PRODUCTION OF LIQUID HYDROCARBONS

Abstract

The invention relates to a process for the conversion of hydrogen and one or more oxides of carbon to hydrocarbons, which process comprises: contacting hydrogen and one or more oxides of carbon with a catalyst in a reaction zone; removing from the reaction zone an outlet stream comprising unreacted hydrogen, unreacted one or more oxides of carbon and one or more hydrocarbons and feeding the outlet stream to a separation zone in which the outlet stream is divided into at least three fractions, in which; a first fraction predominantly comprises unreacted hydrogen, unreacted one or more oxides of carbon and hydrocarbons having from 1 to 4 carbon atoms; a second fraction predominantly comprises hydrocarbons having 5 to 9 carbon atoms, at least a portion of which hydrocarbons having from 5 to 9 carbon atoms are olefinic; and a third fraction predominantly comprises hydrocarbons having 10 or more carbon atoms; characterised in that at least a portion of the second fraction is recycled to the reaction zone.

Claims

1. A process for the conversion of hydrogen and one or more oxides of carbon to hydrocarbons, which process comprises: contacting hydrogen and one or more oxides of carbon with a catalyst in a reaction zone; removing from the reaction zone an outlet stream comprising unreacted hydrogen, unreacted one or more oxides of carbon and one or more hydrocarbons and feeding the outlet stream to a separation zone in which the outlet stream is divided into at least three fractions, in which; a first fraction predominantly comprises unreacted hydrogen, unreacted one or more oxides of carbon and hydrocarbons having from 1 to 4 carbon atoms; a second fraction predominantly comprises hydrocarbons having 5 to 9 carbon atoms, at least a portion of which hydrocarbons having from 5 to 9 carbon atoms are olefinic; and a third fraction predominantly comprises hydrocarbons having 10 or more carbon atoms; characterised in that at least a portion of the second fraction is recycled to the reaction zone.

2. A process as claimed in claim 1, in which the reaction zone is maintained at a temperature in the range of from 150 to 400° C. and a pressure in the range of from 10 to 100 bara (1.0 to 10.0 MPa), preferably a temperature in the range of from 150 to 400° C. and a pressure in the range of from 10 to 85 bara (1.0 to 8.5 MPa), more preferably at a temperature in the range of from 170 to 400° C. and a pressure in the range of from 35 to 85 bara (3.5 to 8.5 MPa), and more preferably at a temperature in the range of from 250 to 400° C. and a pressure in the range of from 45 to 85 bara (4.5 to 8.5 MPa).

3. A process as claimed in claim 1, in which the reaction zone comprises a solid, fixed bed Fischer-Tropsch catalyst.

4. A process as claimed in claim 1, in which the catalyst comprises iron.

5. A process as claimed in claim 4, in which the catalyst comprises one or more promoters selected from a manganese promoter, a potassium promoter, a lanthanide promoter such as a cerium promoter, and a copper promoter.

6. A process as claimed in claim 5, in which the catalyst comprises a manganese promoter, a potassium promoter, a cerium promoter and a copper promoter.

7. A process as claimed in claim 1, in which the separation zone comprises a flash separation zone and a fractionation zone, in which the outlet stream from the reaction zone is fed to the flash separation zone to produce a gaseous fraction which is the first fraction, and a liquid fraction predominantly comprising hydrocarbons having 5 or more carbon atoms, which liquid fraction is fed to the fractionation zone to produce the second fraction predominantly comprising hydrocarbons having 5 to 9 carbon atoms at least a portion of which are olefinic, and a third fraction comprising hydrocarbons having 10 or more carbon atoms.

8. A process as claimed in claim 1, in which at least a portion, but not all, of the first fraction is recycled to the reaction zone.

