JP2016056132A - Method for producing 1,2-difluoroethylene - Google Patents

Abstract

PROBLEM TO BE SOLVED: To provide a method for producing 1,2-difluoroethylene efficiently and economically in an industrially practicable method.SOLUTION: A method for producing 1,2-difluoroethylene includes the reaction between 1-chloro-1,2-difluoroethylene and hydrogen in a gas phase in the presence of a hydrogenated catalyst.SELECTED DRAWING: None

Description

  The present invention relates to a method for producing 1,2-difluoroethylene.

  In recent years, due to the increase in greenhouse gas regulations, an alternative chlorofluorocarbon having a lower global warming potential (GWP) has been demanded, and 1,2-difluoroethylene (HFO-1132) has attracted attention. In the present specification, for halogenated hydrocarbons, an abbreviation of the compound is described in parentheses after the compound name, but in the present specification, the abbreviation may be used instead of the compound name as necessary.

As a method for producing the above HFO-1132, for example, 1,2-dichloro-1,2-difluoroethylene (CFO-1112) is obtained by the hydrodechlorination reaction in the presence of a hydrogenation catalyst as follows. A method based on the reaction formula (3) is known (see Patent Document 1).

  However, in order to obtain HFO-1132 by hydrodechlorination reaction of CFO-1112, two atoms of chlorine atoms in the raw material CFO-1112 are replaced with hydrogen atoms, so hydrogen is supplied to the excess amount reactor. There was a need to do. Therefore, it has been a problem that the reactor size becomes large. In addition, for example, the selectivity of HFO-1132 as the target product shown in Patent Document 1 is about 55%, which is not high.

Furthermore, when chlorine is replaced, hydrogen chloride is by-produced, and 2 mol of hydrogen chloride is produced when 1 mol of HFO-1132 is produced. .
Thus, the present situation is still unknown about the method of manufacturing refrigerant | coolant HFO-1132 efficiently industrially.

JP 62-30730 A

  An object of the present invention is to provide a method for industrially and efficiently producing 1,2-difluoroethylene (HFO-1132).

  The present invention relates to a process for producing 1,2-difluoroethylene (HFO-1132) in which 1-chloro-1,2-difluoroethylene (HCFO-1122a) and hydrogen are reacted by contacting them with a hydrogenation catalyst in a gas phase. I will provide a.

  According to the production method of the present invention, 1,2-difluoroethylene (HFO-1132) can be produced industrially efficiently. In addition, according to the method of the present invention, compared with the theoretical equivalent when HFO-1132 is obtained by hydrodechlorination of CFO-1112 represented by the reaction formula (3) of the above-mentioned known technique, it is supplied to the reactor. There is an advantage in that the total amount of raw material gases to be reduced is reduced to 2/3. Further, the amount of hydrogen chloride generated is also advantageous in that it is halved compared to the method according to reaction formula (3).

Embodiments of the present invention will be described below.
The method for producing HFO-1132 of the present invention is a hydrodechlorination reaction of HCFO-1122a represented by the following reaction formula (1).

  HCFO-1122a is a compound known to be used as a raw material for producing a fluorine-containing compound or an intermediate. For example, as shown in Zhurnal Organicheskoi Khimii, 18 (5), 938-45; 1982, dehydrochlorination of 1,2-dichloro-1,2-difluoroethane (HCFC-132) using an organic alkaline reagent The method of making it known is known.

  For example, HCFO-1122a can also be produced by dehydrochlorination by bringing HCFC-132 into contact with an alkaline aqueous solution in the presence of a phase transfer catalyst.

  Furthermore, it is possible to use HCFO-1122a obtained together with HFO-1132 by hydrodechlorination reaction of CFO-1112 shown in Patent Document 1. HCFO-1122a is not limited to what was manufactured by said each method, but can be used.

  In the production method of the present invention, the method for reacting HCFO-1122a and hydrogen in the gas phase in the presence of a hydrogenation catalyst is not particularly limited. The catalytic reaction is performed by supplying HCFO-1122a and hydrogen into a reaction field in which a hydrogenation catalyst is disposed, usually into a reactor containing the hydrogenation catalyst. In the following description, HCFO-1122a and hydrogen are in a gas phase unless otherwise specified. The reaction field will be described as a reactor containing a hydrogenation catalyst.

