Process and catalyst system for conversion of C6 aromatics to higher aromatics

Abstract

A process for the upgrading of hydrocarbon streams, i.e., processing any hydrocarbon feed streams rich in benzene and sulphur compounds. The process for simultaneous hydrodesulfurization and benzene conversion to higher alkylated aromatic molecules (C.sub.7 to C.sub.10 aromatics), without need of prior treatment like distillation, or sulfur removal. The hydrocarbon feed streams are processed over sulfided metal catalyst impregnated on acid support simultaneously desulfurizes and alkylates the benzene molecules.

Claims

1. A process for controlling a percentage of benzene in a hydrocarbon feed stream comprising benzene, the process comprising: mixing the hydrocarbon feed stream with an alkylating agent and a sulfiding agent to obtain a mixed hydrocarbon feed stream; and routing the mixed hydrocarbon feed stream and hydrogen gas to pass over a catalyst in a catalytic reactor to obtain a processed hydrocarbon stream, wherein the catalyst comprises 15-25 wt % of group VIB metals and 4-7 wt % of group VIIIB metals, and wherein the process converts benzene into higher alkylated aromatic molecules and simultaneously results in hydrodesulfurization of the hydrocarbon feed stream in the catalytic reactor.

2. The process as claimed in claim 1, wherein the benzene content in the hydrocarbon feed stream is 5 vol %.

3. The process as claimed in claim 1, wherein the hydrocarbon feed stream comprises a sulphur compound.

4. The process as claimed in claim 1, wherein the catalytic reactor is selected from the group consisting of a fixed bed plug flow reactor, a continuous stirred tank reactor, a batch reactor, and a semi batch reactor.

5. The process as claimed in claim 1, wherein the catalytic reactor is maintained at a weighted average bed temperature (WABT) of catalyst bed of 350-500 C.

6. The process as claimed in claim 1, wherein the catalytic reactor is maintained at a hydrogen partial pressure of 15-100 kg/cm.sup.2 g.

7. The process as claimed in claim 1, wherein the hydrogen gas to hydrocarbon ratio (H.sub.2/HC) is in a range of 300-1500 Nm.sup.3/m.sup.3.

8. The process as claimed in claim 1, wherein the catalyst is a dual functional catalyst with an acidic and a hydrogenation-dehydrogenation function, wherein the acidic function is imparted by virtue of a zeolite support and the hydrogenation-dehydrogenation function is imparted by metals selected from the group VIB and the group VIIIB.

9. The process as claimed in claim 1, wherein the catalyst comprises: a) a carrier comprising Y-zeolite; b) a binder comprising alumina; c) metals selected from the groups VIB and VIIIB; and d) an additive containing nitrogen and oxygen.

10. The process as claimed in claim 9, wherein the group VIB metals are selected from the group consisting of molybdenum, tungsten, and salts and mixtures thereof, and the group VIIIB metals are selected from the group consisting of nickel, cobalt, and salts and mixtures thereof.

11. The process as claimed in claim 1, wherein the hydrocarbon feed stream is not desulfurized; and the nitrogen content in the hydrocarbon feed stream is below 20 ppmw.

12. The process as claimed in claim 1, wherein the alkylating agent is selected from the group consisting of olefins and alkyl-electrophiles generating agents, wherein the alkyl-electrophiles generating agents comprise a functional group of hydroxyls, halides, thiols, or oxy; and wherein the sulfiding agent is selected from the group consisting of H.sub.2S and Dimethyl Disulfide (DMDS).

13. The process as claimed in claim 1, further comprising generating the alkylating agent in-situ using compound(s) that-react within the catalytic reactor.

14. The process as claimed in claim 1, wherein the processed hydrocarbon stream has a RON at of least 2 units more than the hydrocarbon feed stream; and the benzene content in the processed hydrocarbon stream is 0.5 vol %.

15. The process as claimed in claim 1 comprising: a. routing the hydrocarbon feed stream with the alkylating agent, the sulfiding agent, and the hydrogen gas to the catalytic reactor to pass over the catalyst to produce a catalytic reactor effluent; b. routing the catalytic reactor effluent to a high-pressure separator (HPS) and routing water settled at a bottom of the HPS to a cold high-pressure separator (CHPS); c. routing an effluent hydrocarbon from the HPS to a low-pressure separator (LPS), d. routing and mixing condensed hydrocarbons of the CHPS with the effluent hydrocarbon from the; and e. separating water from the CHPS and obtaining a product from the LPS, wherein CHPS is operated at a lower temperature relative to the HPS and the LPS is operated at a lower pressure and a temperature relative to the HPS.

