TWO-STAGE HYDROCRACKING PROCESS COMPRISING A HYDROGENATION STAGE DOWNSTREAM OF THE SECOND HYDROCRACKING STAGE, FOR THE PRODUCTION OF MIDDLE DISTILLATES

Abstract

The invention relates to the implementation of a multi-stage hydrocracking process comprising a hydrogenation stage located downstream of the second hydrocracking stage, said hydrogenation stage treating the effluent produced in the second hydrocracking stage, in the presence of a specific hydrogenation catalyst. In addition, the hydrogenation and second hydrocracking stages are implemented under specific operating conditions and particularly under very specific temperature conditions.

Claims

1. Process for producing middle distillates from hydrocarbon feedstocks containing at least 20% by volume of compounds boiling above 340° C., said process comprising the following steps: a) a step of hydrotreating said feedstocks in the presence of hydrogen and at least one hydrotreating catalyst, at a temperature of between 200° C. and 450° C., under a pressure of between 2 and 25 MPa, at a space velocity of between 0.1 and 6 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 100 and 2000 Nl/l, b) a step of hydrocracking at least one portion of the effluent resulting from step a), the hydrocracking step b) taking place, in the presence of hydrogen and at least one hydrocracking catalyst, at a temperature of between 250° C. and 480° C., under a pressure of between 2 and 25 MPa, at a space velocity of between 0.1 and 6 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 80 and 2000 Nl/l, c) a step of high-pressure separation of the effluent resulting from the hydrocracking step b) to produce at least a gaseous effluent and a liquid hydrocarbon effluent, d) a step of distilling at least one portion of the liquid hydrocarbon effluent resulting from step c) performed in at least one distillation column, from which step the following are drawn off: a gaseous fraction, at least one petroleum fraction having at least 80% by volume of products boiling at a temperature below 150° C., at least one middle distillates fraction having at least 80% by volume of products having a boiling point between 150° C. and 380° C., preferably between 150° C. and 370° C. and preferably between 150° C. and 350° C., an unconverted heavy liquid fraction having at least 80% by volume of products having a boiling point above 350° C., preferably above 370° C., preferably above 380° C., e) optionally a purging of at least one portion of said unconverted liquid fraction containing HPNAs, having at least 80% by volume of products having a boiling point above 350° C., before the introduction thereof into step f), f) a second step of hydrocracking at least one portion of the unconverted liquid fraction having at least 80% by volume of products with a boiling point above 350° C., resulting from step d) and optionally purged, said step f) being performed in the presence of hydrogen and of at least one second hydrocracking catalyst, at a temperature TR1 of between 250° C. and 480° C., under a pressure of between 2 and 25 MPa, at a space velocity of between 0.1 and 6 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 80 and 2000 Nl/l, g) a step of hydrogenating at least one portion of the effluent resulting from step f) performed in the presence of hydrogen and of a hydrogenation catalyst, at a temperature TR2 between 150° C. and 470° C., under a pressure of between 2 and 25 MPa, at a space velocity of between 0.1 and 50 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 100 and 4000 Nl/l, said hydrogenation catalyst comprising at least one metal from group VIII chosen from nickel, cobalt, iron, palladium, platinum, rhodium, ruthenium, osmium and iridium alone or as a mixture and not containing any metal from group VIB and a support chosen from refractory oxide supports, and in which the temperature TR2 is at least 10° C. below the temperature TR1, h) a step of high-pressure separation of the effluent resulting from the hydrogenation step g) to produce at least a gaseous effluent and a liquid hydrocarbon effluent, i) recycling, into said distillation step d), at least one portion of the liquid hydrocarbon effluent resulting from step h).

2. Process according to claim 1, in which said hydrocarbon feedstocks are chosen from VGOs or vacuum distillates VDs or gas oils, such as the gas oils resulting from the direct distillation of crude or from conversion units, such as FCC, coker or visbreaking units, and also feedstocks originating from units for the extraction of aromatics from lubricating oil bases or resulting from the solvent dewaxing of lubricating oil bases, or else distillates originating from the desulfurization or hydroconversion of ATRs (atmospheric residues) and/or VRs (vacuum residues), or else from deasphalted oils, or feedstocks resulting from biomass, or any mixture of the abovementioned feedstocks.

3. Process according to claim 1, in which the hydrotreating step a) is performed at a temperature of between 300° C. and 430° C., under a pressure of between 5 and 20 MPa, at a space velocity of between 0.2 and 5 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 300 and 1500 Nl/l.

