SINGLE STAGE PROCESS COMBINING NON-NOBLE AND NOBLE METAL CATALYST LOADING

Abstract

The present disclosure relates to a process for reducing the amount of aromatics in a raw feed stream comprising hydrocarbons, more than 200 ppmw sulfur or 1000 ppmw sulfur as either hydrocarbon heteroatoms or as other sulfide compounds as well as at least 10% by weight di-aromatics or poly-aromatics and at least 30% by weight aromatics in total said process comprising the steps of hydrotreating said raw feed stream in the presence of hydrogen and a material catalytically active in hydrotreatment with a severity resulting in a conversion of sulfur hydro-carbon heteroatoms to hydrogen sulfide of at least 50% providing a pre-treated stream, separating said pre-treated stream at least into a second stage feed stream and a stream rich in hydrogen sulfide, directing said second stage feed stream to contact a material catalytically active in hydrocracking and ring opening, and to contact a material catalytically active in saturation of aromatics, wherein the material catalytically active in hydrocracking and ring opening is positioned upstream, downstream or mixed with said material catalytically active in saturation of aromatics, and withdrawing a dearomatized stream, wherein said the amount of aromatics of said dearomatized stream is less than 50%, 70%, 90% or 95% of the amount of aromatics in said raw feed stream, with the associated benefit of said process of providing efficient dearomatization with low yield loss.

Claims

1. A process for reduced the amount of aromatics in a raw feed stream comprising hydrocarbons, more than 200 ppmw sulfur or 1000 ppmw sulfur as either hydrocarbon heteroatoms or as other sulfide compounds as well as at least 10% by weight di-aromatics or poly-aromatics and at least 30% by weight aromatics in total said process comprising the steps of i. hydrotreating said raw feed stream in the presence of hydrogen and a material catalytically active in hydrotreatment with a severity resulting in a conversion of sulfur hydrocarbon heteroatoms to hydrogen sulfide of at least 50% providing a pre-treated stream, ii. separating said pre-treated stream at least into a second stage feed stream and a stream rich in hydrogen sulfide, iii. directing said second stage feed stream to contact a material catalytically active in hydrocracking and ring opening, and to contact a material catalytically active in saturation of aromatics, wherein the material catalytically active in hydrocracking and ring opening is positioned upstream, downstream or mixed with said material catalytically active in saturation of aromatics, and withdrawing a dearomatized stream, wherein said the amount of aromatics of said dearomatized stream is less than 50%, 70%, 90% or 95% of the amount of aromatics in said raw feed stream.

2. The process according to claim 1 wherein said material catalytically active in hydrocracking and ring opening first material comprises a base metal and is provided in presulfided form and said and said material catalytically active in saturation of aromatics comprises a noble metal and is provided in prereduced form.

3. The process according to claim 1, wherein the material catalytically active in hydrocracking and ring opening comprises one or more base metals preferably taken from the group comprising group 6 elements and Ni or Mo.

4. The process according to claim 1, wherein the material catalytically active in hydrocracking and ring opening comprises an acidic support, such as a zeolite or silica-alumina.

5. The process according to claim 1, wherein the material catalytically active in saturation of aromatics comprises one or more one or more noble metals preferably taken from the group comprising Ru, Rh, Pd, Os, Ir and Pt, and more preferably taken from the group comprising Pt and Pd.

6. The process according to claim 1, wherein the material catalytically active in saturation of aromatics comprises an acidic support, such as a zeolite or silica-alumina

7. The process according to claim 1, wherein the maximum temperature of the first catalytically active material is 250 C.-350 C.

8. The process according to claim 1, wherein the maximum temperature of the second catalytically active material is 250 C.-350 C.

9. The process according to claim 1, wherein the difference between the outlet temperature of the first catalytically active material and inlet temperature of the second catalytically active material is less than 40 C.

10. A process plant for conversion of a stream of heavy aromatic hydrocarbon mixture in to a hydrocarbon mixture rich in middle distillate comprising a first stage reactor unit containing a hydrotreatment catalyst, said first stage reactor unit having an inlet and an outlet, a means for gas/liquid separation having an inlet and a gas outlet and a liquid outlet, and a second stage reactor unit comprising one or several reactors and containing at least a presulfided material catalytically active in hydrocracking and a prereduced material catalytically active in hydrodearomatization, said second stage reactor unit having one or more inlets and a single outlet, in which the stream of heavy aromatic hydrocarbon mixture is in fluid communication with an inlet of the first stage reactor, the outlet of the first stage reactor unit is in fluid communication with the inlet of the means for gas/liquid separation, the outlet for liquid of the first means for gas/liquid separation is in fluid communication with the inlet of the second reactor unit, a stream of hydrogen is optionally in fluid communication with a further inlet of the second reactor, the outlet of the second reactor unit provides the hydrocarbon mixture rich in middle distillate.

Description

[0063] FIG. 1 shows a process for conversion of a LCO/gas oil feed mixture to diesel according to the present invention.

[0064] FIG. 2 shows a process for conversion of a LCO/gas oil feed mixture to diesel according to the prior art.