  The production method of the present invention may be a continuous production method or a batch production method. In a continuous production method, supply to a reactor containing a hydrogenation catalyst of HCFO-1122a and hydrogen, contact between a raw material and a hydrogenation catalyst in the reactor, and the reaction of HFO-1132 and HCl The removal from the vessel is continuously performed.

  In batch production, either HCFO-1122a and the hydrogen supply may be first or simultaneous. That is, when one of HCFO-1122a and hydrogen is supplied, even if the other is not supplied into the reactor, the components supplied earlier are supplied later while staying in the reactor. The components are supplied, and HCFO-1122a and hydrogen may be in contact with the hydrogenation catalyst for a predetermined time in the reactor.

  The production method of the present invention is preferably a continuous method in terms of production efficiency. Hereinafter, unless otherwise specified, an embodiment in which the method of the present invention is applied to continuous production will be described.

  When performing the manufacturing method of this invention by a continuous type, the supply of HCFO-1122a and hydrogen to a reactor may be performed separately, and may be mixed beforehand. In the production method of the present invention, the theoretical equivalent of hydrogen required for HCFO-1122a according to the above reaction formula (1) is 1 equivalent.

  In the method of the present invention, the ratio of HCFO-1122a and hydrogen supplied to the reactor is preferably a ratio of 0.01 to 10.0 moles of hydrogen with respect to 1 mole of HCFO-1122a. That is, the molar ratio of the hydrogen supply amount to the supply amount of HCFO-1122a supplied to the reactor (when the supply molar amount of HCFO-1122a and the supply molar amount of hydrogen are represented by “HCFO-1122a” and “hydrogen”, respectively) “Hydrogen / HCFO-1122a” is preferably 0.01 to 10.0.

Further, from the viewpoint of increasing the conversion rate of the raw material components, the hydrogen / HCFO-1122a is preferably 0.1 or more, more preferably 0.3 or more, and particularly preferably 0.5 or more. From the viewpoint of suppressing the generation of by-products, the hydrogen / HCFO-1122a is preferably 4.0 or less, more preferably 2.0 or less, and particularly preferably 1.0 or less.
In the present embodiment employing a continuous production method, hydrogen / HCFO-1122a (molar ratio) is represented by a molar flow rate ratio per unit time.

  The hydrogenation catalyst used in the method of the present invention is not particularly limited as long as it has a function of catalyzing the hydrodechlorination reaction according to the reaction formula (1). Specifically, a catalyst containing palladium such as palladium alone, an alloy containing palladium, or a mixture of palladium and a metal other than palladium is preferable as the hydrogenation catalyst. Among these catalysts, a mixture of palladium and a metal other than palladium is advantageous in terms of durability. A hydrogenation catalyst may be used individually by 1 type, and may use 2 or more types together.

  As an alloy containing palladium, an alloy having a palladium content of usually 50 atomic% or more, preferably 80 atomic% or more is preferable, and platinum and / or rhodium are preferable as a metal combined with palladium.

  As a metal other than palladium (hereinafter, also simply referred to as “other metal”) in a mixture of palladium and a metal other than palladium, a Group 8 element such as iron, ruthenium, osmium, a Group 9 element, For example, cobalt, rhodium, iridium, a Group 10 element such as nickel, platinum, and gold can be used. Another metal may be used individually by 1 type and may be used in combination of 2 or more type. As for the ratio of palladium and another metal, 0.01-50 mass parts is preferable with respect to 100 mass parts of palladium.

  The hydrogenation catalyst is preferably used by being supported on a carrier. When two or more components are used in combination as a hydrogenation catalyst, a mixture in which two or more components are mixed in advance may be supported on a carrier, or each component may be supported on another carrier. May be used. In this specification, the hydrogenation catalyst is composed of two or more mixtures, for example, a mixture of palladium and another metal. In the form of supporting on a carrier, palladium and the other metal are separated from each other. And a mode in which it is supported and mixed.

  The carrier for supporting the hydrogenation catalyst is not particularly limited as long as the carrier can sufficiently support the catalyst. Specific examples include activated carbon and metal oxides. Examples of the metal oxide include alumina, zirconia, and silica. Among these, activated carbon is preferable from the viewpoint of activity, durability, and reaction selectivity.