4. Process according to claim 1, in which the hydrocracking step b) is performed at a temperature of between 330° C. and 435° C., under a pressure of between 3 and 20 MPa, at a space velocity of between 0.2 and 4 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 200 and 2000 Nl/l.

5. Process according to claim 1, in which the hydrocracking step f) is performed at a temperature TR1 of between 320° C. and 450° C., very preferably between 330° C. and 435° C., under a pressure of between 9 and 20 MPa, at a space velocity of between 0.2 and 3 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 200 and 2000 Nl/l.

6. Process according to claim 1, in which said hydrogenation step g) is performed at a temperature TR2 of between 180° C. and 320° C., under a pressure of between 9 and 20 MPa, at a space velocity of between 0.2 and 10 h.sup.−1 and with an amount of hydrogen introduced such that the litre of hydrogen/litre of hydrocarbon volume ratio is between 200 and 3000 Nl/l.

7. Process according to claim 1, in which step g) is performed at a temperature TR2 at least 20° C. lower than the temperature TR1.

8. Process according to claim 7, in which step g) is performed at a temperature TR2 at least 50° C. lower than the temperature TR1.

9. Process according to claim 8, in which step g) is performed at a temperature TR2 at least 70° C. lower than the temperature TR1.

10. Process according to claim 1, in which the hydrogenation step g) is performed in the presence of a catalyst comprising nickel and alumina.

11. Process according to claim 1, in which the hydrogenation step g) is performed in the presence of a catalyst comprising platinum and alumina.

12. The process according to claim 1, wherein said hydrocarbon feedstocks containing at least 80% by volume of compounds boiling above 340° C.

13. The process according to claim 1, wherein said at least one middle distillates fraction has at least 80% by volume of products having a boiling point between 150° C. and 370° C.

14. The process according to claim 1, wherein said at least one middle distillates fraction has at least 80% by volume of products having a boiling point between 150° C. and 350° C.

15. The process according to claim 1, wherein said unconverted heavy liquid fraction have at least 80% by volume of products having a boiling point above 370° C.

16. The process according to claim 1, wherein said unconverted heavy liquid fraction have at least 80% by volume of products having a boiling point above 380° C.

17. The process according to claim 5, wherein hydrocracking step f) is performed at a temperature TR1 of between 330° C. and 435° C.

Description

LIST OF FIGURES

[0122] FIG. 1 illustrates one embodiment of the invention.

[0123] The VGO-type feedstock is sent a via pipe (1) into a hydrotreating step a). The effluent resulting from step a) is sent a via pipe (2) into a first hydrocracking step b). The effluent resulting from step b) is sent a via pipe (3) into a high-pressure separation step c) to produce at least a gaseous effluent (not shown in the FIGURE) and a liquid hydrocarbon effluent which is sent a via pipe (4) into the distillation step d). The following are drawn off in the distillation step d): [0124] a gaseous fraction (5), [0125] at least one petroleum fraction having at least 80% by volume of products boiling at a temperature below 150° C. (6), [0126] at least one middle distillates fraction having at least 80% by volume of products having a boiling point between 150° C. and 380° C. (7), and [0127] an unconverted heavy liquid fraction having at least 80% by volume of products having a boiling point above 350° C. (8).

[0128] Optionally, a portion of the unconverted heavy liquid fraction containing HPNAs is purged in a step e) via pipe (9).

[0129] The optionally purged unconverted heavy liquid fraction is sent via pipe (10) into the second hydrocracking step f). The effluent resulting from step f) is sent a via pipe (11) into a hydrogenation step g). The hydrogenated effluent resulting from step g) is sent a via pipe (12) into a high-pressure separation step h) to produce at least a gaseous effluent (not shown in the FIGURE) and a liquid hydrocarbon effluent which is recycled via pipe (13) into the distillation step d).

EXAMPLES

[0130] The examples that follow illustrate the invention without limiting the scope thereof.

Example No. 1 not in Accordance with the Invention

[0131] A hydrocracking unit treats a vacuum gas oil (VGO) feedstock described in Table 1:

TABLE-US-00001 TABLE 1 Type VGO Flow rate t/h 37 Density — 0.93 Initial boiling point (IBP) ° C. 320 Final boiling point (FBP) ° C. 579 S content wt % 2.71 N content ppm by weight 1510

[0132] The VGO feedstock is injected into a preheating step and then into a hydrotreating reactor under the following conditions set out in Table 2:

TABLE-US-00002 TABLE 2 Reactor R1 Temperature ° C. 385 Total pressure MPa 14 Catalyst — NiMo on alumina HSV h.sup.−1 1.67

[0133] The effluent from this reactor is subsequently injected into a second “hydrocracking” reactor R2 operating under the conditions of Table 3:

