distillation column having more than 50 stages is not preferable for economics of construction of the distillation column, operational difficulty, and safety management. If the number of stages is too small, separation becomes difficult.

  A crude conjugated diene is obtained in the step of separating the conjugated diene. The crude conjugated diene can be further purified by distillation purification or the like. Although the pressure at the time of distillation of the distillation column used here can be set arbitrarily, it is usually preferred that the column top pressure is 0.05 to 0.4 MPaG. More preferably, the tower top pressure is 0.1 to 0.3 MPaG, and particularly preferably in the range of 0.15 to 0.2 MPaG. If the pressure at the top of the column is too low, a large amount of cost is required to condense the conjugated diene distilled off at a low temperature, and if it is too high, the temperature at the bottom of the distillation column increases, resulting in an increase in steam costs. End up.

The tower bottom temperature is usually 30 ° C to 100 ° C, preferably 40 ° C to 80 ° C, more preferably 50 ° C to 60 ° C. If the column bottom temperature is too low, it will be difficult to distill the conjugated diene from the column top. If the temperature is too high, the amount of condensation at the top of the tower increases and costs increase. The reflux ratio may be 1 to 10, and preferably 2 to 4.
As the distillation column, either a packed column or a plate column can be used, but multistage distillation is preferred. In order to separate impurities such as conjugated diene and furan, it is preferable that the number of theoretical stages of the distillation column is 5 or more, particularly 10 to 20 stages. A distillation column having more than 50 stages is not preferable for economics of construction of the distillation column, operational difficulty, and safety management. If the number of stages is too small, separation becomes difficult. The purified conjugated diene thus obtained is a conjugated diene having a purity of 99.0 to 99.9%.

[Oxidation dehydrogenation catalyst]
Hereinafter, the oxidative dehydrogenation catalyst suitably used in the present invention will be described. The catalyst for the oxidative dehydration reaction used in the present invention is not particularly limited as long as it is a catalyst capable of producing a corresponding conjugated diene using a monoolefin having 4 or more carbon atoms and oxygen as raw materials. Thus, a composite oxide catalyst containing at least molybdenum, bismuth, and cobalt is preferable.

Among these, a composite oxide catalyst represented by the following general formula (2) is more preferable.
Mo _a Bi _b Co _c Ni _{d F e} X _f Y _g Z _h Si _i O _j (2)
Wherein X is at least one element selected from the group consisting of magnesium (Mg), calcium (Ca), zinc (Zn), cerium (Ce) and samarium (Sm), and Y is sodium (Na) , Potassium (K), rubidium (Rb), cesium (Cs) and at least one element selected from the group consisting of thallium (Tl), Z is boron (B), phosphorus (P), arsenic (As) And at least one element selected from the group consisting of tungsten (W), and a to j represent atomic ratios of the respective elements, and when a = 12, b = 0.1 to 7, c = 0 to 10, d = 0 to 10 (provided c + d = 1 to 10), e = 0.05 to 3, f = 0 to 2, g = 0.04 to 2, h = 0 to 3, i = 0 to 0 48, and j satisfies the oxidation state of other elements Numerical value is.)
Further, the composite oxide catalyst is a method for producing a component compound source compound constituting the composite oxide catalyst through a process of integrating and heating in a water system, the molybdenum compound,
A catalyst precursor is produced by heat-treating a raw material compound aqueous solution containing at least one selected from the group consisting of an iron compound, a nickel compound and a cobalt compound and, if necessary, silica, or a dried product obtained by drying the aqueous solution. It is preferable that the catalyst precursor, the molybdenum compound and the bismuth compound are integrated with an aqueous solvent, followed by drying and firing. If the composite oxide catalyst is used, a conjugated diene such as butadiene can be produced with high catalytic activity and high yield, and a reaction product gas having a low aldehyde content can be obtained.

Next, a method for producing a composite oxide catalyst suitable for the present invention will be described.
In this method for producing a composite oxide catalyst, the molybdenum used in the preceding step is molybdenum corresponding to a partial atomic ratio (a ₁ ) of the total atomic ratio (a) of molybdenum, The molybdenum used is preferably molybdenum corresponding to the remaining atomic ratio (a ₂ ) obtained by subtracting a ₁ from the total atomic ratio (a) of molybdenum. The a ₁ is preferably a value satisfying 0.5 <a ₁ / (c + d + e) <3. Furthermore, it is preferable that a ₂ is a value satisfying 0 <a ₂ / b <8.

