IMPROVED METHOD FOR THE CATALYZED HYDROISOMERISATION OF HYDROCARBONS

Abstract

The invention relates to an arrangement of several layers of catalysts arranged in series in a reactor for the hydroisomerisation of hydrocarbons, to a method for the hydroisomerisation of hydrocarbons and to the use of the arrangement for the hydroimerisation of hydrocarbons.

Claims

1. Catalyst arrangement in a reactor for hydroisomerisation of hydrocarbons, wherein at least two catalyst layers are arranged in the reactor, wherein the first catalyst layer is arranged upstream and the second catalyst layer is arranged downstream, and wherein the catalyst of the first catalyst layer is a supported precious metal catalyst for a hydrogenation of the reaction fluid and the catalyst of the second catalyst layer is a bifunctional supported precious metal catalyst, the support of which has acidic or basic properties, for the isomerisation of the reaction fluid after passing through the first catalyst layer.

2. Catalyst arrangement according to claim 1, wherein the support of the catalyst of the first catalyst layer comprises an aluminium oxide, silicon oxide, a metal foam, ceramic or a thermally stable polymer.

3. Catalyst arrangement according to claim 1, wherein the catalyst of the second catalyst layer comprises, as active component, an amorphous aluminosilicate, zeolite, chlorinated aluminium oxide, tungstenated zirconium oxide or sulfonated zirconium oxide.

4. Catalyst arrangement according to claim 3, wherein the catalyst of the second catalyst layer comprises, as active component, tungstenated zirconium oxide or sulfated zirconium oxide, and has been promoted with a transition element or rare earth element.

5. Catalyst arrangement according to claim 1, wherein the downstream catalyst has an immobilized acid or ionic liquid on the support.

6. Catalyst arrangement according to claim 1, wherein the active component of the downstream catalyst has been embedded in a thermally stable organic, ceramic or metallic matrix by using a 3D printing method (rapid prototyping).

7. Catalyst arrangement according to claim 1, wherein the catalyst of the first catalyst layer and/or the catalyst of the second catalyst layer has a precious metal content within a range from 0.05% to 5.0% by weight, preferably from 0.1% to 4.0% by weight and more preferably from 0.1% to 3.0% by weight, based on the weight of the catalyst after ignition loss at 900° C.

8. Catalyst arrangement according to claim 1, wherein the catalyst layers are in the same reactor housing or separately from one another in reactor housings arranged in succession.

9. Use of the catalyst arrangement according to claim 1 for catalytic hydroisomerisation of hydrocarbon mixtures in the presence of aromatics, olefins, organic sulfur compounds, organic nitrogen compounds, carbon monoxide, carbon dioxide, carbonyl sulfide or carbon disulfide or mixtures thereof.

10. Process for catalytic hydroisomerisation of hydrocarbon mixtures in the presence of aromatics, olefins, organic sulfur compounds, organic nitrogen compounds, carbon monoxide, carbon dioxide, carbonyl sulfide or carbon disulfide or mixtures thereof, with a catalyst arrangement according to claim 1, wherein the process comprises the following steps: providing a reactor for the hydroisomerisation; arranging at least two catalyst layers, wherein the first catalyst layer is arranged upstream and the second catalyst layer is arranged downstream, and wherein the catalyst of the first catalyst layer is a supported precious metal catalyst for a hydrogenation of the reaction fluid and the catalyst of the second catalyst layer is a bifunctional supported precious metal catalyst, the support of which has acidic or basic properties, for the isomerisation of the reaction fluid after passing through the first catalyst layer, charging the reactor with a hydrocarbon mixture; converting the hydrocarbon mixture under hydroisomerisation conditions; discharging the generated hydroisomerised hydrocarbon from the reactor.

11. Process according to claim 10 for the variation of the boiling curve and density of a hydrocarbon mixture by cracking reactions or rearrangement reactions.