9. A process as claimed in claim 1, in which a portion of the first fraction is separated into C.sub.3-C.sub.4 fraction which comprises an increased concentration of C.sub.3-C.sub.4 hydrocarbons compared to the first fraction, and a lights fraction, which comprises an increased concentration of hydrogen, one or more oxides of carbon and C.sub.1-C.sub.2 hydrocarbons compared to the first fraction.

10. A process as claimed in claim 9, in which at least a portion of the C.sub.3-C.sub.4 fraction is fed to a dehydrogenation zone which is maintained under conditions such that C.sub.3-C.sub.4 alkanes can be converted to corresponding olefins, to produce a C.sub.3.sup.═-C.sub.4.sup.═ fraction that has an increased concentration of C.sub.3-C.sub.4 olefins compared to the C.sub.3-C.sub.4 fraction, at least a portion of which C.sub.3.sup.═-C.sub.4.sup.═ fraction is fed to the reaction zone.

11. A process as claimed in claim 9, in which at least a portion of the lights fraction is fed to a reforming zone, in which at least a portion of the C.sub.1-C.sub.2 hydrocarbons and CO.sub.2 are converted to CO and H.sub.2 to produce a reformed fraction, at least a portion of which reformed fraction is fed to the reaction zone.

12. A process as claimed in claim 1, in which at least a portion of the unrecycled second fraction is used to make gasoline, or is used to produce hydrocarbons that are blended with gasoline.

13. A process as claimed in claim 12, in which the portion of the unrecycled second fraction is isomerised and/or alkylated before being used as or blended with gasoline.

14. A process as claimed in claim 1, in which at least a portion of the third fraction is used to make jet fuel and/or diesel fuel, or is used to produce hydrocarbons that can be blended with jet fuel and/or diesel fuel.

15. A process as claimed in claim 14, in which the portion of the third fraction is hydrogenated before being used as or blended with jet fuel and/or diesel fuel.

16. A process as claimed in claim 15, in which the portion of the third fraction is isomerised either prior to or during hydrogenation.

17. A diesel fuel or aviation fuel as made by the process of claim 14.

18. A gasoline fuel as made by the process of claim 12.

Description

EXPERIMENTAL

[0063] There now follow non-limiting examples illustrating the invention, with reference to the drawings in which:

[0064] FIG. 1 is a schematic overview of an embodiment according to the present invention,

[0065] FIG. 2 is a schematic overview of an embodiment similar to FIG. 1, that includes the removal of water and other oxygen-containing compounds in the separation zone;

[0066] FIG. 3 is a schematic representation of the apparatus used to perform the experiments;

[0067] FIG. 4 is a graph showing the hydrocarbon distribution (based on numbers of carbon atoms) in the products of a reaction involving no recycle of hydrocarbons; and

[0068] FIG. 5 is a graph showing the hydrocarbon distribution (based on numbers of carbon atoms) in the products of a reaction involving recycle of hydrocarbons.

[0069] FIG. 1 shows a process comprising a first section, 1, which relates to the Fischer-Tropsch reactor, associated inlets for feedstocks and recycle lines, and the outlet for the outlet stream, and a second section, 2, which relates to apparatus and process lines associated with separating the outlet stream into various fractions. The first section comprises a syngas inlet, 3, recycle lines from the first, 4, and second, 5, fractions of the separation zone, a reactor, 6 (the reaction zone), containing a fixed, solid particulate catalyst bed, 7, and an outlet for the outlet stream, 8, that leads to the second section. In the embodiment shown, the syngas feedstock is pre-mixed with the recycled components of the first fraction and second fraction from the separation zone before being fed to the reactor.

[0070] The second section comprises the separation zone. The separation zone comprises a flash separator as the flash separation zone, 9, in which a gaseous fraction, 10 (the first fraction), comprising predominantly unreacted hydrogen and one or more oxides of carbon together with C.sub.1-4 hydrocarbons is removed. A portion of this is recycled back to the reactor via recycle line 4, and a portion is removed from the process via purge line 11.