Source compounds of the above component elements include oxides, nitrates, carbonates, ammonium salts, hydroxides, carboxylic acid salts, ammonium carboxylates, ammonium halide salts, hydrogen acids, acetylacetonates, alkoxides of the component elements. The following are mentioned as the specific example.
Examples of Mo supply source compounds include ammonium paramolybdate, molybdenum trioxide, molybdic acid, ammonium phosphomolybdate, and phosphomolybdic acid.

Examples of Fe source compounds include ferric nitrate, ferric sulfate, ferric chloride, and ferric acetate.
Examples of the Co source compound include cobalt nitrate, cobalt sulfate, cobalt chloride, cobalt carbonate, and cobalt acetate.
Examples of the Ni source compound include nickel nitrate, nickel sulfate, nickel chloride, nickel carbonate, nickel acetate and the like.

Examples of Si source compounds include silica, granular silica, colloidal silica, and fumed silica.
Examples of Bi source compounds include bismuth chloride, bismuth nitrate, bismuth oxide, and bismuth subcarbonate. In addition, Bi component in which X component (one or more of Mg, Ca, Zn, Ce, and Sm) and Y component (one or more of Na, K, Rb, Cs, and Tl) are dissolved. It can also be supplied as a complex carbonate compound of the X component and the Y component.

For example, when Na is used as the Y component, a complex carbonate compound of Bi and Na can be obtained by dropping an aqueous solution of a water-soluble bismuth compound such as bismuth nitrate into an aqueous solution of sodium carbonate or sodium bicarbonate. The precipitate can be produced by washing with water and drying.
In addition, the complex carbonate compound of Bi and the X component is prepared by mixing an aqueous solution of a water-soluble compound such as bismuth nitrate and nitrate of the X component with an aqueous solution of ammonium carbonate or ammonium bicarbonate, etc. It can be produced by washing with water and drying.

When sodium carbonate or sodium bicarbonate is used instead of the above ammonium carbonate or ammonium bicarbonate, a complex carbonate compound with Bi, Na and X components can be produced.
Examples of source compounds of other component elements include the following.
Examples of the source compound for K include potassium nitrate, potassium sulfate, potassium chloride, potassium carbonate, and potassium acetate.

Examples of Rb source compounds include rubidium nitrate, rubidium sulfate, rubidium chloride, rubidium carbonate, and rubidium acetate.
Examples of the Cs supply source compound include cesium nitrate, cesium sulfate, cesium chloride, cesium carbonate, and cesium acetate.
Examples of Tl source compounds include thallium nitrate, thallium chloride, thallium carbonate, and thallium acetate.

Examples of the source compound for B include borax, ammonium oxalate, and boric acid.
Examples of P source compounds include ammonium phosphomolybdate, ammonium phosphate, phosphoric acid, phosphorus pentoxide, and the like.
Examples of the source compound for As include dialsenooctammonium molybdate, ammonium dialseno18 tungstate, and the like.

Examples of W source compounds include ammonium paratungstate, tungsten trioxide, tungstic acid, and phosphotungstic acid.
Examples of the Mg supply vibration compound include magnesium nitrate, magnesium sulfate, magnesium chloride, magnesium carbonate, and magnesium acetate.
Examples of the source compound for Ca include calcium nitrate, calcium sulfate, calcium chloride, calcium carbonate, and calcium acetate.

Examples of the Zn source compound include zinc nitrate, zinc sulfate, zinc chloride, zinc carbonate, and zinc acetate.
Examples of the Ce source compound include cerium nitrate, cerium sulfate, cerium chloride, cerium carbonate, and cerium acetate.
Examples of Sm source compounds include samarium nitrate, samarium sulfate, samarium chloride, samarium carbonate, and samarium acetate.