12. Process according to claim 10 for hydroisomerisation of aromatics to alkylated methylcyclopentanes.

13. Process according to claim 10, wherein the at least two catalyst layers are present in separate columns or separately as column packing materials in a single distillation plant for the reactive distillation.

14. Process according to claim 10, wherein the two catalyst layers are present separately in a microstructure reactor or in separate microstructure reactors.

15. Process according to claim 10, wherein at least one of the two catalyst layers is in the form of a catalytically active membrane in a membrane reactor.

16. Process according to claim 10, wherein the inlet temperature is in the range from 220 to 320° C., preferably in the range from 220 to 260° C., more preferably in the range from 230 to 250° C., most preferably in the range from 235 to 245° C.

17. Process according to claim 10, wherein one or more further catalyst layers are arranged downstream of the catalyst layer arranged downstream.

18. Process according to claim 10, wherein the reaction fluid is a light gasoline fraction.

Description

[0077] The invention is described in detail hereinafter by multiple examples with reference to the appended drawings. The drawings show:

[0078] FIG. 1 a schematic diagram of an arrangement of the catalyst layers in a reactor

[0079] FIG. 2 a schematic diagram of a flow apparatus for performance of a process of the invention for hydroisomerisation of hydrocarbons

EXAMPLES

[0080] The determinations of ignition loss in the context of the present invention were effected to DIN 51081 by determining the weight of about 1-2 g of a sample of the material to be analysed, then heating it to 900° C. under ambient atmosphere and storing it at this temperature for 3 h. Subsequently, the sample was cooled down under protective atmosphere and the remaining weight was measured. The difference in weight before and after thermal treatment corresponds to the ignition loss.

[0081] Experimental Apparatus

[0082] The comparative examples and inventive examples were performed using an experimental apparatus as described in FIG. 2. The setup was chosen for virtually adiabatic characteristics of the reactor. The dimensions of the reactor (20) were such that it could accommodate a total catalyst volume of at least 2500 cm.sup.3. It was also designed such that it could be operated at an operating pressure of 15 to 30 bar gauge.

[0083] For exact control of the volume flow rates, standard electronic mass flow regulators, called flow indication and controls FIC (21), were used. Nitrogen (22) served the purpose merely of purging of the plant in order that no explosive air-hydrogen or air-hydrocarbon mixtures could form. The feed oil (23) was initially charged in a cooled vessel (24) that rested on a balance (25) and was pumped by means of a pump (26) together with the hydrogen (27) into the crossflow microscale heat exchanger I (28). The crossflow microscale heat exchanger (28) was chosen such that it was possible to heat a hydrogen stream up to 400° C. in the above-specified pressure range of 1.5 kg/h (min. 5 kW). The pipelines to the reactor (20) were heated by means of temperature control by a temperature indicator and controller TIC (29) such that the desired reactor inlet temperature was maintained. At the reactor outlet was a thermocouple (30) for determining the reactor outlet temperature. The operating pressure was adjusted using a backpressure control valve (31). The reduced-pressure reaction fluid was guided in a pipe connection heated by means of temperature control by a temperature indicator controller TIC (32) to a sample loop (33) in order to analyse the composition of the reaction fluid with the aid of an online gas chromatograph (34). Alternatively, the sample loop connection permitted constant connection of a pipe connection to the crossflow microscale heat exchanger II (35). The reaction fluid was cooled to at least −10° C. by means of temperature controller (36) in order to collect an integral sample for further characterizations in the liquid sampling vessel (37), which likewise rested on a balance (38) to ascertain a mass balance. The escaping gas was supplied to an offgas conduit (39), determining the mass flow rate with a flow indicator FI (40).

[0084] The calculation of the yield Y, i.e. the product fraction based on molecules having a carbon number≥4, was found as the quotient of the mass m(C4+).sub.liq of the molecules having a carbon number≥4 that were collected in the vessel (37), and the mass m(C4+).sub.gas of the molecules having a carbon number≥4 in the offgas stream, which is determined by means of gas chromatography, divided by the mass m(C4+)inlet of the molecules having a carbon number≥4 that were initially charged in the vessel (24):

[00001]$Y=\frac{m\left(C4+\right)\mathrm{liq}+m\left(C4+\right)\mathrm{gas}}{m\left(C4+\right)\mathrm{inlet}}$

[0085] The proportions by weight reported in tables 1 to 5 are each based on the total weight of the C4+ hydrocarbons present in the corresponding sample.

