METHOD FOR PREPARING MONOCYCLIC AROMATIC COMPOUNDS AND LONG-CHAIN OLEFIN COMPOUNDS FROM CARBON DIOXIDE-RICH SYNTHESIS GAS

Abstract

Disclosed is a method for directly synthesizing monocyclic aromatic compounds and long-chain olefin compounds from a carbon dioxide-rich synthetic gas and, specifically, a method for directly synthesizing monocyclic aromatic compounds and long-chain olefin compounds from a carbon dioxide-rich synthetic gas, the method comprising a step of preparing a C.sub.1-C.sub.15 short-chain hydrocarbon by Fischer-Tropsch (FT) synthesis and a step of preparing monocyclic aromatic compounds and long-chain olefin compounds by dehydrogenating the short-chain hydrocarbon products, and maximizing the yield of the short-chain hydrocarbon by using, as a synthetic gas to be used in FT synthesis, a carbon dioxide-rich synthetic gas in which the molar ratio of hydrogen, carbon monoxide and carbon dioxide is delimited to a specific range, and maximizing the yield of the monocyclic aromatic compounds or the long-chain olefin compounds by specifying the composition of a catalyst to be used in the dehydrogenation and the temperature and pressure condition.

Claims

1. A method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds from a carbon dioxide-rich synthesis gas, which comprises: i) a step of preparing a hydrocarbon by conducting a Fischer-Tropsch (FT) synthesis process in the presence of an iron-based catalyst using a carbon dioxide-rich synthesis gas with a CO.sub.2/(CO+CO.sub.2) molar ratio controlled within a range from 0.4 to 0.65 as a source material; ii) a step of separating a C.sub.1-C.sub.15 short-chain hydrocarbon from the hydrocarbon products; and iii) a step of preparing C.sub.6-C.sub.12 monocyclic aromatic compounds and C.sub.6-C.sub.18 long-chain olefin compounds by dehydrogenating the C.sub.1-C.sub.15 short-chain hydrocarbon in the presence of a crystalline aluminosilicate-based catalyst, hydrogen and water.

2. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein the carbon dioxide-rich synthesis gas in the step i) has a H.sub.2/(2CO+3CO.sub.2) molar ratio of 0.85-1.1.

3. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein the iron-based catalyst in the step i) further comprises one or more promoter selected from a group consisting of copper (Cu), manganese (Mn), cobalt (Co), nickel (Ni), zinc (Zn), aluminum (Al) and potassium (K).

4. The method for synthesizing monocyclic aromatic compounds and a long-chain olefin compounds according to claim 1, wherein the step i) is performed under the condition of a reaction temperature of 250-350° C. and a reaction pressure of 10-30 bar.

5. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein, in the step ii), a C.sub.4-C.sub.15 short-chain hydrocarbon separated at a temperature of 0-200° C. based on a pressure of 20 bar is separated.

6. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein the step iii) is performed under the condition of a reaction temperature of 100-450° C. and a reaction pressure of 1-30 bar.

7. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 6, wherein the step iii) is performed under the condition of a reaction temperature of 300-350° C. and a reaction pressure of 5-20 bar to maximize the selectivity of the monocyclic aromatic compounds.

8. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 6, wherein the step iii) is performed under the condition of a reaction temperature of 200-270° C. and a reaction pressure of 10-20 bar or a reaction temperature of 300-350° C. and a reaction pressure of 1-2 bar to maximize the selectivity of the long-chain olefin compounds.

9. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein the crystalline aluminosilicate-based catalyst in the step iii) has a Si/Al molar ratio of 10-50 and comprises mesopores with a size of 10 nm or smaller and micropores with a size of 1-8 Å.

10. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 9, wherein the crystalline aluminosilicate-based catalyst in the step iii) has a Brønsted acid site/Lewis acid site ratio of 0.1-3.2.

11. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein the crystalline aluminosilicate-based catalyst in the step iii) further comprises one or more promoter metal element (A) selected from a group consisting of iron (Fe), copper (Cu), zinc (Zn), tin (Sn), nickel (Ni), molybdenum (Mo), potassium (K), lanthanum (La), platinum (Pt), gold (Au), palladium (Pd), rhodium (Rd), ruthenium (Ru) and silver (Ag).

12. The method for synthesizing a monocyclic aromatic compound and a long-chain olefin compound according to claim 11, wherein the A/Al molar ratio of the promoter metal element (A) based on the aluminum (Al) atom of the crystalline aluminosilicate-based catalyst is 0.01-2.5.

13. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein, in the step iii), the molar ratio of hydrogen is 0.1-10 and the molar ratio of water is 0.1-1.3 based on the short-chain hydrocarbon.

14. The method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds according to claim 1, wherein, in the step iii), a C.sub.1-C.sub.4 light hydrocarbon produced as a byproduct is recycled to a reforming process for preparing the synthesis gas.

Description

BRIEF DESCRIPTION OF DRAWINGS

[0026] FIG. 1 schematically describes a method for synthesizing monocyclic aromatic compounds and long-chain olefin compounds from a carbon dioxide-rich synthesis gas according to the present invention.

BEST MODE FOR CARRYING OUT INVENTION

[0027] Hereinafter, a synthesis method according to the present invention is described in detail referring to FIG. 1.

[0028] In the first step, a hydrocarbon is produced by conducting a Fischer-Tropsch (FT) synthesis process using a synthesis gas as a source material.

[0029] A general FT synthesis process is conducted by hydrogenation of carbon monoxide. However, in the present invention, carbon efficiency is maximized by an FT synthesis process in which carbon monoxide and carbon dioxide are hydrogenated at the same time.

Hydrogenation of carbon monoxide: nCO+2nH.sub.2.fwdarw.—(CH.sub.2).sub.n—+nH.sub.2O  (1)

Reverse water gas shift reaction: CO.sub.2+H.sub.2.fwdarw.CO+H.sub.2O  (2)

[0030] In the FT synthesis process of the present invention, a synthesis gas with a controlled molar ratio of hydrogen, carbon monoxide and carbon dioxide is used as a source material in the presence of an iron-based catalyst.

[0031] The iron-based catalyst is a catalyst commonly used in the FT synthesis process. The iron-based catalyst may further contain one or more promoter selected from a group consisting of copper (Cu), manganese (Mn), cobalt (Co), nickel (Ni), zinc (Zn), aluminum (Al) and potassium (K), if necessary.

[0032] The synthesis gas used as the source material in the present invention is a carbon dioxide-rich synthesis gas which contains hydrogen, carbon monoxide and carbon dioxide and has a CO.sub.2/(CO+CO.sub.2) molar ratio of 0.4-0.65 and a H.sub.2/(2CO+3CO.sub.2) molar ratio of 0.85-1.1. The synthesis gas may be produced through a reforming process of natural gas. Methods for reforming the natural gas may include steam reforming, carbon dioxide reforming, combined reforming, partial oxidation, etc. In particular, a synthesis gas produced by the combined reforming method which allows control of the composition of the synthesis gas may be used.

[0033] In the synthesis gas used in the present invention, if the CO.sub.2/(CO+CO.sub.2) molar ratio is smaller than 0.4, the maximum carbon efficiency desired by the present invention cannot be achieved because the reverse water gas shift reaction whereby carbon dioxide is converted to carbon monoxide does not occur sufficiently. In addition, the catalyst lifetime may decrease due to increased carbon deposition. On the contrary, if the CO.sub.2/(CO+CO.sub.2) molar ratio exceeds 0.65, productivity may decrease because both CO and CO.sub.2 conversion decreased.