  Examples of the activated carbon include activated carbon obtained from plant materials such as wood, charcoal, fruit husk and coconut shell, and mineral raw materials such as peat, lignite and coal. As the carrier for supporting the hydrogenation catalyst, activated carbon obtained from plant raw materials is preferable from the viewpoint of catalyst durability, and coconut shell activated carbon is particularly preferable. Examples of the shape of the activated carbon include formed charcoal having a length of about 2 to 10 mm, crushed charcoal of about 4 to 50 mesh, granular charcoal, and the like, from the point of activity, 4 to 20 mesh of crushed charcoal, or length 2 to 2 5 mm of charcoal is preferred.

  When the hydrogenation catalyst is supported on the carrier, the amount of the hydrogenation catalyst supported is preferably 0.1 to 10 parts by mass, more preferably 0.5 to 2 parts by mass with respect to 100 parts by mass of the carrier. When the supported amount of the hydrogenation catalyst is 0.1 parts by mass or more, the conversion rate in the reaction between the raw material compound HCFO-1122a and hydrogen is sufficiently improved. If the loading amount of the hydrogenation catalyst is 10 parts by mass or less, it is easy to suppress an excessive temperature rise of the hydrogenation catalyst due to reaction heat, and it is easy to reduce the production of by-products. As a method for supporting the hydrogenation catalyst on the carrier, a conventionally known method may be used.

  The reactor for reacting HCFO-1122a with hydrogen is resistant to the temperature and pressure described below, and is a hydrogenation catalyst, preferably a carrier carrying a hydrogenation catalyst (hereinafter also referred to as “catalyst-carrying carrier”). .) Is not particularly limited as long as it can accommodate the catalyst layer and form a catalyst layer.

  Specifically, a known reactor for performing a gas phase reaction using a catalyst can be used without particular limitation, and examples thereof include a cylindrical vertical reactor. Examples of the material of the reactor include glass, iron, nickel, iron or an alloy containing nickel as a main component. The reactor is preferably provided with means capable of adjusting the temperature inside the reactor. For example, a double tube structure provided with a jacket capable of circulating a heat medium or a refrigerant outside the reactor may be used, and heating means such as an electric heater may be provided outside or inside the reactor.

  A hydrogenation catalyst, preferably a catalyst support, is accommodated in such a reactor, and a catalyst layer is formed as a reaction field. The hydrogenation catalyst, preferably the catalyst support, may be accommodated in either a fixed bed type or a fluidized bed type. In addition, in the case of a fixed bed type, it may be either a horizontal fixed bed type or a vertical fixed bed type, but in a mixed gas composed of multiple components, the concentration distribution of each component is generated depending on the location due to the difference in specific gravity. It is preferable to use a vertical fixed floor type.

When using a catalyst-carrying carrier, a catalyst layer is formed in the reactor by filling the catalyst-carrying carrier in a predetermined region in the reactor. Packing density of the catalyst-supporting carrier in the catalyst layer is preferably ^{0.5~1g / cm 3, 0.6~0.8g / cm} 3 is more preferable. If the packing density of the catalyst-carrying carrier is 0.5 g / cm ³ or more, the filling amount of the catalyst-carrying carrier per unit volume is large and the amount of gas to be reacted can be increased, so that productivity is improved. If the packing density of the catalyst-carrying support is 1 g / cm ³ or less, it is easy to suppress an excessive temperature rise of the catalyst layer due to reaction heat, and it is easy to reduce the production of by-products. There may be one packed portion of the catalyst-supporting support in the reactor, or two or more.

  When HCFO-1122a and hydrogen supplied to the reactor are mixed in advance, the HCFO-1122a is mixed as it is in the vicinity of the inlet of the reactor as a raw material mixed gas as it is when the supply is separate from the inlet. It circulates through the catalyst layer in the outlet direction. The gas flowing through the reactor is a raw material mixed gas on the upstream side of the catalyst layer and a reaction product gas on the downstream side. The temperature and pressure of HCFO-1122a and hydrogen supplied to the reactor are preferably the same temperature and pressure, for example, 0 to 200 ° C. or 0 to 2 MPa, whether they are separate or mixed.

  Moreover, in order to suppress an excessive temperature rise and catalyst deterioration, the reaction may be carried out by diluting with an inert gas such as nitrogen.