[0071] The liquid fraction, 12, from the flash separation zone, predominantly comprising C.sub.5+ hydrocarbons, is fed to a fractionation column, 13. From the top of the fractionation zone, a light fraction, 14, predominantly comprising further unreacted hydrogen and one or more oxides of carbon together with C.sub.1-4 hydrocarbons, is removed and combined with the gaseous phase, 10, removed from the flash separation zone. A medium-boiling fraction predominantly comprising C.sub.5-9 hydrocarbons, 15 (the second fraction), at least some of the hydrocarbons being olefinic, is removed from a lower portion of the fractionation column, a portion of which is recycled to the reaction zone via recycle line 5. An unrecycled portion of the second fraction, 16, is optionally further isomerised to produce branched hydrocarbons in the gasoline boiling range.

[0072] A higher boiling fraction, 17, (third fraction) comprising predominantly C.sub.10+ hydrocarbons, and preferably predominantly C.sub.10-20 hydrocarbons, is removed from a lower portion of the distillation column. This is also optionally hydrogenated to produce alkanes in the diesel oil boiling range, optionally after additional isomerisation.

[0073] A heavy fraction, 18, comprising long chain and high boiling point components is removed from the base of the column, and is optionally converted to diesel oil boiling range alkanes using a process such as hydrocracking, or can optionally be used to make high value synthetic base oils for use as or in the production of lubricants.

[0074] FIG. 2 shows a similar process to FIG. 1, in which positions for removal of water and other oxygen-containing compounds formed in the reaction are shown. Thus, at the base of each of the two separation zones, 9 and 13, an aqueous phase comprising water and any other dissolved compounds, typically oxygenated organic compounds such as alcohols, and which is separate from the hydrocarbon-containing phases, is removed (respectively 19 and 20). The hydrocarbon-containing phases, 12 and 18, are removed from the columns at a position above the interface with the aqueous phase.

Experiment 1

[0075] A zeolite-Y supported iron catalyst was prepared according to a procedure described in PCT application PCT/EP2012/070897 (for catalyst A, pages 30-31). The catalyst contained Fe, Ce, and Cu on a zeolite-Y support, and was prepared as follows:

[0076] Y-zeolite was prepared in the Na.sup.+ cation exchanged form (NaY), and ion-exchanged with K. The ion exchange of NaY was carried out by adding 12 g of NaY to a 600 ml of a 0.5M K.sub.2CO.sub.3 solution in doubly deionized water. The amount of K.sub.2CO.sub.3 in the solution represented a 6-fold excess of K.sup.+ with respect to the amount of cation-exchanging sites of the zeolite. The resulting suspension was stirred and heated at 80° C. with reflux cooling for a minimum of 4 hours. Subsequently the resulting ion-exchanged zeolite was filtered and washed with doubly deionized water.

[0077] This ion-exchange procedure was repeated three times, and the resulting material was dried before use. The resulting KY zeolite was impregnated with a suitable amount of solution of Fe(NO.sub.3).sub.2, Ce(NO.sub.3).sub.3 and Cu(NO.sub.3).sub.2. The volume of solution used was equal to the pore volume of the zeolite added. These nitrate salts are highly soluble and allow the impregnation of metals to be carried out simultaneously. The resulting slurry was dried at 120° C. and calcined in air at 550° C. for 18 h.

[0078] The overall composition of the impregnated transition metal ions in the catalyst then reflects the following atomic ratios; Fe:Ce:Cu=86:9.5:4.5. Zeolite-Y with a Si/Al ratio of 2,9 theoretically contains 14,4 wt. % K when fully exchanged.

[0079] The apparatus shown schematically in FIG. 3 was used to perform single pass and recycle experiments.

[0080] 10 g of the Fe/Cu/Ce on KY catalyst, with a particle size of 1-2 mm, was loaded into reactor, 100, having internal diameter 22 mm, to form a catalyst bed, 101, with a length of 100 mm. Three thermocouples were located at the top, middle and bottom of the catalyst bed within a thermowell of 6 mm diameter. Only the central thermocouple, 102, is shown in FIG. 3.