[0034] And, if the H.sub.2/(2CO+3CO.sub.2) molar ratio in the synthesis gas is smaller than 0.85, the maximum carbon efficiency desired by the present invention cannot be achieved because the reverse water gas shift reaction whereby carbon dioxide is converted to carbon monoxide does not occur sufficiently. In addition, the catalyst life time may decrease due to increased carbon deposition. On the contrary, if the H.sub.2/(2CO+3CO.sub.2) molar ratio exceeds 1.1, the yield of the long-chain olefin compounds may decrease because the selectivity of methane (C.sub.1) is increased unnecessarily and, as a result, the selectivity of the short-chain olefin is increased during the following dehydrogenation of the hydrocarbon.

[0035] The FT synthesis process of the present invention is performed under the condition of a reaction temperature of 250-350° C. and a reaction pressure of 10-30 bar, specifically under the condition of a reaction temperature of 300-350° C. and a reaction pressure of 15-25 bar.

[0036] If the reaction temperature in the FT synthesis process is below 250° C., the maximum carbon efficiency desired by the present invention cannot be achieved because the reverse water gas shift reaction whereby carbon dioxide is converted to carbon monoxide does not occur sufficiently. In addition, the yield of the long-chain olefin compounds may decrease because the selectivity of methane (C.sub.1) is increased unnecessarily. On the contrary, if the reaction temperature exceeds 350° C., the catalyst life time may decrease due to increased carbon deposition and the yield of the long-chain olefin compounds may decrease because the selectivity of methane (C.sub.1) is increased.

[0037] If the reaction pressure in the FT synthesis process is below 10 bar, the maximum carbon efficiency desired by the present invention cannot be achieved because the reverse water gas shift reaction does not occur sufficiently. In addition, the yield of the long-chain olefin compounds may decrease because the selectivity of methane (C.sub.1) is increased and the catalyst life time may decrease due to increased carbon deposition. On the contrary, if the reaction pressure exceeds 30 bar, the productivity of the monocyclic aromatic compounds and the long-chain olefin compounds decreases due to decreased selectivity of the short-chain hydrocarbon which is a reactant in the following dehydrogenation of the hydrocarbon. In addition, the operation of the process for maintaining the high pressure is complicated.

[0038] In the second step, a C.sub.1-C.sub.15 short-chain hydrocarbon is separated from the hydrocarbon produced in the FT synthesis process.

[0039] Because the reaction condition of the FT synthesis process of the present invention is controlled to improve the selectivity of the short-chain hydrocarbon, the selectivity of the C.sub.1-C.sub.15 short-chain hydrocarbon can be increased to 80-100%. Accordingly, the preparation method of the present invention achieves the effect of simplifying the process because an additional process for upgrading the hydrocarbon produced from the FT synthesis process is not necessary.

[0040] In the second step of the present invention, the C.sub.1-C.sub.15 short-chain hydrocarbon is separated and purified through simple separation in order to further increase the selectivity of the monocyclic aromatic compounds and the long-chain olefin compounds produced in the following dehydrogenation process.

[0041] The separation of the short-chain hydrocarbon is achieved through a separation process. Assuming that the pressure inside a reactor is 20 bar, the separation temperature may be maintained at 0-200° C., specifically 120-160° C.

[0042] If the separation temperature is below 0° C., the selectivity of the monocyclic aromatic compounds and the long-chain olefin compounds may decrease and the separator may be damaged due to freezing of water which is a reaction byproduct. On the contrary, if the separation temperature exceeds 200° C., the catalyst life time may decrease due to deposition of the long-chain hydrocarbon on the catalyst in the following dehydrogenation process.

[0043] In the third step, the monocyclic aromatic compounds and the long-chain olefin compounds are produced by dehydrogenating the short-chain hydrocarbon.

[0044] The dehydrogenation reaction consists of a step of reacting the short-chain hydrocarbon with hydrogen (H.sub.2) and water (H.sub.2O) in the presence of a crystalline aluminosilicate-based catalyst.

[0045] Based on the short-chain hydrocarbon used as the reactant, the hydrogen (H.sub.2) is used with a molar ratio of 0.1-10, specifically 1-8, and the water (H.sub.2O) is used with a molar ratio of 0.1-1.3, specifically 0.1-0.8.

