PROCESS AND PLANT FOR CONVERSION OF OXYGENATES

Abstract

A process for producing a C3 olefin product stream and a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range, the process comprising: passing a feedstock stream comprising oxygenates over a catalyst active in the conversion of oxygenates for producing a first olefin stream; conducting the first olefin stream to a first separation step and withdrawing thereof a liquid hydrocarbon fraction comprising at least 50 wt % of the C3-olefins contained in the first olefin stream; conducting the liquid hydrocarbon fraction to a fractionation step and separating therefrom said C3 olefin product stream and the olefin product stream; and converting the olefin product stream into the hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range, particularly sustainable aviation fuel (SAF), by subsequent oligomerization and hydrogenation. The invention provides also a plant for conducting the process.

Claims

1. A process for producing a C3 olefin product stream and a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range, said process comprising: i) passing a feedstock stream comprising oxygenates over a catalyst active in the conversion of oxygenates, at a pressure of 1-100 bar and temperature of 240-400 C.; thereby producing a first olefin stream; ii) conducting the first olefin stream to a first separation step and withdrawing thereof a liquid hydrocarbon fraction comprising at least 50 wt % of the C3-olefins contained in said first olefin stream; iii) conducting the liquid hydrocarbon fraction to a fractionation step and separating therefrom said C3 olefin product stream and an olefin product stream; iv) passing at least a portion of the olefin product stream through an oligomerization step over an oligomerization catalyst for thereby producing an oligomerized stream; v) passing at least a portion of the oligomerized stream through a hydrogenation step over a hydrogenation catalyst, for thereby producing said hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range.

2. Process according to claim 1, wherein: step iv) further comprises subsequently conducting a separation step for thereby producing said oligomerized stream; and/or step v) further comprises subsequently conducting a separation step, for thereby producing said hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range.

3. Process according to claim 1, wherein in step i) the pressure is in the range 1-30 bar, and the temperature is in the range 240-360 C.

4. Process according to claim 1, wherein the catalyst in step i) comprises a zeolite with a framework having a 10-ring pore structure, in which said 10-ring pore structure comprises: (a) a unidimensional (1-D) pore structure, and/or a two-dimensional (2-D) pore structure, and/or (b) a three-dimensional (3-D) pore structure.

5. Process according to claim 1, wherein the catalyst in step i) comprises a zeolite with a framework having a 10-ring pore structure, in which said 10-ring pore structure comprises: (a) a unidimensional (1-D) pore structure, said unidimensional (1-D) pore structure is selected from any of *MRE (ZSM-48), MTT (ZSM-23), TON (ZSM-22), or combinations thereof; the pressure is 1-25 bar, and the temperature is 240-360 C.

6. Process according to claim 1, wherein the catalyst in step i) comprises a zeolite with a framework having a 10-ring pore structure, in which said 10-ring pore structure comprises: (b) a three-dimensional (3-D) pore structure.

7. Process according to claim 1, wherein the zeolite has a silica-to-alumina ratio (SAR) of up to 240.

8. Process according to claim 1, wherein step i) is conducted isothermally.

9. Process according to claim 1, wherein said feedstock stream is combined with a diluent, the feedstock stream is methanol and/or dimethyl ether (DME), and the feedstock is diluted to a methanol and/or DME concentration in the feedstock of 1-30 vol.

10. Process according to claim 9, wherein the diluent is a recycle stream resulting from the process, in which the process further comprises in the separation step (step ii): withdrawing a gaseous fraction comprising C2-C3 olefins (C2=-C3=) as said recycle stream.

11. Process according to claim 1, wherein the separation step (step ii) further comprises withdrawing a water stream and the separation is conducted in a separation unit at 20-80 C., and 5-50 bar.

12. Process according to claim 1, wherein the fractionation step (step iii) is a flash step being conducted in a flashing unit, at 20-80 C., and 5-50 bar.

13. Process according to claim 1, wherein: the oligomerization catalyst in step iv) is: as solid phosphoric acid (SPA), ion-exchange resins or a zeolite catalyst, and the oligomerization step is conducted at a pressure of 30-100 bar, and a temperature of 100-350 C.

14. Process according to claim 1, wherein: the hydrogenation catalyst in step v) is: a Ni-based hydrogenation catalyst; or a Cu-based hydrogenation catalyst; and the hydrogenation step is conducted at a pressure of 1-100 bar and a temperature of 0-350 C.

15. Process according to claim 1, wherein the oligomerization step (step iv) and hydrogenation step (step v) are combined in a single hydro-oligomerization step (OLI/HYDRO), wherein the OLI/HYDRO is conducted in a single reactor having a stacked reactor bed where a first bed comprises the oligomerization catalyst, and a subsequent bed comprises the hydrogenation catalyst.