16. The process as claimed as claimed 15, wherein an unutilized hydrogen gas is separated in the HPS and removed and recycled back to the catalytic reactor.

17. The process as claimed as claimed 15, wherein off gases are separated from the LPS, and from the CHPS.

Description

BRIEF DESCRIPTION OF THE DRAWING

(1) To further clarify advantages and aspects of the invention, a more particular description of the invention will be rendered by reference to specific embodiments thereof, which is illustrated in the appended drawing(s). It is appreciated that the drawing(s) of the present invention depicts only typical embodiments of the invention and are therefore not to be considered limiting of its scope.

(2) FIG. 1: illustrates schematic diagram of process flow scheme of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

(3) For promoting and understanding of the principles covered by present invention, reference will now be made to the specific embodiments of the present invention further illustrated in the drawings and specific language will be used to describe the same. The foregoing general description and the following detailed description are explanatory of the present disclosure and are not intended to be restrictive thereof. It will nevertheless be understood that no limitation of the scope of the present disclosure is thereby intended, such alterations and further modifications in the illustrated composition, and such further applications of the principles of the present disclosure as illustrated herein being contemplated as would normally occur to one skilled in the art to which the present disclosure relates. Unless otherwise defined, all technical and scientific terms used herein have the same meaning as commonly understood by one of ordinarily skilled in the art to which this present disclosure belongs. The methods, and examples provided herein are illustrative only and not intended to be limiting.

(4) The main embodiment of the present invention provides a process for conversion of benzene molecules available in any refinery stream to alkylated benzene molecules thereby reducing the concentration of benzene in the stream without much alteration in the total aromatic content of that stream.

(5) Any hydrocarbon feed stream containing benzene molecules along with alkylating agent, sulfiding agent, and hydrogen gas are feed to a fixed bed catalytic reactor. These streams can be selected from the streams forming part of the gasoline pool, such as FCC gasolines, coker gasolines, reformates, straight run naphtha's, hydrocracker naphtha's, etc. The streams going to gasoline pool have benzene concentration up to 8 vol % as in reformates or can be in the range 0.5 to 2 vol % as in FCC and coker gasolines or can be 0.8 to 8 vol % as in straight run naphtha or in the range of 1 to 5 vol % as in the hydrocracker naphtha.

(6) The benzene concentration in the total gasoline pool is to be maintained below 1 vol %, therefore, the feed streams need to be processed to reduce their benzene concentrations. Alternatively, there are many feed streams which are being sent to aromatics maximization or petrochemicals production. These feed streams may contain benzene concentration as low as 30 wt % along with other aromatic compounds. It is preferable to reduce the benzene concentrations in these feed streams going for petrochemicals production, but not by saturating them but it is desirable to alkylate them.

(7) In one of the embodiment the present invention, the feed stream is any hydrocarbon stream generated through any refinery processes with boiling range C.sub.5-210 C., preferably C.sub.5 to 160 C. and most preferably C.sub.5 to 95 C.

(8) In another embodiment of the present invention, the benzene concentration in feed stream should be at least 5 vol % and preferably more than 5 vol %.

(9) In yet another embodiment of present invention, the C.sub.6 fraction is used as feed stream. Further, the process has no limitation with respect to maximum benzene concentration of the feed stream.

(10) In another embodiment of the present invention, the feed stream along with alkylating agent and hydrogen is passed over catalyst system present in a catalytic reactor. The catalytic reactor is selected from fixed bed plug flow reactor, continuous stirred tank reactor, batch reactor or semi batch reactor. In one of the embodiment the fixed bed catalytic reactor is a fixed bed plug plow reactor (PFR). The reactor and the catalyst system are maintained at a predetermined temperature and under predetermined hydrogen pressure.

(11) In the process of present invention, suitable alkylating agents known in the art may be used. However, the primary alkylating agents used are olefins. Use of olefins as alkylating agents have advantage of achieving alkylation of target molecule (i.e., alkane or aromatic) in a single step without the need of metal function in the catalyst. But the process requires very strong acidity, which makes the process cumbersome due to acid handling and storage issues. Other alkylating agent are selected from heteroatom containing compounds. Therefore, the alkylating agent in the present invention can be any alkyl-electrophiles generating agent from the functional group of hydroxyls, halides, thiols, oxy, but not limited to. The alkylating agent can also be generated in-situ using compound(s) that may react within the reactor system to generate alkyl-electrophiles.