  In order to cause HCFO-1122a and hydrogen to react in the gas phase, the catalyst layer temperature is set to a temperature equal to or higher than the dew point of the raw material mixed gas composed of HCFO-1122a and hydrogen. Here, in this specification, the catalyst layer temperature is the temperature of the catalyst layer that maintains the temperature using the heating means attached to the reactor. Moreover, since the reaction rate of the raw material tends to decrease when the catalyst layer temperature is low from the viewpoint of the reaction rate of the raw material, the temperature at which the catalyst layer is maintained is preferably 40 ° C. or higher in order to carry out the reaction efficiently. More preferably, it is 60 ° C. or higher, and most preferably 70 ° C. or higher. Further, since the durability of the catalyst is lowered at a high temperature, the catalyst layer temperature is desirably maintained at 400 ° C. or lower, preferably 350 ° C. or lower. In the present specification, the catalyst layer temperature is also referred to as an initial reaction temperature.

  When the raw material mixed gas supplied to the catalyst layer reacts, reaction heat is generated. Since the raw material mixed gas reacts in a partial region of the catalyst layer, a partial region of the catalyst layer has a higher temperature than the other catalyst layer regions. The region in which the raw material mixed gas reacts is called the reaction zone, and the maximum temperature in the reaction zone is called the reaction zone temperature. In the present invention, the reaction zone temperature of the catalyst layer is set to the reaction temperature of the supplied raw material mixed gas.

  Since the catalytic activity of the reaction zone decreases with time, the reaction zone usually moves gradually from the inlet of the raw material mixed gas to the downstream side in the gas flow direction. Therefore, in the initial stage of the operation, the reaction zone temperature (reaction temperature) is set such that the measurement part of the insertion-type thermometer is positioned on the inlet side of the catalyst layer, and the measurement part is moved to the outlet side as the reaction proceeds. It is desirable to measure it.

  In addition, as described above, the catalytic activity decreases with time, and as the reaction zone moves to the outlet side of the catalyst layer, the reaction zone temperature decreases and the reaction rate of the raw material gas decreases. It is desirable to maintain the reaction zone temperature by suppressing the decrease in the reaction rate of the raw material gas. For example, the reaction zone temperature can be maintained by gradually increasing the catalyst layer temperature using the heating means.

  The reaction zone temperature, that is, the reaction temperature is preferably 180 ° C. or higher, more preferably 200 ° C. or higher, from the viewpoint of the conversion rate of the raw material. Moreover, 200 degreeC or more is preferable from a viewpoint of a selectivity, 220 degreeC or more is more preferable, and 240 degreeC or more is further more preferable. In view of catalyst durability, the reaction zone temperature (reaction temperature) is preferably 400 ° C. or less, more preferably 370 ° C. or less, and particularly preferably 350 ° C. or less. As a combination of the upper and lower limits, 180 to 400 ° C is preferable, 200 to 370 ° C is more preferable, and 250 to 350 ° C is particularly preferable.

  In the above reaction, the reaction zone temperature changes due to the load on the reactor depending on the amount of raw material consumption and the change in the catalyst activity described above. Therefore, in order to maintain the temperature of the reaction zone within the above range, it is desirable to adjust the catalyst layer temperature using the heating means in accordance with the load on the reactor and the catalyst activity depending on the raw material consumption.

  In the embodiment of the present invention, the pressure in the reactor during the hydrodechlorination reaction of HCFO-1122a is preferably 0 to 2.0 MPa, and more preferably 0 to 1.0 MPa in terms of gauge pressure.

  The contact time (T) between the HCFO-1122a and the hydrogenation catalyst of the raw material mixed gas comprising hydrogen is represented by the following formula (A), and is generally preferably 4 to 120 seconds, and 15 seconds from the viewpoint of the conversion rate of the raw material. The above is more preferable, and 30 seconds or more is particularly preferable. From the viewpoint of reactor efficiency, the contact time is more preferably 90 seconds or shorter, and particularly preferably 60 seconds or shorter. This contact time is calculated from the amount of gas introduced into the reactor and the volume of the catalyst layer. In the present invention, the contact time (T) is the reaction time of the raw material mixed gas.

Contact time (T) = (W / 100) × V _G / V (A)
Wherein (A), W is the concentration of the raw material mixed gas in the total gas flowing through the catalyst layer (mol%), V _G is an a total gas flowing through the catalyst layer flow (mL / sec) , V is the volume (mL) of the catalyst layer.

The linear velocity (u) in the catalyst layer of the raw material compound to be reacted with hydrogen in the hydrodechlorination reaction, HCFO-1122a in the present invention, is represented by the following formula (B).
Linear velocity (u) = (W / 100) × V _G / S (B)
Wherein (B), W is the concentration of the raw material compound gas in the total gas flowing through the catalyst layer (mol%), V _G is the total gas flowing through the catalyst layer at a flow rate (cm 3 ^/ sec) Yes, S is the cross-sectional area (cm ² ) of the catalyst layer with respect to the gas flow direction.