16. Plant for conducting the process according to claim 1, said plant comprising: an oxygenate conversion reactor comprising a catalyst active in the conversion of oxygenates, wherein the oxygenate conversion reactor is arranged to receive a feedstock stream comprising oxygenates and withdraw said first olefin stream, the oxygenate conversion reactor further arranged to operate at temperature of 240-400 C.; a first separation unit arranged to receive the first olefin stream and withdraw a liquid hydrocarbon fraction comprising at least 50 wt % of the C3-olefins contained in said first olefin stream; a fractionation unit arranged to receive the liquid fraction and withdraw said C3 olefin product stream and olefin product stream; an oligomerization reactor comprising an oligomerization catalyst, wherein the oligomerization reactor is arranged to receive at least a portion of the olefin product stream and withdraw an oligomerized stream; a hydrogenation reactor comprising a hydrogenation catalyst, wherein the hydrogenation reactor is arranged to receive at least a portion of the oligomerized stream and withdraw a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range.

Description

BRIEF DESCRIPTION OF THE DRAWINGS

[0200] FIG. 1 is a simplified figure showing an embodiment of the invention for the conversion of a feedstock comprising oxygenates to olefins and further conversion to jet fuel.

[0201] FIG. 2 shows the product distribution of a first olefin stream exiting MTO in accordance with Example 1.

[0202] FIG. 3 shows a plot of boiling points of a number of compounds at 15 bar.

[0203] FIG. 4 shows a plot of methanol conversion as a function of temperature with a neat feed of methanol (no co-feed i.e. no diluent) compared to a feed comprising propylene as co-feed (diluent), in accordance with Example 2.

[0204] FIG. 5 shows a plot of the effect of propylene on the jet yield and selectivity during oligomerization.

DETAILED DESCRIPTION

[0205] With reference to FIG. 1, a schematic layout of process and plant 100 for producing jet fuel, in particular SAF, and propylene (MTJP process/plant) is shown. A methanol synthesis gas stream 1 containing CO.sub.2 and H.sub.2, or CO, CO.sub.2 and H.sub.2 is introduced to a methanol reactor 10 from which a methanol stream 3 is produced. A part of this is recycled to the methanol reactor 10 as stream 5 while a portion is withdrawn as water stream 7. The resulting methanol stream 9 may be passed to an optional dehydration reactor 12 for converting methanol to dimethyl ether (DME). From the exiting stream 11 of the dehydration reactor 12, a water stream 15 is withdrawn while a portion 13 is recycled. A feedstock 17 comprising oxygenates (MeOH and/or DME) is formed, which is then diluted with a recycle stream 21 comprising ethylene and propylene (C2-C3 olefins), which acts not only as diluent to reduce the exothermicity of the MTO step in MTO reactor 14, but also to enable a lower inlet temperature to the MTO reactor 14 as the onset of the MTO reaction may then take place at lower temperature. After combining with the recycle stream 21, the feedstock 19 is passed to MTO reactor 14, suitably as a plurality of MTO reactors arranged in parallel, thereby producing a first olefin stream 23. The first olefin stream 23 is conducted, suitably after compression, to a first separation step in separation unit 16, suitably a 3-phase separator, and withdrawing therefrom a liquid hydrocarbon fraction 25 comprising a portion, suitably at least 50 wt %, of the C3-olefins contained in said first olefin stream 23; as well as a water stream 27 and said recycle stream 21 as a gaseous fraction containing C2-C3 olefins. The recycle stream 21 may also comprise methane, ethane, propane, carbon monoxide, carbon dioxide and hydrogen.

[0206] The liquid hydrocarbon fraction 25 is conducted to a fractionation step in e.g. a distillation unit, suitably a flashing unit 18, such as a flash distillation unit, thereby easily separating therefrom a C3 olefin product stream 29 having e.g. a propylene purity of at least 93 vol. % (% propene in stream 29), and thus being withdrawn as chemically grade propylene. An olefin product stream 31 is also produced, which after optional evaporation and compression, is conducted to an oligomerization step in oligomerization reactor (OLI reactor) 20; thereby producing an oligomerized stream 33, a part of which may be recycled as stream 35. The oligomerized stream 33 is then add-mixed with hydrogen 37 and passed as stream 30 through a hydrogenation step (HYDRO) in a hydrogenation reactor 22 (HYDRO reactor) for thereby producing a hydrocarbon stream 41 comprising hydrocarbons boiling in the jet fuel range, particularly as SAF. Suitably, the OLI and HYDRO step are combined in a single step (OLI/HYDRO) in a single reactor (not shown) having a stacked reactor bed where a first bed comprises an oligomerization catalyst and a subsequent bed comprises a hydrogenation catalyst. The HYDRO reactor 22 may also be provided as another hydroprocessing reactor, such as a hydrotreating reactor or a hydrocracking reactor.

Example 1

[0207] This example illustrates the product distribution of the MTO reaction and how it can be used advantageously according to the invention. MTO tests were run in a fixed catalyst bed (fixed bed) reactor with a zeolite catalyst ZSM-48 (EU-2) having a 1-D pore structure and a silica to alumina ratio (SAR) of 110, and at the following operating conditions: zeolite catalyst load: 250 mg cat/750 mg SiC (inert diluent), pressure=1 barg (2 bar), space velocity (WHSV)=2 h.sup.1, total flow=3.5 NL/h (59 NmL/min); methanol concentration in the feed (C.sub.MeOH)=10% (volume basis) with nitrogen as the diluent. Thus, P.sub.MeOH is 0.2 bar. The temperature used was in the range 320-360 C.