[0065] In FIG. 1 according to an embodiment of the present invention, a LCO/gas oil feed mixture in combination with hydrogen 2 is directed to contact a hydrotreatment catalyst 4 in a in a pretreater unit 6, in order to provide a pretreated feed 8. A gas phase 12, including hydrogen sulfide, is removed in a means of gas/liquid separation 10 (such as an interstage stripper or a flash drum), and the pretreated hydrocarbon feed 14 is combined with a stream of unconverted oil 40 and directed as second stage feed 16 to contact a base metal hydrocracking catalyst 18 in a hydrocracking catalyst bed in a second stage reactor 22. The entire effluent of the hydrocracking catalyst bed is transferred to contact a noble metal hydrodearomatization catalyst 20 in a separate catalyst bed or possibly in a separate reactor. The base metal catalyst 18 and the noble metal catalyst 20 can be operated in the same reactor, because both have been activated ex-situ and because a moderate level of sulfur is present in the reaction mixture. The second stage product stream 24 is first separated in a gas phase 28 and a liquid phase product 30 in a gas/liquid separator, and then fractionated in a fractionator 32 into naphtha 34, jet fuel 36 and an unconverted oil (UCO) fraction 38 dominated by product boiling in the diesel range. A part of the UCO 40 is directed to contact the hydrocracking catalyst and hydrodearomatisation catalyst again. This process has the benefit that the pressure and temperature of the HDC step 18 and the HDA step 20 may be optimized independently of the HDS step 4, and thus an increased specificity and yield may be obtained in the product 42 which is an unconverted oil stream comprising diesel.

[0066] As the hydrocracking catalyst active in hydrodearomatization as well as hydrocracking, if it is operated at moderate temperature it will catalyze dearomatization by ring opening while it will catalyze hydrocracking at elevated temperature. Therefore, the inlet temperature to reactor 22 will make it possible to control the extent of hydrocracking, and accordingly it will be uncomplicated to switch between active hydrocracking (at elevated temperatures) and ring opening and moderate hydrocracking at intermediate temperatures.

[0067] In FIG. 2 according to the prior art, a LCO/gas oil feed mixture in combination with hydrogen 2 is directed to contact a hydrotreatment catalyst 4 in a in a pretreater unit 6, in order to provide a pretreated feed 8, A gas phase 12, including hydrogen sulfide, is removed in a means of gas/liquid separation 10 (such as an interstage stripper or a flash drum), and the pretreated hydrocarbon feed 14 is combined with a stream of unconverted oil 40 and directed as second stage feed 16 to contact a base metal hydrocracking catalyst 18 in a hydrocracking catalyst bed in a second stage reactor 22. The secand stage product stream 24 is first separated in a gas phase 28 and a liquid phase product 30 in a gas/liquid separator, and then fractionated in a fractionator 32 into naphtha 34, jet fuel 36 and an unconverted oil (UCO) fraction 38 dominated by product boiling in the diesel range. A part of the UCO 40 is directed to contact the hydrocracking catalyst and hydrodearomatisation catalyst again. Compared to the process of FIG. 1, this process does not contain a specific HDA step, and thus the only reduction of aromatics will be due ring opening by the HDC catalyst. Accordingly, this process relative to the process of FIG. 1 will result in a lower diesel quality (due to the higher aromatics content and the lower cetane number) as well as a higher yield loss (due to more severe process conditions, selected to increase the extent of ring opening.

[0068] In FIG. 3 according to an embodiment of the prior art, a LCO/gas oil feed mixture in combination with hydrogen 2 is directed to contact a hydrotreatment catalyst 4 and a base metal hydrocracking catalyst 18 in a in a pretreater unit 6, in order to provide a pretreated feed 8. A gas phase 12, including hydrogen sulfide, is removed in a means of gas/liquid separation 10 (such as an interstage stripper or a flash drum), and the pre-treated hydrocarbon feed 14 is combined with a stream of unconverted oil 40 and directed as second stage feed 16 to contact a noble metal hydrodearomatization catalyst 20 The second stage product stream 24 is first separated in a gas phase 28 and a liquid phase product 30 in a gas/liquid separator, and then fractionated in a fractionator 32 into naphtha 34, jet fuel 36 and an unconverted oil (UCO) fraction 38 dominated by product boiling in the diesel range. A part of the UCO 40 is directed to contact the hydrodearomatisation catalyst again.

[0069] In this configuration, the temperature of the hydrocracking catalyst is not independently controlled, as the effluent of hydrotreatment will be directed immediately to the hydrocracking catalyst. In addition, the presence of a high amount of sulfur will result in decreased catalyst activity, which therefore would require an elevated temperaturewith the consequence of reduced ring opening selectivity. This will cause an elevated diesel yield loss at similar dearomatization levels compared to the configuration of FIG. 1.

EXAMPLE 1

[0070] Example 1 discloses operation of HDS, HDC and HDA process in accordance with FIG. 1, i.e. in a configuration where HDC and HDA are operated in the same stage, independently of the HDS stage,

[0071] In Table 1 an example of operation according to the present process scheme (e.g. in accordance with FIG. 1) is shown. The results were consistent over more than 2000 run hours. The example clearly shows the ability to convert LCO/gas oil into a high amount of quality diesel with good cetane properties, even in the presence of moderate levels of sulfur (25 ppm).

