PROCESS FOR PRODUCING LPG AND BTX FROM MIXED HYDROCARBONS FEED

Abstract

The present invention relates to a process for producing LPG and BTX from a mixed feedstream comprising C5-C12 hydrocarbons by contacting said feedstream in the presence of hydrogen with a first hydrocracking catalyst and contacting the thus obtained first hydrocracked product in the presence of hydrogen with a second hydrocracking catalyst to produce a second hydrocracked product stream comprising LPG and BTX.

Claims

1. A process for producing LPG and BTX comprising: (a) contacting a feedstream comprising C5-C12 hydrocarbons and bicyclic hydrocarbons in the presence of hydrogen with a first hydrocracking catalyst to produce a first hydrocracked product stream, wherein the first hydrocracking catalyst comprises 0.01-2 wt-% hydrogenation metal in relation to the total catalyst weight and an aluminosilicate zeolite having a pore size of 6-8 Å and a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 1-150; and (b) contacting the first hydrocracked product stream in the presence of hydrogen with a second hydrocracking catalyst to produce a second hydrocracked product stream comprising LPG and BTX, wherein the second hydrocracking catalyst comprises 0.01-1 wt-% hydrogenation metal in relation to the total catalyst weight and an aluminosilicate zeolite having a pore size of 5-6 Å and a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 5-200, wherein step (a) is performed at a temperature of 325-425° C., a pressure of 10-4000 kPa gauge, a Weight Hourly Space Velocity (WHSV) of 0.5-24 h.sup.−1 and a hydrogen:hydrocarbon (H.sub.2:HC) molar ratio of 1-4; and step (b) is performed at a temperature of 325-500° C., a pressure of 10-4000 kPa gauge, a WHSV of 0.5-20 .sup.−1 and a H.sub.2:HC molar ratio of 1-4.

2. The process according to claim 1, wherein the hydrogenation metal of the first hydrocracking catalyst is at least one element selected from Group 10 of the periodic table of Elements, preferably Pd.

3. The process according to claim 1, wherein the aluminosilicate zeolite comprised in the first hydrocracking catalyst is selected from the group consisting of zeolite Y, faujasite (FAU), beta zeolite (BEA), and chabazite (CHA).

4. The process according to claim 1, wherein the hydrogenation metal of the second hydrocracking catalyst is at least one element selected from Group 10 of the periodic table of Elements.

5. The process according to claim 1, wherein the aluminosilicate zeolite comprised in the catalyst used in the second hydrocracking step (b) is selected from the group consisting of MCM-22 (MWW), mordenite (MOR) and ZSM-5 (MFI).

6. The process according to claim 1, wherein the feedstream comprises straight run naphtha, hydrocracked gasoline, light coker naphtha and coke oven light oil, FCC gasoline or mixtures thereof.

7. The process according to claim 1, wherein steps (a) and (b) are performed in a single reactor.

8. The process according to claim 7, wherein the reactor has a first catalyst layer comprising the first hydrocracking catalyst and a second catalyst layer comprising the second hydrocracking catalyst, wherein a space or an inert layer is present between the first catalyst layer and the second catalyst layer.

9. The process according to claim 7, wherein the reactor has a first catalyst layer comprising the first hydrocracking catalyst and a second catalyst layer comprising the second hydrocracking catalyst, wherein the first catalyst layer is in contact with the second catalyst layer.

10. The process according to claim 1, wherein the first hydrocracking catalyst is a mixture of the zeolite and the hydrogenation metal on a support of an amorphous alumina.

11. he process according to claim 1, wherein the first hydrocracking catalyst comprises the hydrogenation metal on a support of the zeolite.

12. The process according to claim 1, wherein the second hydrocracking catalyst is a mixture of the zeolite and the hydrogenation metal on a support of an amorphous alumina.

13. The process according to claim 1, wherein the second hydrocracking catalyst comprises the hydrogenation metal on a support of the zeolite.

14. The process according claim 1, further comprising subjecting the second hydrocracked product stream to separation to provide a C6+ stream and contacting the thus obtained C6+ stream in the presence of hydrogen with a third hydrocracking catalyst at temperature of 325-600° C., a pressure of 10-4000 kPa gauge and a Weight Hourly Space Velocity of 0.1-30 h.sup.−1 to produce a third hydrocracked product stream comprising LPG and BTX, wherein the third hydrocracking catalyst comprises 0.01-1 wt-% hydrogenation metal in relation to the total catalyst weight and a zeolite having a pore size of 5-6 Å and SiO.sub.2/Al.sub.2O.sub.3 ratio of 5-200.

15. A fixed bed reactor comprising, in this order, (i) an inlet, (ii) a first reaction zone comprising a first hydrocracking catalyst comprising 0.01-2 wt-% hydrogenation metal in relation to the total catalyst weight and an aluminosilicate zeolite having a pore size of 6-8 Å and a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 1-150, (iii) a second reaction zone comprising a second hydrocracking catalyst comprising 0.01-1 wt-% of the hydrogenation metal in relation to the total catalyst weight and an aluminosilicate zeolite having a pore size of 5-6 Å and a SiO.sub.2/Al.sub.2O.sub.3 molar ratio of 5-200 and (iv) an outlet.