[0208] Methanol was evaporated and mixed with nitrogen at a ratio of 1:9 (thus C.sub.MeOH=10% volume basis) and fed to the reactor. The reaction was carried out a 320 C. and 360 C., respectively, at full methanol conversion. Products were analyzed by gas chromatography. The product distribution in wt % of the thus obtained first olefin stream at the two temperatures is shown in FIG. 2. The left-hand column is at 320 C. and the right-hand column 360 C. It is to be emphasized that ethylene (C2=) accounts for approximately 1% at 360 C. and is barely detectable at 320 C. Propane is not shown in FIG. 2, but the concentration was found to be very low; at most 0.5%. At both temperatures there is high content of the desired olefins (C3= and C4=-C8=), the latter olefins being particularly suitable for subsequent oligomerization and hydrogenation.

[0209] The boiling point of propylene (propene i.e. C3=) at normal pressure is 47 C., which is impractically low for distillation. At 15 bar however, the boiling point of propene is 34 C. The next higher boiling component of the MTO products, apart from propane which is present in an acceptable concentration for at least chemical grade propene, is isobutene. FIG. 3 shows the boiling points at 15 bar of relevant components. Isobutene has a boiling point of 84 C. at 15 bar, which allows for an easy separation of a propylene-rich gaseous phase and a liquid phase rich in higher olefins and suitable for further conversion to jet fuel, as the first olefin stream is free of aromatics, yet rich in the higher olefins C3-C8=.

Example 2

[0210] This example shows the effect of adding the lower olefin propylene (propene) into the methanol feed to the MTO as recycle stream, corresponding to recycle stream 21 in FIG. 1.

[0211] The comparison is conducted with the same zeolite ZSM-48 (EU-2) at the same reaction conditions in the MTO with 1 mole % propylene (propene), and without (i.e. neat methanol feed). Operating conditions: zeolite catalyst load: 250 mg cat/750 mg SiC, pressure=2 barg (3 bar), space velocity (WHSV)=2 h.sup.1, total flow=3.5 NL/h (59 mL/min); methanol concentration in the feed (C.sub.MeOH)=10% (volume basis) with nitrogen as the diluent.

[0212] FIG. 4 shows that adding propylene (propene) as co-feed (diluent) and thus as a recycle stream, see upper line in the figure, significantly promotes the kick-off or initiation of the oxygenate (methanol) conversion. In the operation of the MTO, there will be significant amounts of light olefins, namely C2-C3 olefins, in particular propylene in the recycle stream, suitably as a portion of the first olefin stream, and which may be utilized anyway for temperature control in the MTO due to its exothermicity. The addition of the lower olefin, e.g. as recycle stream, to the methanol feed enables a significant reduction of the inlet temperature to the MTO, whereby the content of aromatics and paraffins (as used herein, also incl. methane) decreases, while the average olefin chain length increasesand hence the content of higher olefins. Furthermore, the recycle and thus co-feed with the lower olefin significantly increases catalyst longevity, for instance as measured by catalyst cycle time. Moreover, hydrogen transfer reactions are minimized, and not least the lower temperature of the MTO, e.g. 320 C., enables operation at the higher pressure range in the MTO, e.g. 15 or 20 bar or higher, which may be also advantageous by i.a. enabling a higher throughput in the MTO and reduction of equipment size in the plant for producing SAF.

Example 3

[0213] This example shows the effect of propylene (C3=) to oligomerization on the jet yield, more specifically the yield and selectivity to C8-C20.

[0214] The feed to oligomerization was changed from 1-pentene (C5=) to a mixed olefin (C3:C4:C5), i.e. (C3=, C4=, C5=), with the respective proportions 1:1:1 and 1:1:4.

[0215] For each feed, e.g. a feed consisting of C5= in the left-hand part of the plot of FIG. 5 (1-pentene), the yield and selectivity is shown with respect to three different catalysts for oligomerization, thus giving rise to six columns for each feed. From left to right under each feed: column 1 represents the selectivity with a Beta catalyst, column 2 represents the yield with the Beta catalyst; column 3 represents the selectivity with a ZSM-48 catalyst; column 4 represents the yield with the ZSM-48 catalyst; column 5 represents the selectivity with a ZSM-5 catalyst; column 6 represents the yield with the ZSM-5 catalyst.

[0216] It is shown that changing from 1-pentene to a 1:1:1 feed results in much lower yield and slightly lower selectivity. Changing from 1-pentene to a 1:1:4 feed results in similar yield and selectivity. Hence, a feed in which the proportion of C3=increases conveys lower yield and selectivity, and is thus not suitable for oligomerization. A feed with higher olefins is preferred.