EXAMPLE 2

[0072] Example 2 discloses the influence of sulfur impurities in the feed to the second stage of a process such as the one shown in Example 1.

[0073] Table 2 shows the effect of increased H.sub.2S on 2nd stage effluent properties. Experiments A and B were carried out with a commercial base metal HDC catalyst loaded in a first reactor, from which the entire effluent was directed to a second reactor loaded with a commercial noble metal HDA catalyst. From the results it is seen that in this setup with either 17 wppm S or 163 wpp S in the feed mixture, the yield loss (i.e. the fraction boiling below 221 C.) is the same and that the dearomatization is close to complete in both cases as well.

EXAMPLE 3

[0074] Example 3 discloses in Table 3 the influence of carbon monoxide impurities in the feed to the second stage of a process such as the one shown in Example 1.

[0075] Experiments C and D were carried out in a setup similar to that of Example 2but with the catalysts loaded in the same reactor, hence the results are not directly comparable with those of Experiments A and B. Here the experiments shows that the presence of 140 ppm CO results in the same dearomatization, but a slightly increase in yield loss of 1.5%.

EXAMPLE 4

[0076] Example 4 discloses in Table 4 the characteristics of two alternative pretreatments of the same LCO feed, corresponding to reactor 6 of FIGS. 1 and 3 respectively, using the same HDT and HDC catalysts as in Example 1. Hydrotreatment followed by hydrocracking in accordance with reactor 6 of FIG. 3 (the column HDT+HDC) according to the prior art has a yield loss (i.e. the fraction boiling below 221 C.) of 34.7%, whereas hydrotreatment alone in accordance with reactor 6 of FIG. 1 (the column HDT only) only has a yield loss of 10.3%. When examining the effect of the HDC step of example 1, a similar Diesel yield loss is observed, but the dearomatization is much lower. To obtain a dearomatization effect of the HDC step of Example 4 corresponding to that reported in the column HDC eff. of Example 1, would demand a significant further increase in yield loss.

[0077] If the liquid fraction of the effluents of the two pre-treatments of Example 4 would be directed to a process corresponding to Example 1, either directly to HDA (for the combined HDT+HDC effluent according to the prior art) or to HDC/HDA (for the HDT only effluent according to the present disclosure), it is clear that the high dearomatization activity of the base metal HDC catalyst in the semi-sweet environment of reactor 22 of FIG. 1 is able to dearomatize more efficiently at lower or comparable yield loss than the similar catalyst in the sour environment of reactor 6 of FIG. 3.

[0078] The dearomatization observed over the HDC catalyst is assumed to be due to partial cracking processes, in which aromatic rings are opened, without decomposing the molecules further. If the process severity is increased further, such ring opening may be increased, but so will decomposition of the molecules, resulting in increased yield loss.

[0079] The results of Example 4 in combination with Example 1 therefore confirm the benefit of upgrading LCO to diesel in accordance with the present disclosure, as the yield loss for similar dearomatization is lower.

TABLE-US-00001 TABLE 1 HDT 2nd-stg HDC HDA FEED eff FEED eff eff Sulfur, wt ppm 4200 25 17 <1 <1 Nitrogen, wt ppm 1000 2 1 <1 <1 Aromatics, D6591 1 ring, wt % 21.7 52.8 32.7 8.7 0.6 2 rings, wt % 30 2.3 1.5 <0.1 <0.1 3+ rings, wt % 5.9 0.2 0.1 <0.05 <0.05 Fraction <221 C., wt % 11.1 15.9 10.9 35.4 35.8 Yield loss 5.4% 27.5% 0.6%

TABLE-US-00002 TABLE 2 CASE A B Sulfur feed, wt ppm 17 163 Sulfur, wt ppm <1 <1 SG 60/60 F 0.852 0.853 Aromatics, D6591 1 ring, wt % 0.57 0.55 2 rings, wt % <0.1 <0.1 3+ rings, wt % <0.05 <0.05 Fraction <221 C., wt % 24.6 24.3

TABLE-US-00003 TABLE 3 CASE C D CO feed, wt ppm 0 140 Sulfur, wt ppm <1 <1 SG 60/60 F 0.859 0.858 Aromatics, D6591 1 ring, wt % 0.56 0.55 2 rings, wt % <0.1 <0.1 3+ rings, wt % <0.05 <0.05 Fraction <221 C., wt % 19.8 21.3

TABLE-US-00004 TABLE 4 FEED HDT only HDT + HDC Sulfur, wt ppm 3626 <1 <1 Nitrogen, wt ppm 744 <1 <1 Aromatics, D6591 1 ring, wt % 29.1 55.7 30.5 2 rings, wt % 34.2 2.8 0.7 3+ rings, wt % 11.5 0.3 0.1 Yield, wt % FF <221 C. wt % 27.1 34.6 52.4 Yield Loss 10.3% 34.7%