The raw material compound aqueous solution used in the previous step is an aqueous solution, water slurry or cake containing at least one of molybdenum (corresponding to a ₁ in the total atomic ratio a), iron, nickel or cobalt as a catalyst component. In addition, this aqueous solution, water slurry, or cake may contain the silica as needed.
The raw material compound aqueous solution is prepared by integrating the source compound in an aqueous system. Here, the integration of each component element source compound in an aqueous system means that the aqueous solution or aqueous dispersion of each component element source compound is mixed or / and aged in stages. Say. (B) a method for lumping each of the above-mentioned source compounds in a lump, (b) a method for ramming each of the above-mentioned source compounds in a lump and aging treatment, (c) each of the above-mentioned source compounds Supply of each component element is a method of stepwise mixing, (d) a method of repeatedly mixing and aging the above-mentioned source compounds stepwise, and a method of combining (b) to (d). It is included in the concept of integration of the source compound in an aqueous system. Here, aging refers to the processing of industrial raw materials or semi-finished products under specific conditions such as constant temperature for a certain period of time to obtain the required physical and chemical properties, increase or advance the prescribed reaction, etc. refers to the operation to achieve a a certain time is usually in the range of 10 minutes to 24 hours, it means a normal boiling range of from room temperature to the aqueous solution or dispersion to the constant temperature.

As a specific method of the above integration, for example, a solution obtained by mixing an acidic salt selected from catalyst components, and a solution obtained by mixing a basic salt selected from catalyst components, Specific examples include a method in which a mixture of an iron compound and a nickel compound and / or a cobalt compound is added to an aqueous solution of a molybdenum compound while heating. In addition, it is preferable to perform addition and mixing of silica in this previous step as required.

The raw material compound aqueous solution (slurry) thus obtained is heated to 60 to 90 ° C. and aged.
This aging means that the catalyst precursor slurry is stirred at a predetermined temperature for a predetermined time. This aging increases the viscosity of the slurry, reduces the sedimentation of the solid components in the slurry, and is particularly effective in suppressing the heterogeneity of components in the subsequent drying process. The catalytic activity such as the raw material conversion rate and selectivity of the product catalyst becomes better.

  60-90 degreeC is preferable and the temperature in the said ripening has more preferable 70-85 degreeC. When the aging temperature is less than 60 ° C., the aging effect is not sufficient, and good activity may not be obtained. On the other hand, when it exceeds 90 ° C., the water is often evaporated during the aging time, which is disadvantageous for industrial implementation. Further, if the temperature exceeds 100 ° C., a pressure vessel is required for the dissolution tank, and handling becomes complicated, which is extremely disadvantageous in terms of economy and operability.

The aging time is preferably 2 to 12 hours, and preferably 3 to 8 hours. If the aging time is less than 2 hours, the activity and selectivity of the catalyst may not be sufficiently developed. On the other hand, the aging effect does not increase even if it exceeds 12 hours, which is disadvantageous for industrial implementation.
Any method can be adopted as the stirring method, and examples thereof include a method using a stirrer having a stirring blade and a method using external circulation using a pump.

  The aged slurry is subjected to heat treatment as it is or after drying. There is no particular limitation on the drying method in the case of drying and the state of the dried product to be obtained. For example, a powdery dried product may be obtained using a normal spray dryer, slurry dryer, drum dryer, etc. Alternatively, a block-like or flake-like dried product may be obtained using a normal box-type dryer or a tunnel-type firing furnace.

  The raw material salt aqueous solution or granules or cakes obtained by drying the aqueous salt solution are heat-treated in air at a temperature of 200 to 400 ° C., preferably 250 to 350 ° C. for a short time. There are no particular limitations on the type and method of the furnace at that time, and for example, it may be heated with a dry matter fixed using a normal box-type furnace, tunnel-type furnace, etc., or a rotary kiln. It is possible to heat the dried product while flowing it.

The ignition loss of the catalyst precursor obtained after the heat treatment is preferably 0.5 to 5% by weight, more preferably 1 to 3% by weight. By setting the ignition loss within this range, a catalyst having a high raw material conversion rate and high selectivity can be obtained. Note that the loss on ignition is a value given by the following equation as described above.
Burning loss (%) = [(W ₀ −W ₁ ) / W ₀ ] × 100
W ₀ : Weight (g) of the catalyst precursor after drying at 150 ° C. for 3 hours to remove adhering moisture
W ₁ : Weight (g) after heat-treating the catalyst precursor excluding adhering moisture at 500 ° C. for 2 hours
In the step after the, the integration of the above procatalyst precursor and the molybdenum compound obtained in step (remaining a ₂ corresponds to minus a ₁ equivalent of the total atomic ratio a) and bismuth compounds, carried out in an aqueous solvent . At this time, it is preferable to add ammonia water. The addition of the X, Y, and Z components is also preferably performed in the subsequent step. Further, the bismuth source compound of the present invention is bismuth which is hardly soluble or insoluble in water. This compound is preferably used in the form of a powder. These compounds as the catalyst production raw material may be particles larger than the powder, but are preferably smaller particles in view of the heating step in which thermal diffusion should be performed. Therefore, if these compounds as raw materials are not such particles, they should be pulverized before the heating step.