[0081] The apparatus comprised three gas feed lines, for nitrogen (as a purge), 103a, for syngas, 103b, and for carbon dioxide, 103c. The flows were controlled respectively by isolation valves 104a, 104b and 104c, pressure regulators 105a, 105b and 105c, and mass flow control valves, 106a, 106b and 106c. The pressure regulators and mass flow control valves formed part of the control system, 107, represented generally in FIG. 3 by dashed lines. The syngas feed was a mixture of hydrogen and carbon monoxide, with a H.sub.2:CO molar ratio of 2:1.

[0082] Compressor, 108, pressurised the gases to the desired reaction pressure. The gases were heated at heater, 109, before passing to the reactor, 100. Temperature controller, 110, interfaced with heat exchanger, 109, was used to maintain a desired temperature in the catalyst bed based on the temperature at thermocouple 102. The reactor comprised a cooling jacket, 111, to avoid large temperature excursions. In the examples described below, reaction pressure was maintained in the range of from 30 to 35 barg.

[0083] The hydrocarbon-containing outlet stream from the reactor was cooled via heat exchanger, 112, to near ambient temperature, and then fed to three-phase separator, 113, operating at a pressure of 10 to 15 barg, controlled by pressure regulator, 114. The heat exchanger was regulated based on a temperature measurement, 112a, in the separator 113.

[0084] The vapour phase from the separator, comprising unreacted syngas components, and light hydrocarbons, typically in the C.sub.1 to C.sub.4 range, was then removed from the system through vent, 115, or recycled back to reactor, 100 via compressor, 108. The proportion of vented or recycled components was controlled by pressure controller, 116.

[0085] A liquid phase comprising water and oxygen -containing compounds was removed from the base of separator 113, and passed via a separation vessel, 117, where vapours were removed via vent, 118, and the remaining water and oxygenate-containing liquid phase being removed from the system via 119. Level control at 120a was used to regulate removal of this base stream through valve, 120.

[0086] A separate liquid phase stream comprising predominantly C.sub.5+ hydrocarbons was also removed from the first separator, 113, at a position above the interface with the aqueous phase, and fed to a second separator, 121, regulated by valve 122 based on level control at 122a.

[0087] A vapour fraction comprising C.sub.5 hydrocarbons was removed from the top of the column. Compressor, 123, was used to control the pressure, measured at pressure sensor 123a, in the second separator, 121, to less than 6 barg. The temperature of this second column was higher than that of the first column, to increase the proportion of C.sub.5 hydrocarbons in the vapour fraction.

[0088] This vapour fraction was either recycled to reactor, 100, via compressor, 108. Alternatively, for single pass operation, the fraction was passed to vent, 116, by opening manual control valve, 124, and closing manual control valve 125.

[0089] From the base of the second separator, 121, a liquid water and oxygenate-containing phase was removed through a water boot, 126, and passed to vessel 118, as described above for the corresponding liquid phase from the base of the first separator, 113. Flow of the base stream to vessel 117 through valve, 127, was based on level control at 127a.

[0090] Hydrocarbon liquid phase, comprising the desired product hydrocarbons, was removed from second separator, 121, and split into two streams. One stream formed a recycle loop, which was used to maintain temperature, measured at 128a, in second separator, 121. This recycle stream was pumped via pump, 129, through heater, 128, and back to the separator together with hydrocarbon phase from the first separator, 113. The other stream was passed to vessel, 130, where vapours were removed through vent, 131, and product removed through 132. Control of this stream from second separator 121 to vessel 130 was achieved by control of valve 133 based on level measured at 133a.

[0091] Product removed at 132 was vaporised and analysed by gas chromatography, using a flame induction detector, and using a device fitted with a 25 m, 0.15 mm inner diameter CP-Sil 5 non-polar column.

