Method of treating a hydrocarbonaceous raw material with hydrogen

Description of the invention

The field of technology

The present invention relates to methods of upgrading the middle distillate boiling fraction that is obtained from VGO hydrotreating or moderate severity hydrotreating reactors. The invention encompasses a multi-step single-loop hydrogen process.

State of the art

Crude oil refining processes use vacuum gas oil hydrotreating (VGO) reactors and hydrocracking reactors to remove contaminants such as sulfur, nitrogen and metals from crude oil. Typically, the medium distillate boiling material (250 ° F-735 ° F) from VGO hydrotreating or moderate severity hydrocrackers does not meet the smoke point, cetane number, and flavor specifications. In most cases, this middle distillate is separately upgraded in a moderate strength hydrotreating reactor, or alternatively the middle distillate is formulated into a general pool of diesel fuels or used as domestic fuel oil. There are also jets in the diesel boiling range from other units such as fluid catalytic cracking, coking or visbreaking that require upgrading. Very often, existing diesel hydrotreating reactors are not designed for the pressure limits required to process these streams and the mild hydrocracking unit provides the possibility of upgrading these streams simultaneously.

Some processes have previously been disclosed in which the hydroprocessing occurs in a single hydrogen loop. International Patent Application WO 97/38066 (PCT / US97 / 04270), published October 16, 1997, discloses a process with an inverse configuration of stages in hydroprocessing reactor systems. This hydroprocessing reactor system comprises two reaction zones stacked on top of each other in one reaction loop. In a preferred embodiment, the hydrocarbon feed is fed to the denitrification and desulfurization zone, which is the downstream zone. The effluent from this zone cools and separates the gases therefrom. The liquid product is then fed to the top zone where hydrocracking or hydrotreating can take place. The deeper processing preferably takes place in the upper zone.

U.S. Patent No. 5,980,729 discloses a configuration similar to that of WO 97/38066. However, a hot stripper is located downstream of the denitrification / desulfurization zone. This evaporator is followed by an additional hydrotreating reactor. Also located downstream from the denitrification / desulfurization zone is a finish refining zone for saturating aromatics. U.S. Patent No. 6,106,694 discloses a configuration similar to U.S. Patent No. 5,980,729, but without a stripper hydrotreating reactor and no finish refining zone.

Brief Description of the Invention

In the process of the invention, the middle distillate is hydrotreated in the same high pressure loop as the vacuum gas oil hydrotreating or moderate severity hydrocracking reactor, but without using the reverse stage configuration used in the cited prior art publications. The capital cost savings and / or operating savings due to the use of a single hydrogen loop are significant as no separate middle distillate hydrotreating reactor is required. Other advantages include optimal hydrogen pressures for each stage as well as optimal hydrogen consumption and utilization of each product. The efficiency of the upgraded product is also maximum, without the use of a recirculating liquid. The invention is discussed below.

The hydroprocessing method of a hydrocarbon feed using at least two reaction zones in a single reaction loop comprises the following steps:

(a) feeding a hydrocarbon feed to a first hydroprocessing zone having one or more catalytic beds containing a hydroprocessing catalyst, the hydroprocessing zone being maintained under hydroprocessing conditions in which the feed is in contact with the catalyst and the hydrogen;

(b) feeding the effluent from step (a) directly to a hot high pressure separator where the effluent is contacted with a hot hydrogen rich stripping gas to produce a vapor stream containing hydrogen, hydrocarbon compounds boiling below the boiling range of the hydrocarbon feed, hydrogen sulfide and ammonia , and a liquid stream containing hydrocarbon compounds boiling approximately in the range of said hydrocarbon feed;

(C) feeding the steam stream from step (b), after cooling and partial condensation, to a hot hydrogen stripper containing at least one hydrotreating catalyst bed in which it is in countercurrent contact with the hydrogen and the liquid stream from step (b) ) is fed to fractionation;

(d) feeding the steam stream from the hot hydrogen stripper of step (c), after cooling and contacting with water, the vapor overhead stream containing hydrogen, ammonia and hydrogen sulfide, together with light gases and gasoline, to a cold separator high-pressure, in which hydrogen, hydrogen sulfide and light hydrocarbon gases are removed from the overhead, ammonia is removed from the cold high-pressure separator as ammonium hydrogen sulfide in the acid water stripper, and gasoline and middle distillates are fed to the fractionation;

(e) feeding the hot hydrogen stripper liquid stream from step (cc to a second hydroprocessing zone, the second hydroprocessing zone containing at least one hydroprocessing catalyst bed suitable for saturating aromatics and ring opening, wherein the liquid is contacted under the treatment conditions) hydrogen with a hydroprocessing catalyst in the presence of hydrogen;

(f) feeding the sys- tem from the high pressure separator of the eeap (d) to an absorber in which the hydrogen sulfide is removed before the hydrogen is compressed and returned to the hydrogen treatment vessels in the loop; and (g) feeding the effluent from step (e) to the cold high pressure separator of step (d).