[0046] If the hydrogen/short-chain hydrocarbon molar ratio during the dehydrogenation reaction is below 0.1 or exceeds 10, the selectivity of the monocyclic aromatic compounds and the long-chain olefin compounds may decrease because the dehydrogenation reaction does not occur sufficiently.

[0047] If the water/short-chain hydrocarbon molar ratio is below 0.1, the aluminosilicate-based catalyst may be deactivated quickly because it is difficult to remove carbon deposition with the water. And, if the water/short-chain hydrocarbon molar ratio exceeds 1.3, the catalyst may be deactivated quickly because an excess amount of water is adsorbed to the aluminosilicate-based catalyst.

[0048] As the catalyst used in the dehydrogenation reaction, a crystalline aluminosilicate-based catalyst with a Si/Al molar ratio of 10-50 is used. If the Si/Al molar ratio is below 10, the productivity of the polycyclic aromatic compounds is high because the dehydrogenation reaction proceeds vigorously. On the contrary, if the Si/Al molar ratio exceeds 50, the productivity of the monocyclic aromatic compounds decreases because the chain growth reaction becomes dominant.

[0049] Also, as the crystalline aluminosilicate-based catalyst, one with a Brønsted acid site/Lewis acid site ratio controlled to 0.1-3.2 is used. If the acid site ratio is below 0.1, the productivity of the polycyclic aromatic compounds is high because the dehydrogenation reaction proceeds vigorously. On the contrary, if the acid site ratio exceeds 3.2, the productivity of the monocyclic aromatic compounds decreases because the chain growth reaction becomes dominant.

[0050] Also, as the crystalline aluminosilicate-based catalyst, a crystalline porous catalyst containing mesopores with a size of 10 nm or smaller and micropores with a size of 1-8 Å. If the pore size of the crystalline catalyst does not satisfy the above range, the productivity of the monocyclic aromatic compounds decreases.

[0051] Although the crystalline aluminosilicate-based catalyst may be used alone, the crystalline aluminosilicate-based catalyst may further contain one or more promoter selected from a group consisting of iron (Fe), copper (Cu), zinc (Zn), tin (Sn), nickel (Ni), molybdenum (Mo), potassium (K), lanthanum (La), platinum (Pt), gold (Au), palladium (Pd), rhodium (Rd), ruthenium (Ru) and silver (Ag), if necessary. Specifically, the content of the metal element (A) contained as the promoter may be maintained such that the A/Al molar ratio based on the aluminum (Al) atom of the crystalline aluminosilicate-based catalyst is 0.01-2.5. Specifically, the molar ratio of the promoter metal element (A) based on the aluminum atom (i.e., A/Al molar ratio) may be 0.1-1.

[0052] The dehydrogenation process of the present invention is performed under the condition of a reaction temperature of 100-450° C. and a reaction pressure of 1-30 bar. By changing the reaction temperature and the reaction pressure of the dehydrogenation process, the selectivity of the monocyclic aromatic compounds or the long-chain olefin compounds can be controlled.

[0053] That is to say, in order to increase the selectivity of the monocyclic aromatic compounds, it is recommended to maintain the reaction temperature high and maintain the reaction pressure low within the above-described ranges of the reaction temperature and the reaction pressure. Specifically, the selectivity of the monocyclic aromatic compounds may be maximized by maintaining the reaction temperature at 300-350° C. and the reaction pressure at 5-20 bar.

[0054] And, in order to increase the selectivity of the long-chain olefin compounds, it is recommended to maintain the reaction temperature low and maintain the reaction pressure high within the above-described ranges of the reaction temperature and the reaction pressure. Specifically, the selectivity of the C.sub.6-C.sub.18 long-chain olefin compounds may be maximized by maintaining the reaction temperature at 200-270° C. and the reaction pressure at 10-20 bar or by maintaining the reaction temperature at 300-350° C. and the reaction pressure at 1-2 bar.