  Next, the obtained slurry is sufficiently stirred and then dried. The dried product thus obtained is shaped into an arbitrary shape by a method such as extrusion molding, tableting molding or support molding. Next, this is preferably subjected to a final heat treatment for about 1 to 16 hours under a temperature condition of 450 to 650 ° C. As described above, a composite oxide catalyst having a high activity and a desired oxidation product in a high yield can be obtained.

[Process embodiment]
Hereinafter, with reference to the drawings, embodiments of a process for the production method of the present onset Ming conjugated diene, be described by way of example for producing butadiene.
FIG. 1 shows one embodiment of the process of the present invention.
In FIG. 1, 1 is a reactor, 2 is a quench tower, 3, 6 and 13 are coolers (heat exchangers), 4, 7 and 14 are drain pots, 8A and 8B are dehydration towers, and 9 is a heater (heat (Exchanger), 10 is a solvent absorption tower, 11 is a degassing tower, 12 is a solvent separation tower, and 100 to 126 are pipes.

A mixture containing n-butene as a raw material or n-butene such as the above-described BBSS is gasified by a vaporizer (not shown) and introduced from the pipe 101, and from the pipes 102, 103, and 104, nitrogen gas, After introducing air (molecular oxygen-containing gas) and water (water vapor), and heating the mixed gas to about 150 to 400 ° C. with a preheater (not shown), the catalyst was filled from the pipe 100. A multi-tubular reactor 1 (oxidation dehydrogenation reactor) is supplied. The reaction product gas from the reactor 1 is supplied to the quench tower 2 through the pipe 105 and cooled to about 20 to 99 ° C. Cooling water is introduced into the quench tower 2 through the pipe 106 and is in countercurrent contact with the product gas. And the water which cooled the product gas by this countercurrent contact is discharged | emitted from the piping 107. FIG. The cooling waste water is cooled by a heat exchanger (not shown) and is circulated again in the quench tower 2. The product gas cooled in the quench tower 2 is distilled from the top of the tower, and then cooled to room temperature via the cooler 3 from the pipe 108. The condensed water generated by cooling is separated into the drain pot 4 through the pipe 109. The gas after water separation is further pressurized to about 0.1 to 0.5 MPa by the compressor 5 through the pipe 110, and the pressurized gas is cooled again to about 10 to 30 ° C. by the cooler 6 through the pipe 111. Condensed water generated by cooling is separated from the pipe 112 into the drain pot 7. Compressed gas after water separation, dehydration column 8A drying agent such as molecular sieve blanking is filled and is introduced into 8B is dehydrated. In the dehydration towers 8A and 8B, dehydration of the compressed gas and regeneration by heating and drying of the desiccant are performed alternately. That is, the compressed gas is first introduced into the dehydration tower 8A through the pipes 113 and 113a, dehydrated, and supplied to the solvent absorption tower 10 through the pipes 114a and 114. During this time, nitrogen gas heated to about 150 to 250 ° C. is introduced into the dehydration tower 8B through the pipe 122, the heater 9, and the pipes 123, 123a, and 123b, and moisture is desorbed by heating the desiccant. . The nitrogen gas containing the desorbed water is cooled to room temperature by the cooler 13 through the pipes 124 a, 124 b, and 124, and the condensed water is separated from the pipe 125 into the drain pot 14 and then discharged from the pipe 126.