  The linear velocity u in the catalyst layer of HCFO-1122a is preferably 0.1 to 100 cm / second, and more preferably 1 to 30 cm / second. When the linear velocity u of HCFO-1122a is 0.1 cm / second or more, productivity is improved. If the linear velocity u of HCFO-1122a is 100 cm / sec or less, the conversion rate in the reaction between HCFO-1122a and hydrogen is improved.

  Here, the catalyst layer that has been continuously used as described above and whose catalytic activity of the entire layer from the upstream side to the downstream side of the catalyst layer has decreased is replaced with a new one or regenerated by a known regeneration technique. Used. Thereby, it is possible to maintain the catalyst activity of the catalyst layer in the reactor at a predetermined level.

  The reaction product gas discharged from the outlet of the reactor contains various by-products and HCl in addition to the target substance HFO-1132. In order to remove HCl, the reaction product gas is usually washed with an alkali and then subjected to a dehydration treatment. In the present specification, the gas after de-HCl obtained in this way is referred to as outlet gas.

  In the hydrodechlorination reaction of HCFO-1122a, a composition containing HFO-1132 can be obtained as the outlet gas. As compounds other than HFO-1132 contained in the outlet gas, in addition to HCFO-1122a which is an unreacted raw material, 1,2-difluoroethane (HFC-152), fluoroethane (HFC-161), methane, 1, 1-difluoroethylene (HFO-1132a, VdF), 1,1-difluoroethane (HFC-152a), 1-chloro-1,1-difluoroethane (HCFC-142b), 1,2-dichloro-1,2-difluoroethane ( HCFC-132), fluoroethylene (HFO-1141) and the like.

  The above components other than HFO-1132 contained in the outlet gas can be removed to a desired extent by known means such as distillation. The separated HCFO-1122a can be recycled as part of the raw material.

  EXAMPLES The present invention will be specifically described below with reference to examples, but the present invention is not limited to these.

  The conversion rate described below is when the molar composition of HCFO-1122a relative to each component derived from HCFO-1122a calculated from the GC area obtained by gas chromatography analysis of the reaction gas is X% ( 100-X)% is referred to as the conversion of HCFO-1122a, and the selectivity of each component refers to the percentage of each of the reacted HCFO-1122a that has been converted to each component other than HCFO-1122a. . The selectivity of each component is calculated by the molar composition of each component / conversion rate of HCFO-1122a. The conversion rate and selectivity of CFO-1112 in the comparative example are calculated similarly.

(Example 1)
A catalyst made of stainless steel having an inner diameter of 2.3 cm and a length of 50 cm is filled with palladium-supported activated carbon in which 0.5 parts by mass of palladium is supported on 100 parts by mass of coconut shell activated carbon, and a catalyst having a height of 40 cm A layer was formed. The catalyst layer volume was 166 mL. The packing density of the palladium-supported activated carbon in the catalyst layer was 0.74 g / cm ³ .

  The catalyst layer in the reactor thus formed was set to an initial reaction temperature of 80 ° C. by an electric heater, an internal pressure (gauge pressure) of 0.04 MPa, a raw material composition consisting of HCFO-1122a and hydrogen (hereinafter referred to as raw material (I )) / HCFO-1122a (molar ratio) was set to 1.0 and supplied to the reactor at a flow rate of 0.445 mol / hr. Hereinafter, all the pressures are assumed to be gauge pressures. The reaction proceeded and the temperature in the reaction zone of the catalyst layer (reaction temperature) was 226 ° C. The temperature in the reaction zone was measured by inserting a thermocouple in the catalyst layer. Further, the contact time (T) (reaction time) of the raw material (I) with the catalyst layer was 60 seconds from the flow rate of the raw material (I) and the volume of the catalyst layer.

  The reaction product gas discharged from the outlet of the reactor was subjected to dehydration treatment after alkali washing, and then analyzed by gas chromatography to calculate the molar composition of gas components contained in the outlet gas. Moreover, the conversion rate of HCFO-1122a and the selectivity of each component were calculated | required based on the molar composition of exit gas. Moreover, the production amount of HFO-1132 was calculated from the obtained conversion rate of HCFO-1122a and the selectivity of HFO-1132.