[0086] For comparative examples 1 and 2 and inventive examples 1 to 3, two light gasoline fractions were used: the olefin-free feed oil A and the olefin-containing feed oil B. The composition and some calculated properties are compiled in table 2.

TABLE-US-00001 TABLE 2 Composition and properties of the feed oils used (RON.sub.THEO: research octane number calculated from the composition) Unit Feed oil A Feed oil B n-Butane % by weight 5.63 5.63 Isobutane % by weight 0.31 0.31 n-Pentane (n-Pn) % by weight 35.63 35.23 Isopentane (i-Pn) % by weight 4.46 4.46 Neopentane % by weight 0.00 0.00 Cyclopentane % by weight 3.63 3.63 n-Hexane % by weight 17.45 17.45 2,2-Dimethylbutane % by weight 0.61 0.61 2,3-Dimethylbutane % by weight 1.76 1.76 2-Methylpentane % by weight 12.89 12.89 3-Methylpentane % by weight 7.81 7.81 Cyclohexane % by weight 1.35 1.35 Methylcyclopentane % by weight 6.26 6.26 n-Heptane % by weight 0.10 0.10 iso-Heptanes % by weight 0.10 0.10 Benzene % by weight 1.99 1.99 Toluene % by weight 0.00 0.00 1-Pentene % by weight 0.00 0.40 Density at 15° C. kg/dm.sup.3 0.657 0.657 Average molecular mass g/mol 78.02 78.01 RON.sub.THEO a.u. 66.90 67.00 iPn/(iPn + nPn) % 11.12 11.24

Comparative Example 1

[0087] The reactor was charged with 1790 g of a commercially available zeolite catalyst, HYSOPAR®-5000 in extrudate form with an average diameter of 1.6 mm and a Pt content of 0.35% by weight from Clariant. The catalyst bed was positioned on an aluminium oxide bed consisting of tablets of dimensions 4.75×4.75 mm.

[0088] After the reactor had been filled, it was sealed pressure tight, and the plant was purged with a nitrogen stream of at least 500 dm.sup.3 (STP)/h versus ambient pressure for one hour. Subsequently, the nitrogen stream and the backpressure regulator were adjusted such that the same gas flow rate was attained at 30 bar gauge. After ten minutes, the gas supply was stopped in order to check the system for leaks. Subsequently, this procedure was repeated with hydrogen. For drying and activation of the catalyst, the reactor inlet temperature was first increased to 150° C. over a period of three hours under a hydrogen gas flow rate of 1000 dm.sup.3 (STP)/h versus ambient pressure. Subsequently, this temperature was maintained for a further three hours. This was followed by a constant increase in the reactor inlet temperature to 300° C. over a period of eight hours. This temperature was subsequently maintained for a further three hours.

[0089] Before the start of the catalytic experiment, the reactor inlet temperature was reduced to 200° C. at a constant cooling rate of 1 K/min and the hydrogen flow rate was adjusted to 905 dm.sup.3 (STP)/h versus 20 bar gauge.

[0090] At the start of the catalytic experiment, the olefin-free feed oil A was supplied at a mass flow rate of 2.628 kg/h and the temperature at the reactor inlet was increased from 200° C. to a first target temperature. After attainment of this temperature, these conditions were not changed over a period of three hours, and then the temperature at the reactor inlet was increased by a desired temperature. The number of possible gas chromatography analyses was determined by the necessary separation time. Typically, three injections were possible within three hours.

Comparative Example 2

[0091] The reactor charge, procedure and experimental conditions corresponded to those of comparative example 1, except that the olefin-containing feed oil B was used.