TABLE-US-00003 TABLE 3 Reactor R2 Temperature ° C. 390 Total pressure MPa 14 Catalyst — Metal/zeolite HSV h.sup.−1 3

[0134] R1 and R2 constitute the first hydrocracking step, the effluent from R2 is then sent to a separation step composed of a chain for recovery of heat and then high-pressure separation including a recycle compressor and making it possible to separate, on the one hand, hydrogen, hydrogen sulfide and ammonia and, on the other hand, the liquid hydrocarbon effluent feeding a stripper and then an atmospheric distillation column in order to separate streams concentrated in H.sub.2S, a petroleum cut, a middle distillates (kerosene and gas oil) cut, and an unconverted heavy liquid fraction (UCO). A purge corresponding to 2% by weight of the flow rate of the VGO feedstock is taken as distillation bottoms from said unconverted heavy liquid fraction.

[0135] Said unconverted heavy liquid fraction is injected into a hydrocracking reactor R3 constituting the second hydrocracking step. This reactor R3 is used under the following conditions set out in Table 4:

TABLE-US-00004 TABLE 4 Reactor R3 Temperature (TR1) ° C. 340 Total pressure MPa 14 Catalyst — Metal/zeolite HSV h.sup.−1 2

[0136] This second hydrocracking step is carried out in the presence of 100 ppm of equivalent sulfur and 5 ppm of equivalent nitrogen, which originate from the H.sub.2S and NH.sub.3 present in the hydrogen and from the sulfur- and nitrogen-containing compounds still present in said unconverted heavy liquid fraction.

[0137] The effluent from R3 resulting from the second hydrocracking step is subsequently injected into the high-pressure separation step downstream of the first hydrocracking step and then into the distillation step.

Example No. 2 in Accordance with the Invention

[0138] Example 2 is in accordance with the invention insofar as it is a two-step hydrocracking process in which the effluent resulting from the second hydrocracking step is sent into a hydrogenation step in the presence of a hydrogenation catalyst comprising Ni and an alumina support and in which the temperature TR2 in the hydrogenation step is at least 10° C. below the temperature TR1 in the second hydrocracking step.

[0139] The hydrotreating step in R1, first hydrocracking step in R2 and second hydrocracking step in R3 are performed on the same feedstock and under the same conditions as in Example 1. A purge corresponding to 2% by weight of the flow rate of the VGO feedstock is also taken as distillation bottoms from the unconverted heavy liquid fraction.

[0140] A step of hydrogenation of the effluent resulting from R3 is performed in a reactor R4 downstream of R3. The operating conditions for R4 are given in Table 5. In this case, TR2 is 60° C. below TR1.

TABLE-US-00005 TABLE 5 Reactor R4 Temperature (TR2) ° C. 280 Total pressure MPa 14 Catalyst — Ni/Alumina HSV h.sup.−1 2

[0141] The catalyst used in the reactor R4 has the following composition: 28 wt % Ni on gamma alumina.

[0142] The hydrogenated effluent resulting from R4 is then sent into a high-pressure separation step before being recycled into the distillation step.

Example No. 3 in Accordance with the Invention

[0143] Example 3 is in accordance with the invention insofar as it is a two-step hydrocracking process in which the effluent resulting from the second hydrocracking step is sent into a hydrogenation step in the presence of a hydrogenation catalyst comprising Pt and an alumina support and in which the temperature TR2 in the hydrogenation step is at least 10° C. below the temperature TR1 in the second hydrocracking step.

[0144] The hydrotreating step in R1, first hydrocracking step in R2 and second hydrocracking step in R3 are performed on the same feedstock and under the same conditions as in Example 1. A purge corresponding to 2% by weight of the flow rate of the VGO feedstock is also taken as distillation bottoms from the unconverted heavy liquid fraction.

[0145] A step of hydrogenation of the effluent resulting from R3 is performed in a reactor R4 downstream of R3. The operating conditions for R4 are given in Table 6. In this case, TR2 is 80° C. below TR1.

TABLE-US-00006 TABLE 6 Reactor R4 Temperature (TR2) ° C. 260 Total pressure MPa 14 Catalyst — Pt/Alumina HSV h.sup.−1 2

[0146] The catalyst used in the reactor R4 has the following composition: 0.3 wt % Pt on gamma alumina.

[0147] The hydrogenated effluent resulting from R4 is then sent into a high-pressure separation step before being recycled into the distillation step.