Brief description of the drawings

Figure 1 illustrates the hydroprocessing loop in which the finishing reactor is a middle distillate upgrading reactor that operates at approximately the same pressure as the first stage reactor.

Figure 2 illustrates the hydroprocessing loop where the finishing reactor is the same as that of Figure 1 but operating at a lower pressure than the first stage reactor. A noble metal based catalyst is used in the finishing reactor.

Detailed description of the invention

Description of the preferred embodiment

Description of Figure 1

The feed from stream 1 is mixed with recycle hydrogen and make up hydrogen in stream 42. The feed is preheated in the process heat exchanger system, as are the gaseous streams. The feed and gas mixture, now stream 34, is further heated using heat exchangers 43 and a furnace 49. Stream 34 then enters the first stage 2 fixed bed and downflow reactor. First bed 3 of reactor 2 may contain a VGO hydrotreating catalyst or a moderate severity hydrocracking catalyst. The fixed beds 3 may be sequentially arranged, with interstage cooling jets 4 and 5 supplying hydrogen between the beds.

The effluent 6 of the 1st stage reactor 2, after hydrotreating or partial hydrocracking, contains hydrogen sulfide, ammonia, light gases, gasoline, middle distillate and hydrotreated vacuum gas oil. The effluent enters the hot high pressure separator or deaeration zone 8 operating under heavy oil reactor effluent conditions, where some of the diesel fuel and most of the lighter material are separated from the unconverted oil. The hot high pressure separator has a stack of trays 44, and a hydrogen-rich gas is introduced at the bottom for stripping through stream 46.

Stream 9 is primarily a heavy hydrotreated gas oil boiling above 700 ° F. The valve 10 indicates that the pressure is released in stream 11 prior to sending the unconverted oil to the fractionation section.

Stream 21 contains an overhead from a hot high pressure separator. Stream 21 is then cooled in the exchanger 22 (by steam generation or process heat exchange) before entering the hot hydrogen stripper / reactor 23. Stream 21 flows downstream through a hydrotreating catalyst bed 52 contacting countercurrent hydrogen from stream 51.

Overhead stream 26 contains hydrogen, ammonia, and hydrogen sulfide, and light gases and gasoline. The operating pressure difference between the hot hydrogen stripper / reactor 23 and the cold high pressure separator 17 is maintained by the control valve 50. Stream 26 is cooled in exchanger 27 and merges with stream 14 to form stream 16. To stream 16

Water is injected (stream 36) to remove most of the ammonia as an ammonium bisulfide solution (ammonia and hydrogen sulfide react to form ammonium bisulfide, which is converted into a solution by water injection). The stream is then air-cooled by a cooler 45. Stream 16 enters a cold high pressure separator 17. Hydrogen, light hydrocarbon gases, and hydrogen sulfide are removed at the overhead via stream 19. Hydrogen sulfide is removed from the stream in hydrogen sulfide absorber 20. Ammonia and hydrogen sulfide are removed with an acidic water stream (not shown) from a cold high pressure separator 17.

Stream 40, which contains the hydrogen-rich gas, is compressed in compressor 30 and is split into streams 29 and 32. Stream 32 enters a hot hydrogen stripper / reactor 23. Stream 31 is separated from stream 29 for use as an interstage coolant. Streams 4 and 5 are separated from stream 31. Stream 29, containing hydrogen, is combined with stream of hydrogen 42 before being combined with feed oil stream 1.

Make-up hydrogen 38 is compressed and sent to four separate locations, upstream from reactor 2 for connection with feed stream 1 (via stream 42), to hot high pressure separator 8 via stream 46, to the hot hydrogen stripper / reactor via stream 51. and to a middle distillate upgrading reactor (stream 35) for combination with recirculating gas oil or kerosene or for use as an interstage coolant. Stream 38, containing make-up hydrogen, is passed to make-up hydrogen compressor 37. Streams 35, 42, and 46 are separated from compressor 37 leaving pressurized hydrogen.