(12) In another embodiment of the present invention, the reactor and the catalyst system are maintained at the temperature suitable for caring out the reaction i.e., weighted average bed temperature (WABT) of catalyst bed is preferably maintained between 350 and 500 C., more preferably between 350 to 450 C. and most preferably between 380 and 430 C.

(13) In one of the embodiment of the present invention, the benzene conversion reaction is carried out in hydrogen environment. The hydrogen partial pressure maintained in the reactor is preferably between 15-100 kg/cm.sup.2 g more preferably between 25-75 kg/cm.sup.2g and most preferably between 40-65 kg/cm.sup.2g. Maintaining the hydrogen pressure in the reactor suppress coke formation and therefore reduces catalyst deactivation rate and increase catalyst life. The aromatic hydrogenation reactions are equilibrium controlled and the equilibrium is favored towards hydrogenation reaction at high pressure and low temperature. Since the operating conditions in the alkylation reactor is entirely opposite hence aromatic saturation reactions are not favored.

(14) In yet another embodiment of the present invention, the weighted hourly space velocity (WHSV) of the process is preferably between of 0.5 and 5 h.sup.1, more preferably between 0.7 and 3 h.sup.1 and most preferably between 1 and 2 h.sup.1 and the hydrogen to hydrocarbon ratio (H.sub.2/HC) is preferably between 300 and 1500 Nm.sup.3/m.sup.3, more preferably between 500 and 1200 Nm.sup.3/m.sup.3 and most preferably between 600 and 1000 Nm.sup.3/m.sup.3.

(15) The catalyst of the present invention is dual functional and has both acidic and metal sites. The acidic function is imparted by virtue zeolite support and the hydrogenation-dehydrogenation function is imparted by metals selected from the group of VIB and VIIIB of periodic table. In the base catalyst the metal is impregnated over the support and present in oxide form. The metal site is active in elemental (zero-valent) form of sulfide form. Accordingly, the catalyst is activated either by reduction of metal site to elemental form using any suitable reducing agent or by converting the metal site to sulfide form using any suitable sulfiding agent.

(16) In one of the embodiment of the present invention, the catalyst is comprises of the followings: a) a carrier comprising of Y-zeolite b) a binder comprising of alumina c) metals selected from the groups VIB and VIIIB of periodic table d) additive containing nitrogen and oxygen.

(17) The Y-zeolite acts as the main acid function in the catalyst. The acid function is determined by the number of active sites and its strength. The extent of alkylation reaction highly depends upon the active site per unit area of the catalyst. However, active sites present in a particular type of zeolite are compensated by the strength of the available acid sites.

(18) In another embodiment of the present invention, the alkylation function of the catalyst is promoted by the presence of the metal sites. The metals are selected from group VIB and VIIIB of the periodic table. The metals from group VIB are selected from a group consisting of molybdenum, tungsten, and salts and mixtures thereof. The source of tungsten comprises of salts selected from a group consisting of ammonium tungsten, tungsten trioxide, sodium tungstate, tungstic acid, phospo tungstic acid etc. The group VIIIB metals are selected from a group consisting of nickel, cobalt, and salts and mixtures thereof. The source of nickel salts is selected from a group consisting of nickel nitrate, nickel sulfate, nickel carbonate, nickel acetate, nickel chloride etc.

(19) Yet another embodiment of the present invention the catalyst comprises of 15 to 25 wt % of the metal from VIB and metal from group VIIIB in the range 4 to 7 wt % of the total catalyst present. The metal site in the catalyst is activated to elemental form it is termed as reduced catalyst. The catalyst is activated to reduced catalyst using any suitable reducing agent known in the art and the most prominent reducing agent available in refinery and petrochemical complexes is hydrogen. Hydrogen is used as reducing agent, and the reduction is carried out preferably between 400 and 600 C. and most preferably between 450 and 550 C. Depending upon the exotherm, the purity and the flow rate of hydrogen gas can be controlled.

(20) In another embodiment of the present invention, the metal site is activated to sulfide form, the catalyst is termed as sulfide catalyst. The use of sulfide catalyst is very common to the many refinery applications for hydro-desulfurization and hydrocracking processes. The metallic site from oxide-state is changed to sulfide-state by the use of sulfiding agents known in the art viz. H.sub.2S, DMDS, etc.