[0055] Accordingly, in the dehydrogenation process of the present invention, the yield of the monocyclic aromatic compounds or the C.sub.6-C.sub.18 long-chain olefin compounds may be maximized by optimizing the composition of the catalyst, the reaction temperature and the reaction pressure.

[0056] Among the hydrocarbons produced from the dehydrogenation process, light hydrocarbons (C.sub.1-C.sub.4) are included as byproducts in addition to the monocyclic aromatic compounds and the long-chain olefin compounds. In the present invention, the light hydrocarbons (C.sub.1-C.sub.4) in gas state and the monocyclic aromatic compounds and the long-chain olefin compounds in liquid state can be separated through a simple separation process.

[0057] Specifically, the separation temperature for the gas/liquid separation may be −5 to 5° C. If the temperature is below −5° C., the separator may be damaged due to freezing of water which is a reaction byproduct. And, if the temperature exceeds 5° C., the light hydrocarbon (C.sub.1-C.sub.4) may not be easily separated from the liquid hydrocarbon (C.sub.5+).

[0058] The C.sub.1-C.sub.4 light hydrocarbon separated through the gas/liquid separation may be recycled to a reforming reactor for preparing the synthesis gas.

[0059] The synthesis method of the present invention described above will be described in more detail referring to the following examples. However, the present invention is not limited by them.

BEST MODE FOR CARRYING OUT INVENTION

Examples

Example 1. Fischer-Tropsch Synthesis Processes with Different CO.SUB.2./(CO+CO.SUB.2.) Ratios

[0060] Hydrocarbons were produced by performing a Fischer-Tropsch synthesis process under the condition of a reaction temperature of 300° C. and a reaction pressure of 10 bar by loading 1 g of an iron-based catalyst with a composition of 100Fe-6Cu-16Al-4K in a ½-inch stainless steel fixed-bed reactor and supplying a synthesis gas at a flow rate of 1,800 mL/g cat.Math.h. The hydrocarbon products were analyzed by on-line GC (TCD, FID).

[0061] The result of analyzing the selectivity of the hydrocarbon products depending on the composition of the synthesis gas used as a source material is shown in Table 1.

TABLE-US-00001 TABLE 1 C.sub.2-C.sub.4 hydrocarbon Molar ratio in distribution synthesis gas Conversion (%) Carbon (mol %) CO.sub.2/(CO + H.sub.2/(2CO + CO .fwdarw. CO .fwdarw. CO.sub.2 .fwdarw. deposition C.sub.2-C.sub.4 CO.sub.2) 3CO.sub.2) HC CO.sub.2 HC ratio (%) C.sub.2-C.sub.4 olefin Example 0.25 1 96.8 19.7 — 5.9 32.3 72.7 1-1 Example 0.4 1 95.5 — 8.7 5.4 31.2 73.1 1-2 Example 0.5 1 95.0 — 12.6 4.7 30.1 73.7 1-3 Example 0.6 1 94.1 — 17.6 4 29.4 74.2 1-4 Example 0.75 1 88.4 — 29.7 3.1 28.0 75.1 1-5

[0062] According to the present invention, the Fischer-Tropsch (FT) process and the reverse water gas shift reaction proceeded well, the conversion of CO and CO.sub.2 and the C.sub.2-C.sub.4 hydrocarbon selectivity were high and the carbon deposition ratio was low when the CO.sub.2/(CO+CO.sub.2) molar ratio was 0.4-0.65 and the H.sub.2/(2CO+3CO.sub.2) molar ratio was 1.

[0063] In contrast, when the CO.sub.2/(CO+CO.sub.2) molar ratio was low as 0.25, the reverse water gas shift reaction did not occur although the carbon monoxide (CO) was converted to the C.sub.2-C.sub.4 hydrocarbon with a high conversion. And, when the CO.sub.2/(CO+CO.sub.2) molar ratio was 0.75, the C.sub.2-C.sub.4 hydrocarbon selectivity was low because the CO conversion was decreased in the Fischer-Tropsch (FT) process.