The middle distillate upgrading reactor 12 consists of one or more beds 13 of a hydrotreating / hydrocracking catalyst (such as Ni-Mo, Ni-W and / or noble metal) to saturate the aromas and open the rings to improve the quality characteristics of the diesel fuel such as oil level. aromas and cetane number. In the embodiment shown in Fig. 1 the middle distillate upgrading reactor is operated at approximately the same pressure as the first stage reactor 2. A cooling gas (stream 47) can be introduced to control the temperature of the reactor. Stream 24 may be combined with recirculating diesel or kerosene (stream 48) from the fractionator when no external streams (stream 7) are to be processed or cooled in exchanger 25. The hydrogen from stream 35 is combined with stream 24 prior to its introduction to the middle distillate upgrading reactor. Stream 24 enters the reactor from the top and flows down through the catalyst beds 13.

Stream 14, which is the effluent from the middistillate upgrading reactor 12, is used to heat other process streams in the unit (see exchanger 15) and then merges with stream 26 to form stream 16 which is sent to the effluent air cooler and then to the cold separator high pressure 17. Water is continuously injected into the inlet pipe of the leachate air cooler to prevent salt build-up in the air cooler tubes. Hydrogen, hydrogen sulfide, and ammonia exit the cold high pressure separator via overhead stream 19, while naphtha and middle distillates exit via stream 18 for fractionation (stream 39).

Description of Figure 2

As described for Figure 1, feed in stream 1 is mixed with recycle hydrogen and make-up hydrogen in stream 42. The feed and gaseous streams are preheated in a process heat exchanger system. The feed / gas mixture, now in stream 34, is further heated using heat exchangers 43 and furnace 51. Stream 34 then enters the fixed bed, downflow stage 2 reactor. First bed 3 of reactor 2 may contain a VGO hydrotreating catalyst or a moderate severity hydrocracking catalyst. There may be a series of consecutive fixed beds 3, with interstage cooling jets 4 and 5 supplying hydrogen between the beds.

The effluent 6 of the first stage hydrotreated or partially hydrocracked reactor comprises hydrogen sulfide, ammonia, light gases, gasoline, middle distillate and hydrotreated vacuum gas oil. The effluent enters the hot high pressure separator or deaeration zone 8 operating under heavy oil reactor effluent conditions, where some of the diesel fuel and most of the lighter material are separated from the unconverted oil.

The hot high pressure separator has a stack of trays 44, and a hydrogen-rich gas is introduced at the bottom for stripping through stream 46.

PL 198 388 B1

Stream 9 is primarily a heavy hydrotreated gas oil boiling above 700 ° F. The valve 10 indicates that the pressure is released in stream 11 prior to sending the unconverted oil to the fractionation section.

Stream 21 contains an overhead from the hot high pressure separator and can be made up by external feed 7. Stream 21 is then cooled in exchanger 22 (by steam generation or process heat exchange) before entering the hot hydrogen stripper / reactor 23. Stream 21 flows downward through a hydrotreating catalyst bed 52 contacting countercurrent hydrogen from stream 32.

The overhead stream 26 from the hot hydrogen stripper / reactor 52 contains hydrogen, ammonia, and hydrogen sulfide, as well as light gases and gasoline. It is cooled in an exchanger 27. Water (stream 36) is injected into stream 26 to remove most of the ammonia as the ammonium bisulfide solution (ammonia and hydrogen sulfide react to form ammonium bisulfide which is converted into a solution by spraying water). The stream is then air-cooled by a cooler 45. The air cooler effluent enters the cold high pressure separator 17. Hydrogen, light hydrocarbon gases and hydrogen sulfide are removed overhead with a stream 19. Hydrogen sulfide (stream 51) is removed from the stream in a hydrogen sulfide absorber 20. Ammonia and hydrogen sulfide are removed with an acid water stream (stream 48) from a cold high pressure separator 17. Stream 40, which contains hydrogen, is pressurized in compressor 30 and split into streams 29 and 31. Stream 31 is separated from stream 29 for use as an interstage coolant. Streams 4 and 5 are separated from stream 31. Stream 29, containing hydrogen, is combined with stream of hydrogen 42 before being combined with feed oil stream 1.