(21) In one of the embodiment of the present invention, the initial activity of catalyst in reduced form as well as in sulfide form is similar, however; the coke formation and deactivation rate of reduced catalyst is much greater compared to sulfided catalyst. Thus, the sulfide catalysts are more compatible for refinery-based benzene alkylation process. Further, with use of reduced catalyst the pressure drop starts developing within 4-5 days run and finally the feed stream needs to be cut off due to excessive pressure drop across the catalyst bed, whereas no such observation recorded with sulfided catalyst even after the continuous run of one month, which is good enough to understand the stability of sulfide catalyst.

(22) Another, disadvantage of reduced catalyst is with respect to feed impurity. The reduced catalyst demand feed stream free from any hetero atom impurities, particularly sulphur and nitrogen. However, for sulfide catalyst the only impurity is Nitrogen. The nitrogen compound in the feed reacts with the acid site and deactivate them temporarily.

(23) In another embodiment of the present invention, no prior desulfurization of feed stream is required. It is further disclosed that the process accommodates feed streams containing sulphur between 100 and 5000 ppmw, preferably between 100 and 2000 ppmw and more preferably between 100 and 500 ppmw. If the sulphur content in the feed stream is less than 100 ppmw, DMDS or any H.sub.2S generating agents can be doped along with the feed stream for maintaining sufficient H.sub.2S concentration in the reactor system for keeping the metal in sulfided form. Maintaining H.sub.2S concentration is vital for sulfided catalyst, otherwise catalyst will be permanently deactivated.

(24) In another embodiment of the present invention, the nitrogen content of the hydrocarbon feed stream must be below 20 ppmw. If the feed stream contains N more than 20 ppmw, pretreatment i.e., Hydro-denitrogentation (HDN) of feed is required. The HDN reaction is also associated with hydro-desulfurization (HDS) reaction, and both can be done together in a separate reactor prior to benzene conversion reactor. However, if the feed stream contains N below 20 ppmw and the sulphur as mentioned in previous paragraphs, then no pretreatment is required and both HDS and the benzene conversion reaction will occur simultaneously in the alkylation reactor. This flexibility of process and catalyst design makes this innovation unique and easily implementable in refinery for benzene management.

(25) In another embodiment of the present invention, the operating condition and the catalyst system for alkylation reactor is also favorable for hydrocracking as well as isomerization reactions. However, the hydrocracking rate is low for low boiling fractions. Since the feed steam is boiling between C.sub.5 and 210 C., more preferably between C.sub.5 to 160 C. and most preferably C.sub.5 to 95 C., hence the cracking reaction will not be predominant. It is also known in the art that, the hydrocracking reaction rate is highest for aromatic and lowest for n-paraffin, hence cracking of aromatic side chains will be the more compared to hydrocracking of other molecules viz. naphthenes, iso-paraffins, n-paraffins and olefins. However, the isomerization of n-paraffins will enhance overall RON of the product stream.

(26) In another embodiment of the present invention, the alkylation of other aromatics viz. C.sub.7, C.sub.8, C.sub.9 and C.sub.10 aromatics occurs simultaneously with benzene alkylation, however, due to hydrocracking environment and hydrocracking functionality of catalyst system the side ring cracking or cracking at alpha-carbon will restrict formation of higher aromatics. In the same embodiment it is disclosed that formation aromatics higher than C.sub.10 is very limited. There is no major change in boiling range of reactor outlet product compared to feed stream. Similarly, the change in product density compared to feed stream density is also insignificant.

(27) In yet another embodiment of the present invention, other type of reactions viz. disproportion of toluene molecule trans-alkylation of C.sub.9 with toluene may also occur in the reactor.

(28) In another embodiment of the present invention, RON of the product stream is at least 2 units more preferably 3 units and most preferably 5 units more than the feed steam. It is further disclosed that benzene content of the product stream is preferably less than 1 vol % and most preferably less than 0.5 vol %.

(29) In another embodiment of the present invention, process configurations include fractionation of C.sub.6 fraction from the reactor outlet product and re-circulating the said stream to alkylation reactor for further conversion of un-reacted benzene molecules.

(30) The present process can be also applied for increasing xylene production, through conversion of benzene stream generated in aromatic complex to alkylated benzenes and re-circulating the same in the aromatic complex.