Example No. 4 not in Accordance with the Invention

[0148] Example 4 is not in accordance with the invention in so far as it is a two-step hydrocracking process in which a step of hydrogenation in the presence of a hydrogenation catalyst comprising Pt and an alumina support is carried out upstream of the second hydrocracking step and in which the temperature TR2 in the hydrogenation step is equal to the temperature TR1 of the second hydrocracking step.

[0149] The hydrotreating step in R1, first hydrocracking step in R2 and second hydrocracking step in R3 are performed on the same feedstock and under the same conditions as in Example 1. A purge corresponding to 2% by weight of the flow rate of the VGO feedstock is also taken as distillation bottoms from the unconverted heavy liquid fraction. This time, the unconverted heavy liquid fraction resulting from the distillation is first sent to a hydrogenation step carried out in a reactor R4 upstream of R3. In this case, TR2 in the hydrogenation step is equal to the temperature TR1 in the second hydrocracking step and is 340° C. The operating conditions of R4 are stated in Table 7.

TABLE-US-00007 TABLE 7 Reactor R4 Temperature (TR2) ° C. 340 Total pressure MPa 14 Catalyst — Pt/Alumina HSV h.sup.−1 2

[0150] The catalyst used in the reactor R4 has the following composition: 0.3 wt % Pt on gamma alumina.

[0151] The hydrogenated effluent resulting from R4 is then sent to the second hydrocracking step carried out in the reactor R3 before being sent to the high-pressure separation then being recycled to the distillation step.

Example No. 5 in Accordance with the Invention

[0152] Example 5 is in accordance with the invention insofar as it is a two-step hydrocracking process in which the effluent resulting from the second hydrocracking step is sent into a hydrogenation step in the presence of a hydrogenation catalyst comprising Pt and an alumina support and in which the temperature TR2 in the hydrogenation step is at least 10° C. below the temperature TR1 in the second hydrocracking step.

[0153] The hydrotreating step in R1, first hydrocracking step in R2 and second hydrocracking step in R3 are performed on the same feedstock and under the same conditions as in Example 1. A purge corresponding to 2% by weight of the flow rate of the VGO feedstock is also taken as distillation bottoms from the unconverted heavy liquid fraction.

[0154] A step of hydrogenation of the effluent resulting from R3 is performed in a reactor R4 downstream of R3. The operating conditions for R4 are given in Table 8. In this case, TR2 is 60° C. below TR1.

TABLE-US-00008 TABLE 8 Reactor R4 Temperature ° C. 280 Total pressure MPa 14 Catalyst — Pt/Alumina HSV h.sup.−1 3

[0155] The catalyst used in the reactor R4 has the following composition: 0.3 wt % Pt on gamma alumina.

[0156] The hydrogenated effluent resulting from R4 is then sent into a high-pressure separation step before being recycled into the distillation step.

Example No. 6 not in Accordance with the Invention

[0157] Example 6 is not in accordance with the invention in so far as it is a two-step hydrocracking process in which a step of hydrogenation in the presence of a hydrogenation catalyst comprising Pt and an alumina support is carried out upstream of the second hydrocracking step and in which the temperature TR2 in the hydrogenation step is 60° C. below the temperature TR1 in the second hydrocracking step.

[0158] The hydrotreating step in R1, first hydrocracking step in R2 and second hydrocracking step in R3 are performed on the same feedstock and under the same conditions as in Example 1. A purge corresponding to 2% by weight of the flow rate of the VGO feedstock is also taken as distillation bottoms from the unconverted heavy liquid fraction. This time, the unconverted heavy liquid fraction resulting from the distillation is first sent to a hydrogenation step carried out in a reactor R4 upstream of R3. In this case, TR2 in the hydrogenation step is 60° C. below the temperature TR1 in the second hydrocracking step and is 280° C. The operating conditions of R4 are stated in Table 9.

TABLE-US-00009 TABLE 9 Reactor R4 Temperature (TR2) ° C. 280 Total pressure MPa 14 Catalyst — Pt/Alumina HSV h.sup.−1 3

[0159] The catalyst used in the reactor R4 has the following composition: 0.3 wt % Pt on gamma alumina.

[0160] The hydrogenated effluent resulting from R4 is then sent to the second hydrocracking step carried out in the reactor R3 before being sent to the high-pressure separation then being recycled to the distillation step.

Example No. 7 in Accordance with the Invention

[0161] Example 7 is in accordance with the invention insofar as it is a two-step hydrocracking process in which the effluent resulting from the second hydrocracking step is sent into a hydrogenation step in the presence of a hydrogenation catalyst comprising Ni and an alumina support and in which the temperature TR2 in the hydrogenation step is at least 10° C. below the temperature TR1 in the second hydrocracking step.