Example 2. Fischer-Tropsch Synthesis Processes with Different H.SUB.2./(2CO+3CO.SUB.2.) Values

[0064] Fischer-Tropsch synthesis was conducted under the same condition as in Example 1 by fixing the CO.sub.2/(CO+CO.sub.2) value to 0.5 and varying the H.sub.2/(2CO+3CO.sub.2) value.

[0065] The result of analyzing the selectivity of the hydrocarbon products depending on the composition of the synthesis gas used as a source material is shown in Table 2.

TABLE-US-00002 TABLE 2 Molar ratio in synthesis gas CO .fwdarw. HC CO.sub.2/(CO + H.sub.2/(2CO + conversion Selectivity (%) CO.sub.2) 3CO.sub.2) (%) Hydrocarbon CO.sub.2 Example 0.5 0.4 86.2 64.7 29.4 2-1 Example 0.5 1 94.5 100.0 −10.7 2-2

[0066] As seen from Table 2, when the H.sub.2/(2CO+3CO.sub.2) molar ratio was 1, the hydrocarbon selectivity was 100%, the CO conversion was high and CO.sub.2 was consumed. In contrast, when the H.sub.2/(2CO+3CO.sub.2) molar ratio was low as 0.4, the CO conversion and the hydrocarbon selectivity were low and CO.sub.2 was produced.

Example 3. Dehydrogenation Processes with Different Catalyst Conditions

[0067] Hydrocarbons produced by performing a Fischer-Tropsch synthesis process under the condition of a reaction temperature of 320° C. and a reaction pressure of 20 bar by loading 1 g of an iron-based catalyst with a composition of 100Fe-6Cu-16Al-4K in a ½-inch stainless steel fixed-bed reactor and supplying a synthesis gas with CO.sub.2/(CO+CO.sub.2)=0.5 and H.sub.2/(2CO+3CO.sub.2)=1 at a flow rate of 1,800 mL/g cat.Math.h were dehydrogenated. Before dehydrogenating the hydrocarbon products, C.sub.1-C.sub.15 short-chain hydrocarbons were separated through a separation process. The temperature and pressure inside the separator were maintained at 136° C. and 20 bar.

[0068] For the dehydrogenation, monocyclic aromatic compounds and long-chain olefin compounds were produced under the condition of a reaction temperature of 300° C. and a reaction pressure of 10 bar by loading 0.6 g of a crystalline aluminosilicate-based catalyst in a ½-inch stainless steel fixed-bed reactor. The composition of the products was analyzed by on-line GC (TCD, FID) and GC/MS.

[0069] The result of analyzing the composition of the hydrocarbon products depending on the composition of the catalyst used in the dehydrogenation process is shown in Table 3.

TABLE-US-00003 TABLE 3 Dehydrogenation condition Promoter (A) Product distribution (mol %) A/Al C.sub.6-C.sub.18 Main molar C.sub.6-C.sub.18 long- paraffin Other catalyst Promoter ratio BTEX.sup.1) chain olefins isomers byproducts Example H- — — 10.2 28.7 19.5 41.6 3-1 ZSM5.sup.2) Example H- — — 19.9 22.3 13.1 44.7 3-2 ZSM5.sup.3) Example H- — — 22.4 1.5 9.7 66.4 3-3 ZSM5.sup.4) Example H- Zn 0.3 14.8 4.2 19.2 61.8 3-4 ZSM5.sup.4) Example H- Zn 0.5 16.3 3.0 23.8 56.9 3-5 ZSM5.sup.4) .sup.1)BTEX: benzene, toluene, ethylbenzene and xylene .sup.2)H-ZSM5: Si/Al = 40, specific surface area = 454.6 m.sup.2/g, Ratio of Brønsted to Lewis acid site = 1.21 .sup.3)H-ZSM5: Si/Al = 25, specific surface area = 421.3 m.sup.2/g, Ratio of Brønsted to Lewis acid site = 1.35 .sup.4)H-ZSM5: Si/Al = 15, specific surface area = 402.1 m.sup.2/g, Ratio of Brønsted to Lewis acid site = 0.81