(31) In yet another embodiment of the present invention, the conversion of the benzene is at least more than 70% more preferably more than 60% and most preferably more than 50%. However, with C.sub.6 fraction recycle, the benzene conversion more than 95%, more preferably more than 90% and most preferably more than 85% can be achieved.

(32) Description of FIG. 1:

(33) FIG. 1 represent schematic process diagram of the subject invention. Feed containing benzene, methanol and DMDS have been introduced in the reactor 3 through line 1. Hydrogen is also added in the reactor through line 2. The reactor effluent coating converted and unconverted part of the hydrocarbon, process off gas and water is sent to a high-pressure separator (HPS) 5 through line 4. In the HPS, the water is settled in the bottom of the vessel and sent via line 8 to another separator 10 operated at lower temperature and called Cold High-Pressure Separator (CHPS). However, the effluent hydrocarbon from HPS is sent through line 7 to another separator 13 operated at lower pressure and temperature and called Low Pressure Separator (LPS). The condensed hydrocarbons from low temperature of the CHPS via line 12 are mixed with line 7. The unutilized hydrogen gas is separated in HPS and taken out of the system through line 6 and recycled back in the reactor. The process off gas is separated in the LPS and taken out of the system through line 14 and through line-9 from the CUPS. The product water is separated from the CUPS through line-11 and the product is taken out through line 15.

Example-1

(34) Experiment conducted in a fixed bed micro flow unit (MFU). The feed containing pure benzene and pure methanol and DMDS of 2.5 wt % of the total hydrocarbon feed in brought into contact with catalyst in the reactor. Methanol and benzene ratio in the feedstock maintained at 11:9 (wt:wt). Hydrogen is also added in the reactor. The WABT of the reactor is 390 C., whereas the pressure of the reactor maintained at 50 barg. The LHSV of the liquid stream in the reactor is 2 h.sup.1. The H.sub.2/HC ratio of the experiment is 500 Nm.sup.3/m.sup.3. The hydrocarbon product properties are given in the Table-1.

(35) TABLE-US-00003 TABLE 1 Hydrocarbon product properties Hydrocarbon Type (PIONA Analysis) Unit Values n-paraffins wt % 0 i-paraffins 1 Naphthenes 1 Benzene 89.5 Toluene 5.5 Xylene/EB 1.5 Others 1.5 Total 100 Conversion % 10.5 (C7 + C8) selectivity % 66.7

Example-2

(36) For the experiment the feed composition is same as described in Example-1. The WABT of the reactor is 420 C., whereas the pressure of the reactor maintained at 50 barg. The LHSV of the liquid stream in the reactor is 1 h.sup.1. The H.sub.2/HC ratio of the experiment is 1000 Nm.sup.3/m.sup.3. The hydrocarbon product properties are given in the Table-2.

(37) TABLE-US-00004 TABLE 2 Hydrocarbon product properties Hydrocarbon Type (PIONA Analysis) Unit Values n-paraffins wt % 1.86 i-paraffins 4.75 Naphthenes 3.55 Benzene 37.96 Toluene 20.48 Xylene/EB 17.87 Others 13.53 Total 100 Conversion % 62.04 (C7 + C8) selectivity % 61.8

Example-3

(38) In the experiment, the WABT increased to 440 C. However, the feed composition remains same. The pressure of the reactor maintained at 50 barg. The LHSV of the liquid stream in the reactor is 1 h.sup.1. The H.sub.2/HC ratio of the experiment is 1000 Nm.sup.3/m.sup.3. The hydrocarbon product properties are given in the Table-3.

(39) TABLE-US-00005 TABLE 3 Hydrocarbon product properties Hydrocarbon Type (PIONA Analysis) Unit Values n-paraffins wt % 1.31 i-paraffins 1.98 Naphthenes 1.95 Benzene 52.61 Toluene 20.46 Xylene/EB 13.13 Others 8.56 Total 100 Conversion % 47.39 (C7 + C8) selectivity % 70.9

(40) The distinct advantages of the present process are: i. It can process any refinery stream with benzene content more than 5 vol % ii. process has no restriction with respect to feed sulphur or H.sub.2S. iii. process can be utilized simultaneously for hydrodesulfurization and benzene conversion iv. process has been developed considering the refinery constrains and can be utilized for both benzene management as well as production of petrochemical (aromatic complex) feedstock.