[0162] The hydrotreating step in R1, first hydrocracking step in R2 and second hydrocracking step in R3 are performed on the same feedstock and under the same conditions as in Example 1. This time, a purge corresponding to 1% by weight of the flow rate of the VGO feedstock is taken as distillation bottoms from the unconverted heavy liquid fraction.

[0163] A step of hydrogenation of the effluent resulting from R3 is performed in a reactor R4 downstream of R3. The operating conditions for R4 are given in Table 10. In this case, TR2 is 60° C. below TR1.

TABLE-US-00010 TABLE 10 Reactor R4 Temperature (TR2) ° C. 280 Total pressure MPa 14 Catalyst — Ni/Alumina HSV h.sup.−1 2

[0164] The catalyst used in the reactor R4 has the following composition: 28 wt % Ni on gamma alumina.

[0165] The hydrogenated effluent resulting from R4 is then sent into a high-pressure separation step before being recycled into the distillation step.

Example 9: Process Performance

[0166] Table 11 summarizes the performance of the processes described in Examples 1 to 7 in terms of yield of middle distillates, cycle time of the process, cetane number of the gas oil fraction obtained, and overall conversion of the process. The conversion of coronene (HPNA containing 7 aromatic rings) performed in the hydrogenation step is also reported.

TABLE-US-00011 TABLE 11 1 (not in 2 (in 3 (in 4 (not in 5 (in 6 (not in 7 (in accordance accordance accordance accordance accordance accordance accordance with the with the with the with the with the with the with the Examples invention) invention) invention) invention) invention) invention) invention) Scheme R3 alone R3 + R4 R3 + R4 R4 + R3 R3 + R4 R4 + R3 R3 + R4 Catalyst in — 28% Ni/ 0.3% Pt/ 0.3% Pt/ 0.3% Pt/ 0.3% Pt/ 28% Ni/ R3 alumina alumina alumina alumina alumina alumina Purge (%) 2 2 2 2 2 2 1 TR1 (° C.) 340 340 340 340 340 340 340 TR2 (° C.) — 280 260 340 280 280 280 Coronene 0 86 94 7 81 67 86 conversion (%) (1) Yield of Base Base Base Base Base Base Base + 1 middle point distillates Cycle time Base Base + 6 Base + 7 Base + 1 Base + 4 Base + 3 Base + 4 months months months months months months Cetane Base Base + 4 Base + 5 Base Base + 3 Base Base + 4 number points points points points Overall 98 98 98 98 98 98 99 conversion (%) [0167] (1) The coronene conversion is calculated by dividing the difference in the amounts of coronene measured upstream and downstream of the hydrogenation reactor by the amount of coronene measured upstream of this same reactor. The amount of coronene is measured by high-pressure liquid chromatography coupled to a UV detector (HPLC-UV), at a wavelength of 302 nm for which coronene has maximum absorption.

[0168] These examples illustrate the advantage of the process according to the invention which makes it possible to obtain improved performances in terms of yield of metal distillates, cycle time, overall conversion of the process or cetane number of the gas oil fraction obtained.

[0169] Thus, with the process of Example 2 using a hydrogenation reactor downstream of the second hydrocracking step, the cycle time is lengthened by 6 months relative to a process without a hydrogenation reactor (illustrated by Example 1) and the cetane number of the gas oil fraction is increased by 4 points. Specifically, at 280° C., the Ni/alumina hydrogenation catalyst makes it possible to greatly convert the aromatic compounds and in particular the HPNAs. The deactivation of the catalyst of the second hydrocracking step is therefore slowed down, which allows a longer cycle. Since the aromatic compounds of the gas oil fraction are hydrogenated, the cetane number is improved.

[0170] Examples 3 and 5 show the effect of the temperature of the hydrogenation reactor on the conversion of the aromatic compounds and the HPNAs, with their impact on the cycle time and the quality of the gas oil obtained.

[0171] Conversely, with the processes not in accordance with the invention, of Examples 4 and 6, the performances are much worse: the hydrogenation reactor located upstream of the second hydrocracking reactor makes it possible to convert the HPNAs (with a strong temperature dependency), but since the hydrocarbon feedstock processed in this reactor has not yet been cracked, the effect on the hydrogenation of the aromatic compounds of the gas oil fraction is not obtained and the cetane number is not improved.

[0172] Example 7 illustrates that the process according to invention also makes it possible to reduce the degree of purge, since the HPNAs are hydrogenated in the hydrogenation reactor, which leads to an increase in the overall conversion and in the yield of middle distillates, while retaining a lengthened cycle time and an improved cetane number.