Example 1

[0092] The reactor was charged with 1432 g of a commercially available zeolite catalyst, HYSOPAR®-5000 in extrudate form with an average diameter of 1.6 mm and a Pt content of 0.35 Pt from Clamant. In addition, a further bed consisting of 250 kg of HYSOPAR®-1000 type catalyst in the form of a porous, weakly acidic aluminium oxide and with a Pt content of 0.30% by weight was introduced onto this catalyst bed. The bed of the HYSOPAR®-5000 catalyst was positioned on an aluminium oxide bed of tablets of dimensions 4.75×4.75 mm.

[0093] The procedure and experimental conditions corresponded to those of experimental example 1; the olefin-free feed oil A was likewise used.

Example 2

[0094] The reactor was charged with 1432 g of a commercially available zeolite catalyst, HYSOPAR®-5000 in extrudate form with an average diameter of 1.6 mm and a Pt content of 0.35% by weight from Clamant. In addition, a further bed consisting of 250 kg of HYSOPAR®-1000 catalyst in the form of a porous, weakly acidic aluminium oxide and with a Pt content of 0.30% by weight from Clariant was introduced onto this catalyst bed. The bed of the HYSOPAR®-5000 catalyst was positioned on an aluminium oxide bed of tablets of dimensions 4.75×4.75 mm.

[0095] The procedure and experimental conditions corresponded to those of comparative example 1, except that the olefin-containing feed oil B was used.

Example 3

[0096] The reactor was charged with 1432 g of a commercially available zeolite catalyst, HYSOPAR®-5000 in extrudate form with an average diameter of 1.6 mm and a Pt content of 0.25% by weight from Clariant. In addition, a further bed of 250 kg of HYSOPAR®-1000 catalyst in the form of a porous, weakly acidic aluminium oxide and with a Pt content of 0.30% by weight from Clariant was introduced onto this catalyst bed. The bed of the HYSOPAR®-5000 catalyst was positioned on an aluminium oxide bed of tablets of dimensions 4.75×4.75 mm.

[0097] The procedure and conditions corresponded to those of comparative example 1, except that the olefin-containing feed oil B was used.

[0098] Table 3 collates the results from the analysis of liquid products that were generated at different reactor inlet temperatures. The results show that, in the case of the inventive examples, higher yields were already achieved at lower inlet temperatures than in the comparative examples. Moreover, this result required a smaller amount of costly platinum overall.

TABLE-US-00002 TABLE 1 Summary of the reactor temperatures and the essential properties of the resultant product streams from comparative examples 1 and 2 and inventive examples 1 to 3: Comparative example Example 1 2 1 2 3 Feed oil A B A B B Inlet temperature ° C. 255 265 245 242 242 Outlet temperature ° C. 270 285 260 260 260 Proportion of the following hydrocarbons [% by weight] n-Butane 2.95 3.00 2.91 2.90 2.90 Isobutane 3.30 3.26 3.33 3.33 3.33 n-Pentane 14.40 15.44 14.37 14.36 14.36 Isopentane 26.39 25.33 26.44 26.45 26.45 Neopentane 0.00 0.00 0.00 0.00 0.00 Cyclopentane 2.67 2.66 2.67 2.67 2.67 n-Hexane 10.11 10.34 9.93 9.91 9.90 2,2-Dimethylbutane 6.76 6.64 7.02 7.04 7.07 2,3-Dimethylbutane 3.22 3.58 3.24 3.25 3.25 2-Methylpentane 13.24 13.01 13.24 13.24 13.24 3-Methylpentane 8.48 8.22 8.37 8.36 8.35 Cyclohexane 2.82 2.62 2.82 2.82 2.82 Methylcyclopentane 5.56 5.57 5.55 5.55 5.55 n-Heptane 0.06 0.06 0.06 0.06 0.06 iso-Heptanes 0.06 0.06 0.06 0.06 0.06 Benzene 0.00 0.20 0.00 0.00 0.00 Toluene 0.00 0.00 0.00 0.00 0.00 Density [kg/dm.sup.3] 0.6565 0.6566 0.6564 0.6564 0.6564 RON [a.u.] 78 78 78 79 79 Yield [% by weight] 95 92 96 97 97