Description

EXAMPLE

[0083] Aspen simulations have been performed at different temperatures (at 1379 kPa (200 psig) and H.sub.2:HC ratio of 3) to assess the effect induced by this parameter on the extent with which decalin is dehydrogenated in a hydrocracking reaction train. The results obtained from these simulations are summarized in Table 1 showing important differences in the extent with which tetralin is dehydrogenated at 375° C. and 450° C. Despite the fact that the concentration of decalin is low in both cases, the relative concentration of tetralin and naphthalene is strongly influenced by temperature conditions. Specifically, at 450° C. 80% of the decalin occurring in the feed is transformed into naphthalene, whereas at 375° C. the concentration of naphthalene is less than half of the amount of decalin present in the feed.

TABLE-US-00001 TABLE 1 Aspen simulations, performed using the Peng Robinson thermodynamic equation of state to assess the extent with which decalin is dehydrogenated to naphthalene at different temperature conditions. Feed Outlet 1 Outlet 2 Temperature (° C.) 375 375 450 Pressure (kPa) 1379 1379 1379 Mass Flow (kg/h) 360.75 360.75 360.75 Mole Flow (kmol/h) Cis-decalin 2.5 0.02 <0.001 Trans-decalin 0.156 <0.001 Naphthalene 1.14 2 Tetralin 1.19 0.193 Hydrogen 7.5 16.8 19.61

[0084] These differences are of outmost importance to the invention given that aromatic molecules (such as naphthalene) cannot be ring opened without been partially hydrogenated (i.e., tetralin). Consequently, these results are informative of the fact that in order to maximize the ring contraction and ring opening reactions of the di-ring structures present in the feedstream comprising C5-C12 hydrocarbons and polyaromatics, the pre-layer of wide pore zeolite and hydrogenation catalyst should operate at the lower possible hydrocracker temperatures, namely at the beginning of the series of hydrocracking reactors, as depicted in FIG. 1. FIG. 1 describes a series of 4 subsequent reactors (2), (6), (10), and (13), wherein in reactor (2) the hydrocarbon feed (1) is subjected to the first hydrocracking step (a) followed by the second hydrocracking step (b) as described herein in reactor beds (3), respectively (4). The product of reactor (2) is fed to reactor (6) via line (5) and is again subjected to the first hydrocracking step (a) followed by the second hydrocracking step (b) as described herein in reactor beds (7), respectively (8). The product of reactor (10) is fed to reactor (13) via line (12) and is again subjected to the third hydrocracking step as described herein in reactor bed (14), producing the process product (15). Optionally, the process of the present invention comprises a separation step between reactor (6) and reactor (10) to provide the C6+ stream that is subjected to the third hydrocracking step.

[0085] To corroborate the above described aspen simulations experiments were performed with a naphtha feed containing increasing amounts of decalin (up to 2.5 wt %) at 1379 kPa (200 psig), WHSV of 2.6 h.sup.−1, H.sub.2:HC ratio of 3 and 450° C. The results obtained from these experiments are shown in FIG. 2 and the feed composition details summarized in Table 2. These results indicate that under the selected reaction conditions all the decalin present in the feed is transformed into tetralin and naphthalene. Moreover, the relative concentration of naphthalene and tetraline is similar to that obtained in the simulations.

TABLE-US-00002 TABLE 2 Naphtha feed composition Naphtha feed (wt-%) Butane 1.54 Methyl Butane 3.37 Pentane 5.16 UK2 0.49 2-methylpentane 9.57 3 methylpentane 6.11 Hexane 13.89 Methylcyclopentane 5.24 2.4 Dimethylpentane 0.12 Benzene 1.79 Cyclohexane 5.62 trans 1-3 Dmcp 1.12 1-3 Dmcp 3.74 1-2 Dmcp 1.03 2,2,4 Trimethylpentane 1.95 Heptane 5.12 Methylcyclohexane 4.15 Ethylcyclopentane 0.58 2,3,3 Trimethylpentane 0.00 Toluene 3.96 Octane 4.31 Ethyl cyclohexane 0.91 Ethylbenzene 1.06 Meta/Para 1.11 OX 0.74 Nonane 3.00 iso Propyl Benzene 0.26 Propyl Benzene 0.29 1-Methyl-3-Ethylbenzene 0.61 1-Methyl-4-Ethylbenzene 0.00 1,3,5-Trimethylbenzene 0.25 1M2E Benzene 0.15 Pseudocumene 1.01 Indane 0.12 Indene 0.03 1,3 De Benzene 0.01 Butyl Benzene 0.07 2ethanyl 1,4 Dm Benzene 0.00 1,4 Dm 2 E Benzene 0.01 1M Indene 0.01 Naphthalene 0.00 1,3,5 Triethylbenzene 0.00 Methyl Naphthalene (1) 0.00 Methyl Naphthalene (2) 0.00