[0092] In the Examples below, the catalyst was pre-reduced in a flow of pure hydrogen at a gas hourly space velocity of 2000, a pressure of 20 barg, and a temperature of 500° C. for 2.5 hours, and allowed to cool to a temperature of 340-350° C. for 30 minutes before being contacted with syngas and brought up to the reaction pressure of 30-35 barg. The syngas flow was started at 180 minutes.

COMPARATIVE EXAMPLE 1

[0093] This example used a single-pass configuration, such that there was no recycle of vapour fractions from the first or second separators to the reactor.

[0094] From a time period of 180 minutes to 208 minutes on stream, the flow of fresh syngas feed (H.sub.2:CO mole ratio of 2:1) was maintained at 200 ml/g catalyst/min (volume based on STP), i.e. a total volume of 2000 ml/min.

[0095] The hydrocarbon distribution in the product from 132 collected over the course of this period on stream, based on the numbers of carbon atoms in the hydrocarbon molecules, is shown in FIG. 4. This shows that shorter chain hydrocarbons, predominantly C.sub.2-C.sub.6 hydrocarbons are the major components of the product.

EXAMPLE 1

[0096] At 208 minutes on-stream, the apparatus was switched to recycle mode, such that a recycle stream comprising vapour fraction from the first and second separation zone was co-fed to the reactor in addition to fresh syngas. Table 1 shows the different volume ratios of the recycled gases to fresh syngas feed at various stages of reaction (measurements taken at the specified time on stream), together with the temperature readings at the top, middle and bottom of the catalyst bed.

TABLE-US-00001 TABLE 1 Effects of Recycle Stream on Catalyst Bed Temperature Profile Time on Recycle Temperature (° C.) Temperature stream (min) Ratio Bottom Middle Top Gradient (° C.) 208 0 357.0 321.7 291.8 65.2 260 2:1 359.7 337.0 312.7 40.0 304 4:1 355.8 340.9 322.2 33.6 372 8:1 337.1 335.0 330.6 6.5

[0097] For the period 180-208 minutes on stream, fresh syngas only was used (there was no recycle), and a temperature gradient of 65.2° C. was observed across the catalyst bed. The gradient arises as a result of the exothermic reaction associated with the conversion of syngas to hydrocarbons.

[0098] Between 208 and 260 minutes on stream, a ratio of recycled gas to fresh syngas of 2:1 was employed. At 260 minutes, just before changing the recycle ratio, a temperature gradient across the catalyst bed of 40° C. was observed, lower than the gradient without any recycle. At 260 minutes, the recycle ratio was changed to 4:1, and at 304 minutes, just before a further change in recycle ratio, the temperature gradient was 34.6° C. Between 304 and 372 minutes on-stream, a recycle ratio of 8:1 was employed, and the temperature gradient at 372 minutes was 6.5° C.

[0099] Thus, increases in the proportion of recycled gas compared to fresh syngas feed resulted in lower temperature gradients across the catalyst bed, demonstrating the efficacy of medium sized hydrocarbons in the recycled stream in achieving temperature control in the catalyst, and enabling control of reaction temperature by control of recycle ratio. With reference to FIG. 3, then temperature control in the reactor can be achieved through a variety of mechanisms, for example variation of recycle flow (by control of valves 116, 124 and 125), syngas feed flow (via pressure and mass flow controllers 105b and 106b), heater control (at heater 109), reactor cooling (at 111) and reactor pressure (via compressor 108).

[0100] FIG. 5 shows the hydrocarbon distribution resulting from the combined liquid hydrocarbon product collected at point 132 in FIG. 3 over the whole period of reaction where recycle was employed, i.e. in the period from 208 to 372 minutes on stream. A clear shift to longer hydrocarbon chain lengths is observed demonstrating that not only is improved temperature/exotherm control possible, but also improved product selectivity to higher (C10+) hydrocarbons can be achieved.