Example 4

[0099] The reactor was charged with 860 g of a commercially available zeolite catalyst, HYSOPAR®-7000 in extrudate form with an average diameter of 1.6 mm and a Pt content of 0.25% by weight from Clariant. In addition, a further bed of 900 g of HYSOPAR®-1000 catalyst in the form of a porous, weakly acidic aluminium oxide and with a Pt content of 0.30% by weight from Clariant was introduced onto this catalyst bed. The bed of the HYSOPAR®-5000 catalyst was positioned on an aluminium oxide bed of tablets of dimensions 4.75×4.75 mm.

[0100] The procedure corresponded to that of comparative example 1, except that the hydrogen flow rate was adjusted to 839 dm.sup.3 (STP)/h versus 30 bar gauge, and a benzene-containing feed oil C with the following composition and properties was used:

[0101] Feed oil C: 94% by weight of n-hexane and 6% by weight of benzene [0102] RON.sub.THEO=32 [0103] Density at 15° C.=0.6811 kg/dm.sup.3 [0104] Average molecular mass 98.875 g/mol

[0105] Table 4 collates the results from the analysis of liquid products that were generated in two experimental procedures A and B at different reactor inlet temperatures.

TABLE-US-00003 TABLE 4 Summary of the reactor temperatures and the essential properties of the product streams obtained from example 4 Parameter Unit A B Inlet temperature ° C. 220 240 Outlet temperature ° C. 270 290 Benzene % by weight 0 0 Cyclohexane % by weight 2 2 Methylcyclopentane % by weight 3 3 n-Hexane % by weight 34 32 i-Hexane % by weight 61 61 C4-05 alkanes % by weight 0 2 Yield % by weight 97 93 RON a.u. 63 62 Density kg/dm.sup.3 0.6672 0.6674

[0106] It can be seen from the data from table 4 that the arrangement of the invention enables lowering of the inlet temperature with simultaneously improved yield and elevated RON.

Example 5

[0107] The catalyst and procedure corresponded to those of example 4, except that a feed oil D having the following composition and properties was used: [0108] Feed oil D: Kerosene fraction having a density of 0.7691 kg/dm.sup.3 at 15° C., 30 ppm by weight of sulfur and simulated boiling characteristics to ASTM D-2887 as in table 5.

TABLE-US-00004 TABLE 5 Boiling curve to ASTM D-2887 for the feed oil D used Boiling progression in % by weight Temperature [° C.] Start 98.00 5 140.30 10 158.70 20 175.40 30 185.40 50 204.30 70 227.60 80 237.70 90 255.30 95 266.10 End 287.50

[0109] According to M. R. Riazi, Characterization and Properties of Petroleum Fractions, ASTM (2005) 1st edition, page 131, the “freeze point” is calculated from boiling progression and density to be FRP=−35° C.

[0110] The experimental conditions corresponded to those of example 4.

[0111] Table 6 collates the results from the analysis of liquid product streams that were generated in experimental procedures A, B and C at different reactor inlet temperatures.

TABLE-US-00005 TABLE 6 Summary of the reactor temperatures and the essential properties of the product streams obtained from example 5 A B C Inlet temperature [° C.] 200 250 280 Outlet temperature [° C.] 240 290 320 Boiling progression in % by weight Temperature [° C.] (to ASTM D-2887) Start 95 48 22 5 133 125 48 10 152 144 89 90 255 254 246 End 287 288 289 FRP [° C.] -37 -39 -51 Yield [% by weight] 99.3 98.6 98.9 Density [kg/dm.sup.3] 0.7700 0.7694 0.7594

[0112] It can be seen from table 6 that the arrangement of the invention can achieve lowering of the FRP. It is also found that the yield of C4+ hydrocarbons can be increased when the process is performed at a lower inlet temperature.