GB1604179A - Distillate hydrogenation process - Google Patents

Description

(54) DISTILLATE HYDROGENATION PROCESS (71) I, ROBERT LEARD IRVINE of Rob Nes, Pyle Hill, Woking, Surrey, a citizen of the United States of America; do hereby declare the invention, for which I pray that a patent may be granted to me, and the method by which it is to be performed, to be particularly described in and by the following statement: The present invention relates to distillate hydrogenation process.

A similar upflow type of reactor system to that disclosed and claimed in UK Patent Number 1,239,972, together with economy of hydrogen circulation as disclosed and claimed in UK Patent Number 1,294,239, is employed, but the present invention is directed towards distillate hydrogenation applications. The resulting reactor system has been found particularly suitable for application to distillates which are hydrogenated, with chemical hydrogen consumption being above 350 standard cubic feet per barrel, and in which the distillate remains substantially in the liquid phase in the reactor at the low hydrogen circulation rates of the present invention. As extensive aromatic saturation requires a high hydrogen partial pressure, such distillate applications are particularly suitable.

Accordingly, the present invention provides a process for the hydrogenation of distillates which comprises the steps of: i) passing the crude hydrocarbon feed and hydrogen through a plurality of fluidised catalyst bed reactor stages operated in a progressive temperature pattern with the lowest bed outlet temperature at the feed inlet and the highest bed outlet temperature at the final reactor outlet, said plurality of fluidised catalyst bed reactor stages having a liquid phase maintained at a near hydrogen-saturated state, and the temperature of the inlet feed in the first reactor stage being regulated to control the outlet stream of said first reactor stage; and ii) passing the products from the outlet of the final reactor stage to a separator stage.

Distillates have insignificant metal contents and also limited Ramsbottom carbon residue precursors as compared with crude or residue feedstocks. The number of stages may therefore be reduced with the individual stage inventories being directed towards decreasing the utilities as compared with those obtained with existing downflow fixed bed reactor systems. Catalyst replacement is not generally a problem since commercial downflow fixed beds have experienced onstream periods in excess of three years for feedstocks similar to the two described in the illustrative Examples which are discussed hereinater and compared with the process of the present invention from the viewpoints of hydrogen circulation, high pressure heat-exchange surface and the utilities required.

Catalyst replacement considerations only become important for feedstocks such as higher boiling lubestocks where downflow fixed beds must normally regenerate catalyst between scheduled maintenance requirements. In comparison with prior art methods, it should be stated that, at the same hydrogen partial pressure and reactor inventory and, for a given product quality from a given feedstock, the process of the present invention provides a more gradual decline in catalyst activities because of the lower maximum temperatures in the feed entry beds which contact the most reactive feed and also in subsequent beds.

The process of the present invention normally provides the entire hydrogen requirement for the feed inlet from the makeup hydrogen required, which itself provides both the hydrogen chemically consumed and the hydrogen in solution in the reactor liquid product streams leaving the high pressure separator. Conventional downflow fixed bed reactor systems require comparatively large volumes of gas per unit volume of feed to adequately distribute the feed, which remains substantially in the liquid phase at reactor inlet conditions, and to provide sufficient turbulence for good kinetics. The process of the present invention, however, requires a gas phase only to ensure that the hydrogen is being consumed within the stage by chemical reactions, and to maintain the liquid phase in a near hydrogen-saturated state.In the process of the present invention, the liquid phase serves as the primary distribution medium (incomplete catalyst wetting simply cannot occur in an upflow bed which has a substantial liquid phase), and provides the required turbulence through fluidization of the catalyst particles since the bed is normally partially expanded by at least 6 percent over the settled bed height.

Several stages are provided in the process of the present invention and these stages are operated in a progressive temperature pattern with the lowest bed outlet temperature at the feed inlet and the highest bed outlet temperature at the final reactor outlet. Normally, hydrogen recycle is injected interstage to remove the heat of the reaction and to maintain the bed outlet of the stages following the feed bed entry stage. For the feed entry inlet stage, the temperature of the incoming feed is regulated to control the outlet stream of the feed entry stage. In the two illustrative Examples below, such control is achieved by bypassing part of the feed around the feed inlet-reactor product heat exchangers.

The above-mentioned temperature pattern is possible in the process of the present invention using the same catalyst inventory that a conventional fixed bed would employ because the average bed temperature in the process of the present invention is always close to the controlled bed outlet temperature as a result of fluidized heat transfer, i.e. the average bed temperature is practically independent of the entering temperature in the process of the present invention. This contrasts with a static bed process wherein the average bed temperature is a function not only of the entering inlet feed temperature to the bed but also of the outlet bed temperature. The downflow static bed reactor, therefore, generally employs the same bed outlet temperature for all the beds.Too low an outlet feed temperature is undesirable because the low reaction rates at the beginning of the feed entry bed cause excessive catalyst requirements for a given duty. It may be seen that the hydrogen quench injected interstage in the process of the present invention is always lower than in a downflow static bed process to remove the same reaction heat. This is because of the additional heat required in a progressive temperature pattern due to the heat requirements of the previous stage product in being heated from the preceding bed outlet temperature level to the higher bed outlet temperature of the stage concerned.

Figures 1 and 2 show two different arrangements of apparatus for use with the process of the present invention in the two illustrative examples discussed hereinbelow.

Two embodiments of the process of the present invention, as applied to distillate hydrogenation, are illustrated in the accompanying drawings, in which: Figure 1 shows an arrangement in which the stages of the process are contained in individual reactors; and Figure 2 shows an arrangement in which the stages of the process are stacked in one reactor.