[0070] As seen from Table 3, when only the crystalline aluminosilicate-based catalyst was used, the selectivity of monocyclic aromatic compounds was higher as the Si/Al ratio was smaller and the selectivity of long-chain olefin compounds was higher as the Si/Al ratio was larger.

[0071] And, when zinc (Zn) was included as a promoter while using the crystalline aluminosilicate-based catalyst as a main catalyst, the selectivity of paraffin isomers was increased in addition to those of the monocyclic aromatic compounds and the long-chain olefin compounds. This effect was enhanced as the Zn/Al ratio was increased.

Example 4. Dehydrogenation Processes with Different Temperature and Pressure Conditions

[0072] Hydrocarbon produced by performing a Fischer-Tropsch synthesis process was dehydrogenated in the same manner as in Example 3. Before performing the dehydrogenation process, C.sub.1-C.sub.15 short-chain hydrocarbons were separated through a separation process.

[0073] For the dehydrogenation, monocyclic aromatic compounds and long-chain olefin compounds were produced by loading 0.6 g of a crystalline aluminosilicate-based catalyst in a fixed-bed reactor. The composition of the products was analyzed by on-line GC (TCD, FID) and GC/MS.

[0074] The result of analyzing the composition of the hydrocarbon products depending on the reaction temperature and reaction pressure conditions of the dehydrogenation process is shown in Table 4.

TABLE-US-00004 TABLE 4 Dehydrogenation condition Reaction Reaction Product distribution (mol %) temperature pressure C.sub.6-C.sub.18 long- Other (° C.) (bar) Catalyst BTEX.sup.1) chain olefins byproducts Example 300 1 H-ZSM5.sup.2) 3 6.5 90.5 4-1 Example 300 5 H-ZSM5.sup.2) 22.0 1.9 76.1 4-2 Example 300 10 H-ZSM5.sup.2) 22.4 0.0 77.6 4-3 Example 300 20 H-ZSM5.sup.2) 18.4 0.6 81.0 4-4 Example 250 20 H-ZSM5.sup.2) 1.8 8.7 89.5 4-5 Example 300 10 Zn/H- 16.3 1.5 82.2 4-6 ZSM5.sup.3) Example 350 10 Zn/H- 18.5 0.0 81.5 4-7 ZSM5.sup.3) Example 400 10 Zn/H- 13.9 1.6 84.5 4-8 ZSM5.sup.3) Example 450 10 Zn/H- 11.6 2.1 86.3 4-9 ZSM5.sup.3) .sup.1)BTEX: benzene, toluene, ethylbenzene and xylene .sup.2)H-ZSM5: Si/Al = 15, specific surface area = 402.1 m.sup.2/g, Brønsted acid point = 0.81 .sup.3)Zn/H-ZSM5: Si/Al = 15, Zn/Al = 0.5

[0075] As seen from Table 4, the selectivity of monocyclic aromatic compounds and long-chain olefin compounds in the products was changed by the reaction temperature and reaction pressure conditions of the dehydrogenation process. That is to say, when the reaction temperature was 300° C., the selectivity of monocyclic aromatic compounds was increased as the reaction pressure decreased below 10 bar. And, the selectivity of long-chain olefin compounds was the highest when the reaction condition was 250° C. and 20 bar.

[0076] Those skilled in the art will appreciate that the conceptions and specific embodiments disclosed in the foregoing description may be readily utilized as a basis for modifying or designing other embodiments for carrying out the same purposes of the present invention. Those skilled in the art will also appreciate that such equivalent embodiments do not depart from the spirit and scope of the invention as set forth in the appended claims.