When the desiccant in the dehydration tower 8A reaches saturation, the gas flow path is switched, the compressed gas is dehydrated in the dehydration tower 8B, and the desiccant in the dehydration tower 8A is regenerated.
The regeneration time of the desiccant in the dehydration tower in the dehydration step is not particularly limited, but is usually 6 to 48 hours, preferably 12 to 36 hours, and more preferably 18 to 30 hours.
The dehydrated gas from the dehydration towers 8A and 8B is cooled to about 10 to 30 ° C. by a cooler (not shown) as necessary, and then sent to the solvent absorption tower 10 to be sent from the pipe 115 to the solvent (absorption) Solvent). Thereby, the conjugated diene and the unreacted raw material gas in the dehydrated gas are absorbed by the absorption solvent. The component (off gas) that has not been absorbed by the absorption solvent is discharged from the top of the solvent absorption tower 10 via the pipe 117 and is combusted and discarded. At this time, if a solvent having a relatively low boiling point such as toluene is used as the absorbing solvent, an amount of the solvent that cannot be ignored economically may be volatilized through the pipe 117. In such a case, a step of recovering a solvent having a low boiling point using a solvent having a higher boiling point may be provided at the end of the pipe 117. In this solvent absorption tower 10, the solvent absorption liquid in which butadiene and unreacted source gas are absorbed by the absorption solvent is extracted from the bottom of the solvent absorption tower 10 and fed to the deaeration tower 11 through the pipe 116. Since a certain amount of nitrogen and oxygen are also absorbed in the solvent absorption liquid of butadiene obtained in the solvent absorption tower 10, the solvent absorption liquid is then supplied to the deaeration tower 11 and heated. Gasify and remove dissolved nitrogen and oxygen. At this time, some of the butadiene, the raw material gas, and the solvent may be gasified. Therefore, this is liquefied by a capacitor (not shown) provided at the top of the degassing tower 11 to be solvent. Collect in absorbent. Uncondensed raw material gas, butadiene, and the like are extracted from the pipe 118 as a mixed gas of nitrogen and oxygen, and are circulated to the inlet side of the compressor 5 and processed again in order to increase the recovery rate of the conjugated diene. On the other hand, the degassed treatment liquid from which the solvent absorption liquid has been degassed is sent to the solvent separation tower 12 through the pipe 119.

  In the solvent separation column 12, the conjugated diene is distilled and separated by a reboiler and a condenser, and a crude butadiene fraction is extracted from the top of the column via a pipe 120. The separated absorption solvent is extracted from the bottom of the tower through a pipe 121 and is circulated and used as the absorption solvent of the solvent absorption tower 10.

EXAMPLES Hereinafter, although an Example is given and this invention is demonstrated more concretely, this invention is not limited at all by the following Examples.
<Production Example 1: Preparation of composite oxide catalyst>
54 g of ammonium paramolybdate was dissolved in 250 ml of pure water by heating to 70 ° C. Next, 7.18 g of ferric nitrate, 31.8 g of cobalt nitrate, and 31.8 g of nickel nitrate were dissolved in 60 ml of pure water by heating to 70 ° C. These solutions were gradually mixed with thorough stirring.

Next, 64 g of silica was added and stirred thoroughly. This slurry was heated to 75 ° C. and aged for 5 hours. Thereafter, the slurry was dried by heating and then subjected to heat treatment at 300 ° C. for 1 hour in an air atmosphere.
The resulting catalyst precursor particulate solid (ignition loss: 1.4 wt%) was ground and dispersed in a solution of ammonia water 10ml was added ammonium paramolybdate 40.1g of pure water 150 ml. Next, 0.85 g of borax and 0.36 g of potassium nitrate were dissolved in 40 ml of pure water under heating at 25 ° C., and the slurry was added.

Next, 58.1 g of bismuth carbonate in which 0.45% of Na was dissolved was added and mixed with stirring. After the slurry was heat-dried at 130 ° C. for 12 hours, the obtained granular solid was formed into tablets with a diameter of 5 mm and a height of 4 mm using a small molding machine, and then baked at 500 ° C. for 4 hours. The catalyst was obtained. The catalyst calculated from the charged raw materials was a complex oxide having the following atomic ratio.
Mo: Bi: Co: Ni: Fe: Na: B: K: Si = 12: 5: 2.5: 2.5: 0.4: 0.35: 0.2: 0.08: 24
The atomic ratio of _{a 1} and _{a 2} of molybdenum during catalyst preparation was 6.9 and 5.1 respectively.

<Example 1>
A multi-tubular reactor having 113 stainless steel reaction tubes having an inner diameter of 27.05 mm and a length of 3500 mm, each reaction tube, 622 ml of the composite oxide catalyst produced in Production Example 1 and 383 ml of inert balls (Chipton) And mixed. When one arbitrary reaction tube was selected from 113 reaction tubes, the catalyst layer length L of this reaction tube was 2.78 m, and the above formula (3)
The dilution rate of the catalyst layer indicated by is 60 vol% in the catalyst layer from the raw material supply part to 1/4 of the catalyst layer length in the gas flow direction, the next 1/4 is 40%, and the remaining 1/2 is 0%. Met. The fluid viscosity μ was 0.000022 kg / m / s, and the packed particle diameter d _p was 0.005 m.