(Example 2)
The same operation as in Example 1 was performed except that the hydrogen / HCFO-1122a (molar ratio) in the raw material (I) was changed to 0.5. In Example 2, the temperature (reaction temperature) in the reaction zone of the catalyst layer was 224 ° C.

(Example 3)
The same operation as in Example 1 was performed except that the hydrogen / HCFO-1122a (molar ratio) in the raw material (I) was changed to 2.0. In Example 3, the temperature (reaction temperature) in the reaction zone of the catalyst layer was 231 ° C.

Example 4
The same operation as in Example 1 was performed except that the catalyst layer in the reactor was changed to 130 ° C. with an electric heater. In Example 4, the temperature (reaction temperature) in the reaction zone of the catalyst layer was 281 ° C.

(Example 5)
The same operation as in Example 1 was performed except that the raw material (I) having a hydrogen / HCFO-1122a (molar ratio) of 1.0 was changed to a flow rate of 0.890 mol / hr. In Example 5, the temperature (reaction temperature) in the reaction zone of the catalyst layer was 243 ° C. Further, the contact time (T) (reaction time) with the catalyst layer of the raw material (I) was set to 30 seconds from the flow rate of the raw material (I) and the volume of the catalyst layer. Table 1 shows production conditions and outlet gas analysis results in each example.

  In Table 1 above, the amount of HCl by-product is the amount of HCl by-produced when converted from HCFO-1122a to HFO-1132, methane, VdF, HFO-1141, HFC-152, and HFC-161. , HCFO-1122a conversion rate × sum of selectivity of each component.

  The HCl by-product unit in Table 1 above is the mass of HCl produced as a by-product when 1 g of HFO-1132 is produced in each example. The smaller the value, the smaller the HCl by-product that needs to be processed. It means that there is little raw quantity.

(Comparative Example) Hydrogen Reduction of CFO-1112 Nickel cylindrical reaction having an inner diameter of 2.0 cm and a volume of 160 mL filled with palladium-supported activated carbon supporting 2.0 parts by mass of palladium with respect to 100 parts by mass of granular activated carbon. The temperature of the vessel was adjusted to 350 ° C., and the hydrogen / HCFO-1112 (molar ratio) of the raw material composition (hereinafter also referred to as raw material (II)) composed of CFO-1112 and hydrogen was set to 5.0 under atmospheric pressure. The reactor was fed at a flow rate of 0.5 mol / hr. The organic vapor discharged from the outlet of the reactor was analyzed by gas chromatography, and the molar composition of the gas component contained in the outlet gas was calculated. The contact time (T) of the raw material (II) with the catalyst layer was 51 seconds. The results are shown in Table 2.

  According to the production method of the present invention, HFO-1132 useful as a refrigerant can be produced efficiently and economically by an industrially feasible method.

Claims (8)

  A process for producing 1,2-difluoroethylene, comprising reacting 1-chloro-1,2-difluoroethylene and hydrogen in the gas phase in the presence of a hydrogenation catalyst.   2. The method for producing 1,2-difluoroethylene according to claim 1, wherein the hydrogenation catalyst is at least one selected from the group consisting of simple palladium, an alloy containing palladium, and a mixture of palladium and a metal other than palladium. .   The method for producing 1,2-difluoroethylene according to claim 1 or 2, wherein the hydrogenation catalyst is supported on a carrier, and the carrier is at least one selected from the group consisting of activated carbon, alumina, zirconia and silica.   The method for producing 1,2-difluoroethylene according to claim 3, wherein the activated carbon is activated carbon obtained from a plant raw material and / or activated carbon obtained from a mineral raw material.   The method for producing 1,2-difluoroethylene according to claim 3 or 4, wherein an amount of the hydrogenation catalyst supported on 100 parts by mass of the carrier is 0.05 to 10 parts by mass.   The ratio of 1-chloro-1,2-difluoroethylene to hydrogen subjected to the reaction is 0.1 to 10.0 moles of hydrogen with respect to 1 mole of 1-chloro-1,2-difluoroethylene. The manufacturing method of the 1, 2- difluoroethylene as described in any one of 1-5.   The method for producing 1,2-difluoroethylene according to any one of claims 1 to 6, wherein a reaction temperature between the 1-chloro-1,2-difluoroethylene and hydrogen is 180 ° C to 400 ° C.   The method for producing 1,2-difluoroethylene according to any one of claims 1 to 7, wherein a reaction time between the 1-chloro-1,2-difluoroethylene and hydrogen is 4 to 120 seconds.