Make-up hydrogen 38 is compressed and sent to four separate locations, upstream of reactor 2 for connection with feed stream 1 (via stream 42), to hot high pressure separator 8 via stream 46, to the hot hydrogen stripper / reactor 23, and to the reactor upgrading a middle distillate (stream 35) for combination with recirculating diesel or kerosene or for use as an intercooled. Stream 38, containing make-up hydrogen, is passed to make-up hydrogen compressor 37. Streams 35, 42, and 46 are separated from stream 41 that exits compressor 37, containing pressurized hydrogen.

In this embodiment, the middle distillate upgrading reactor 12 is operated at a pressure lower than the first stage 2 reactor. The liquid pressure (stream 24) from the hot hydrogen stripper 52 is reduced (via valve 28) and is combined with the make-up hydrogen (stream 35) after the second make-up hydrogen compression step in compressor 37. At this point, recirculating diesel fuel or kerosene (stream 50) may be added. After preheating (in exchanger 25), the mixture is sent to the middle distillate upgrading reactor 12, which is preferably loaded with one or more noble metal catalyst beds 13. Part of the make-up hydrogen is available as coolant stream (stream 47) between the beds in when multiple beds are used. The reactor effluent (stream 14) is cooled in a series of heat exchangers 15 and sent to a cold high pressure separator 49.

The overhead vapors 38 from the cold high pressure separator 49 are substantially high purity hydrogen with a small amount of light hydrocarbon gases. Steam is sent to make-up hydrogen compressor 37. Compressed make-up hydrogen (stream 29) is sent to high pressure reactor 2, high pressure separator 8, and hot hydrogen stripper / reactor 23. The still bottoms (stream 18) from the cold high pressure separator 17 after pressure reduction are sent to the fractionation section (stream 53).

Stream 14, which is the effluent from the middle distillate upgrading reactor 12, is used to heat the other process streams in the unit (see exchanger 15) and enters the cold high pressure separator 49. The liquid effluent from the cold high pressure separator 49, stream 39, goes to fractionation.

Raw Materials

Various types of hydrocarbon feedstocks may be used in the present invention. Typical feeds include any heavy or synthetic oil fractions or process streams having a boiling point above 300 ° F (150 ° C). Such feeds include vacuum gas oils, heavy atmospheric gas oils, coking gas oils, visbreaking gas oils, demetallised oils, vacuum residues, atmospheric residues, deasphalted oil, Fischer-Tropsch streams, FCC streams, etc.

PL 198 388 B1

Typical feeds for the first stage reaction will be vacuum gas oils, heavy coke oven gas oils or deasphalted oil. Lighter feeds such as straight run diesel, light recycle oil, light coke oil or visbreak gas oil can be introduced upstream from the hot hydrogen stripper / reactor 23.

Products

Figures 1 and 2 show two different embodiments of the present invention, primarily relating to the production of high quality middle distillate and the production of heavy hydrotreated gas oil.

The process of this invention is particularly useful for producing middle distillate fractions boiling in the range of about 250^ABOUTF-700 ° F (121 ° C-371 ° C). The middle fraction of the distillate is defined as having a boiling range from about 250 ° F to 700 ° F. At least 75 vol%, preferably 85 vol%, of the components in the middle distillate have a normal boiling point above 250 ° F. At least about 75 vol%, preferably 85 vol%, of the components in the middle distillate have a normal boiling point below 700 ° F. The term "middle distillate" includes gas oil, jet oil and kerosene boiling fractions. The boiling point of kerosene or jet oil is between 280 ° F and 525 ° F (138 ° C-274 ° C). The term "diesel boiling range" refers to hydrocarbons boiling in the range of 250 ° F to 700 ° F (121 ° C-371 ° C).

Gasoline or light petrol can also be produced in the process of the invention. Gasoline or light gasoline normally boils below 400 ° F (204 ° C) or C.₅-. The boiling ranges of the various products obtained at any given refinery will depend on factors such as the characteristics of the crude oil source, local refinery markets and product prices.

Heavy diesel fuel, another product of the present invention, typically boils in the range of 550 ° F to

750 ° F.

Conditions

Hydrotreating conditions is a general term that refers primarily in this application to hydrocracking or hydrotreating, preferably hydrocracking. The first stage reactor, as depicted in Figures 1 and 2, may be either a VGO hydrotreating reactor or a moderate severity hydrocracking reactor.