Three of these reaction tubes were provided with an insertion tube having an outer diameter of 6.0 mm, and a thermocouple was installed in the insertion tube to measure the temperature in the reactor. Note that dibenzyltoluene was used as the heat medium.
Nitrogen 36.4 nm ³ / hr, air 69.7Nm ³ / hr, water vapor 22.5 Nm ³ / hr, and the feed gas having the composition shown in Table 1 was 13.2 nm ³ / hr feed, the preheater The mixture was heated to 318 ° C. as a mixed gas (mixed gas composition = nitrogen: 25.7 vol%, air: 49.1 vol%, water vapor: 15.9 vol%, source gas: 9.3 vol%). At this time, the flow rate of the mixed gas was 141.9 Nm ^{3 /} h (u: 0.60, G: 0.80) with respect to the catalyst amount of 0.135 Nm ³ , and the SV was 1050 h ⁻¹ . The differential pressure before and after the catalyst layer at this time was 16.4 kPa, and the porosity ε represented by the above formula (1) at the start of the oxidative dehydrogenation reaction was 0.336.

The product gas from the outlet of the reactor was quenched with water and cooled, and then analyzed by gas chromatography. The n-butene conversion rate (conversion rate in total of 1-butene, cis-2-butene, and trans-2-butene) was 91.6 mol%, and the butadiene selectivity was 89.0 mol%. After the reaction continued for 162 hours, the ε was 0.308, which was 91.7% compared to the porosity ε at the start of the oxidative dehydrogenation reaction. At this time, the supply amount of the raw material gas was gradually reduced, and after the supply amount was set to 0.0 Nm ³ / h over 4 hours, the temperature of the heat medium in the reactor was increased to 360 ° C. The amount of mixed gas at this time was 131.9 Nm ³ / h, and the difference in carbon dioxide concentration between the reactor inlet and outlet was 2730 ppm. After 35 hours had passed in this state, the difference between the carbon dioxide at the reactor inlet and outlet became 100 ppm or less, so the temperature of the reactor heat medium was reduced, and then the raw material gas was supplied to the reactor at 13.3 Nm ³ / hr. At this time, the mixed gas flow rate was 142.0 Nm ^{3 /} h (u: 0.000022, G: 0.80), the differential pressure was 15.7 kPa, and the ε was 0.336. Moreover, the n-butene conversion rate at this time (the conversion rate of 1-butene, cis-2-butene, and trans-2-butene in total) was 94.9 mol%, and the butadiene selectivity was 88.0 mol%. There was no change in the reaction before and after reducing the source gas. After confirming that there was no change in the reaction, the reactor was stopped after 189 hours.

<Comparative Example 1>
The oxidative dehydrogenation reaction was performed using a reactor similar to the reactor described in Example 1. Nitrogen was supplied at 29.9 Nm ³ / hr, air was supplied at 76.4 Nm ³ / hr, water vapor was supplied at 22.4 Nm ³ / hr, and a raw material gas having the composition shown in Table 1 was supplied at 13.3 Nm ³ / hr, inside the preheater The mixture was heated to 321 ° C. as a mixed gas (reactor introduction gas composition = nitrogen: 21.1 vol%, air: 53.8 vol%, water vapor: 15.8 vol%, source gas: 9.3 vol%). At this time, the flow rate of the mixed gas is 142.0 Nm ^{3 /} h (u: 0.000022, G: 0.80) with respect to the catalyst amount of 0.135 Nm ³ , and the flow rate of the mixed gas with respect to the catalyst amount in the reactor The ratio was 1050 h ⁻¹ . At this time, the differential pressure of the catalyst layer was 15.7 kPa, and the porosity ε was 0.338. As for the reaction results, the n-butene conversion rate (the conversion rate of 1-butene, cis-2-butene and trans-2-butene in total) was 94.9 mol%, and the butadiene selectivity was 88.0 mol%. After the reaction continued for 190 hours, the ε was 0.282, which was 83.4 % compared to the porosity ε at the start of the oxidative dehydrogenation reaction. At this time, after the source gas was reduced and then stopped, the temperature of the heating medium in the reactor was increased to 355 ° C. The amount of mixed gas at this time was 132.0 Nm ³ / h, and the difference in carbon dioxide concentration between the reactor inlet and outlet was 2110 ppm. After 35 hours had passed in this state, the difference between the carbon dioxide at the inlet and the outlet of the reactor became 100 ppm or less, so the temperature of the reactor heat medium was reduced, and then the raw material gas was supplied to the reactor at 11.6 Nm ³ / hr. The mixed gas flow rate at this time is 142.8 Nm ^{3 /} h (u: 0.000022, G: 0.80), the differential pressure is 18 kPa, and ε is 0.331, and the gap at the start of the oxidative dehydrogenation reaction It was 97.9 % compared with the rate ε. When the reaction was continued in this state, the pressure difference across the reactor suddenly increased, making it impossible to continue the operation. Further, after the operation was stopped, the reactor was opened and the catalyst in the reaction tube was extracted. As a result, 11% of the catalyst in the catalyst layer in the reactor was cracked.