The hydrotreating conditions include a reaction temperature in the range 400 ° F-900 ° F (204 ° C-482 ° C), preferably 650 ° F-850 ° F (343 ° C-454 ° C); pressure 500 to 5000 psig (pounds per square inch gau^ge) (^3,5-³4^{6 Mp}and) ^korz^ys^voln^ande ^{1000 d}about ^{3000 p}s^ig (7,0-20,⁸ M.^pand)^; ch^ybkabout^life zasharna (LH^SV) ^{0.5 h}¹up to 20 h1 (v / v); and a total hydrogen consumption of 300 to 5000 scf per barrel of liquid hydrocarbon feed (^53,4-³⁵⁶ m^3/m³ wsa^dat).

In the embodiment shown in Figure 1, the first stage reactor and the middle distillate upgrading reactor are operated at the same pressure. In the embodiment shown in Figure 2, the middle distillate upgrading reactor is operated at a pressure lower than the first stage reactor.

Typical hydrocracking conditions include a reaction temperature of 400 ° F-950 ° F (204 ° C-510 ° C), preferably 650 ° F-850 ° F (343 ° C-454 ° C). The reaction pressure is between 500 and 5000 psig (3.5-34.5 MPa),^korz^ys^voln^ande ^{1500 d}about ^{3500 p}s^ig ABOUT^0,4-^{24.2 Mp}and). ^LHSV jes^vol in the field of Fr.^{d 0.1 d}about ^{15 h_1} (Wv) preferably 0.25-2.5 h1. The hydrogen consumption is in the range of 500 to 2500 scf per barrel of liquid coal feed^dorowe^gabout (^89,1-⁴⁴⁵ m^{3 H.}'7m³ wsa^dat).

Catalyst

The hydroprocessing zone may contain only one catalyst or several catalysts in combination.

The hydrocracking catalyst generally comprises a cracking component, a hydrogenation component, and a carrier. Such catalysts are well known in the art. The cracking component may include an amorphous silica / alumina phase and / or a zeolite such as a Y-type or USY zeolite. REX, REY and USY zeolites are often used in catalysts having a high cracking activity. The binding agent is usually silica or alumina. The hydrogenation component will be a Group VI, Group VII or Group VIII metal, or an oxide or sulphide of this metal, preferably one or more of molybdenum, tungsten, cobalt or nickel, or their oxides or sulphides. These hydrogenation components, if present, typically constitute from about 5% to about 40% by weight of the catalyst. Alternatively, metals of the platinum group, especially platinum and / or palladium, may be present as the hydrogenation component, either alone or in combination with the non-noble metal hydrogenation components molybdenum, tungsten, cobalt or nickel. Platinum group metals, if present, will generally comprise from about 0.1% to about 2% by weight of the catalyst.

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The hydrotreating catalyst, if used, will typically be a composite of a Group VI metal or a compound thereof, deposited on a refractory porous substrate such as alumina. Examples of hydrotreating catalysts are alumina-supported cobalt-molybdenum, nickel sulfide, nickel-tungsten, cobalt-tungsten and nickel-molybdenum. Typically such hydrotreating catalysts are pretreated with sulphides.

Example

Post-refining of distillates from a mild hydrocracking reactor to improve the cetane number

<td>Batch</td><td>A distillate from a mild hydrocracking blend vacuum gas oil / gas oil from coking</td><td>The distillate from mild hydrocracking vacuum gas oil from mid-eastern crude oil</td>

<td>Mild Hydrocracking Conversion</td><td>30% liquid by volume &lt;680 ° F</td><td>31 volumetric liquids &lt;700 ° F</td>

<td>Hydrotreating catalyst</td><td>Precious metal / zeolite</td><td>Common metal / aluminum oxide</td>

<td>Hydroco-refining conditions: Catalyst bed temperature,</td><td> 594</td><td> 720</td>

<td>° F LHSV, 1 / h</td><td> 1,5</td><td> 2,0</td>

<td>Gas / Oil ratio SCF / B</td><td> 3000</td><td> 5000</td>

<td>Partial pressure of H2, psia</td><td> 800</td><td> 1900</td>

<td>Increasing the cetane number (typical)</td><td>7 to 15</td><td>2 to 7</td>

The above table illustrates the efficiency of upgrading the effluent from a first stage reactor that has been subjected to mild hydrocracking. The effluent is subjected to hydrotreatment in a middle distillate upgrading reactor. The increase in cetane number (improvement) is greater and under less stringent conditions when using a catalyst having a noble metal hydrogenation component and a zeolite cracking component as a catalyst than when using a catalyst having base metal hydrogenation components on alumina, an amorphous support. The cetane increase may be higher if an external diesel range charge (7) is added upstream of the high pressure hot separator 44 upstream of the process.