When Example 1 is compared with Comparative Example 1, in Example 1, the porosity during the oxidative dehydrogenation reaction can be maintained in the range of 86 to 100 % at the start of the reaction, so that the reaction can be continued stably. While the crushing was prevented, in Comparative Example 1, the porosity deviated from the above range during the oxidative dehydrogenation reaction, so the reaction could not be stably maintained, and the catalyst was also crushed. Recognize.

DESCRIPTION OF SYMBOLS 1 Reactor 2 Quench tower 3, 6, 13 Cooler 4, 7, 14 Drain pot 8A, 8B Dehydrator 9 Heater 10 Solvent absorption tower 11 Deaeration tower 12 Solvent separation tower 101-121 Piping

Claims (5)

A raw material gas containing monoolefin having 4 or more carbon atoms and a molecular oxygen-containing gas are mixed and supplied to a fixed bed reactor having a catalyst layer, and the corresponding conjugated diene is produced by performing an oxidative dehydrogenation reaction. When obtaining the product gas containing the conjugated diene, the catalyst layer is provided with 2 to 5 catalyst layers having different dilution ratios , and the catalyst layer in the oxidative dehydrogenation reaction represented by the following formula (1) Is controlled in the range of 86 to 100% as compared with the porosity of the catalyst layer at the start of the oxidative dehydrogenation reaction.

However,

(In the above formula , Δp is the pressure difference [kPa] between the inlet and outlet of the reactor, L is the catalyst layer length [m], g _c is the gravitational acceleration [kg / m · s ² ], and μ is supplied to the reactor. Viscosity of mixed gas [kg / m /
s], u are the linear velocity of the mixed gas supplied to the reactor [m / s], d _p is the catalyst particle diameter [m], and G is the mass velocity of the mixed gas supplied to the reactor [kg / m ^2]. S]. )   The porosity ε of the catalyst layer is controlled by manipulating the supply amount of the source gas and the supply amount of the molecular oxygen-containing gas to the reactor, or the temperature in the reactor. A method for producing a conjugated diene as described in 1. above.   The method for producing a conjugated diene according to claim 1 or 2, wherein the catalyst is a composite oxide catalyst containing at least molybdenum, bismuth, and cobalt. The said catalyst is a complex oxide catalyst represented by following General formula (2), The manufacturing method of the conjugated diene of any one of Claims 1-3 characterized by the above-mentioned.
Mo _a Bi _b Co _c Ni _{d F e} X _f Y _g Z _h Si _i O _j (2)
Wherein X is at least one element selected from the group consisting of magnesium (Mg), calcium (Ca), zinc (Zn), cerium (Ce) and samarium (Sm), and Y is sodium (Na) , Potassium (K), rubidium (Rb), cesium (Cs) and at least one element selected from the group consisting of thallium (Tl), Z is boron (B), phosphorus (P), arsenic (As) And at least one element selected from the group consisting of tungsten (W), and a to j represent atomic ratios of the respective elements, and when a = 12, b = 0.1 to 7, c = 0 to 10, d = 0 to 10 (provided c + d = 1 to 10), e = 0.05 to 3, f = 0 to 2, g = 0.04 to 2, h = 0 to 3, i = 0 to 0 48, and j satisfies the oxidation state of other elements Numerical value is.)   The raw material gas is produced by dehydrogenation or oxidative dehydrogenation of gas containing 1-butene, cis-2-butene, trans-2-butene or a mixture thereof obtained by dimerization of ethylene, or a mixture thereof. The gas containing a butene fraction or a gas containing a large amount of hydrocarbons having 4 carbon atoms obtained when fluid oil cracking a heavy oil fraction, according to any one of claims 1 to 4, The manufacturing method of conjugated diene of description.

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