Boxed areas for hydrogen recycle cleanup and common equipment such as feed pumps, hydraulic letdowns and separation stages have not been shown in the drawings for the sake of clarity.

In the arrangement of Figure 1, the hydrocarbon feed and the hydrogen makeup feed are introduced into the system by way of lines 20 and 21, respectively. These two fields are then combined and are introduced into the feed entry stage reactor 8 either by way of a line 19, which is a bypass normally controlled by the temperature recorded controller TRC of the feed entry stage reactor 8, or else by way of a reactor effluent heat-exchanger 18, and, if necessary, via a start-up heater 17. Because of the limited hydrogen consumption in feed entry, hydrogen makeup results in a satisfactory gas phase in the outlet stream of the reactor 8 and in adequate saturation by, and availability of, hydrogen.This can be accomplished, for example, by limiting the chemical hydrogen consumption in the stage 8 to approximately 50 So of the overall chemical hydrogen consumption. The reactor 8 also has catalyst withdrawal means 28 and catalyst addition means 22. The effluent from the reactor 8 is fed by way of a line 33 to the bottom of the second stage reactor 9. The reactor 9 also has catalyst withdrawal means 29 and catalyst addition means 23. Effluent from the second stage reactor 9 is fed by way of a line 34 to the bottom of a reactor 35, which comprises both a third stage 10 and a fourth stage 11. The reactor 35 has catalyst withdrawal means 31 and catalyst addition means 24 for the third stage, and catalyst withdrawal means 30 and catalyst addition means 25 for the fourth stage. Feed passes from the bottom of the reactor 35 to its top via the third stage 10 and the fourth stage 11, and the effluent from the reactor 35 is fed by way of a line 27 to a feed inlet heat exchanger 12. The hydrogenated material is then passed by way of an air cooler line 37 to a thermal water cooler CW and a high pressure separator 36. Facilities are provided for the introduction of condensate wash via a line 32 entering the blow between the heat exchanger 12 and the air cooler. Ammoniacal water and hydrogenated liquid phase material are passed to separation via lines 13 and 14, respectively. Hydrogen which is to be recycled is passed from the high pressure separator 36 by way of a line 38 to a hydrogen sulphide removal stage (amine absorber) 15 and, thence, to a methane plus adsorption stage 16.The hydrogen which has been cleaned up in this way may then be recycled interstage via a line 26.

In the arrangement of Figure 2, the inlet feed, having been subjected to preliminary heating at 18 and controlled by the line 19 to the necessary feed inlet temperature, corresponding to the decided feed entry stage temperature as described above, is introduced into the feed entry stage 7 of a stacked reactor by way of a line 1. The material to be hydrogenated then passes up the stacked reactor via the second stage 6, the third stage 5 and finally the fourth stage 4. Effluent from the fourth stage is passed via a line 3 to heat-exchange and product separation, as described above. Cleaned up hydrogen for recirculation may be introduced into the stacked reactor via a line 2, after having been treated as described above. Condensate wash is usually periodically injected (not shown) to remove any ammonia hydrosulphide solid deposits which may sublime in the cooler portion.

Other hydrogenation applications may require a different number of stages and, for certain hydrogenations, e.g. those with lower chemical hydrogen consumption, it may be desirable to introduce a limited amount of hydrogen recycle to the feed inlet. One may also remove part of the reaction heat by means other than the introduction of cold hydrogen quench such as introducing part of the cold liquid feed as a liquid quench medium for the second stage, or removal by direct heat-exchange, e.g., by regenerating high-pressure steam. For high heteroatom content feed-stocks, such as gas oil in the Example, hydrogen quench is preferred kinetically as such because injection reduces the inhibition of ammonia, hydrogen sulfide and water by-products. This is because such inhibition is a function of the gas phase concentrations of these components under reactor conditions.

Whether the stages are contained in individual reactors or are stacked as shown in Figure 2 depends upon the individual situation, reactor availability and also economics. Because the gas phase flow significantly increase in the process of the present invention in the first several successive stages, and also because it is less expensive to provide a given reactor volume with a smaller diameter and compensating with increased length, the reactor diameters are economically preferred to decrease at the feed entry as permitted by the low gas phase volume. With high gas phase volumes, it is also desirable to limit the superficial gas phase velocity of the present invention to below 6 centimeters per second, in order to prevent any fluid-catalyst particle disengaging problems at the top of a bed.The process of the present invention, has because of the low hydrogen circulation rates, a high mass throughput reactor cross-section, and even the final reactor diameter is generally smaller than that used in a downflow fixed bed. This favours the process of the present invention as compared with downflow fixed bed reactors with respect to reactor capital investment because the smaller reactor diameter more than offsets any bed expansion volume under preferred design conditions.

Limited backmixing of the liquid phase occurs within a stage due to particulate fluidization but, with the elimination of product liquid as a result of recirculation to the inlet of the reactor bed, e.g., as in H-oil, the length to diameter ratios of the beds, and also with a limited number of stages, the dispersion or deviation from plug flow behaviour are similar to those of a well-designed trickle flow, downflow fixed bed reactor. Problems associated with downflow fixed beds such as maldistribution or hot spots simply cannot occur with the process of the present invention in the recommended range of operation. Discrete catalyst particle sizes of 0.4 to 1 millimeter in diameter are preferably employed for the process of the present invention.There is nothing unique about the catalyst composition, i.e., any catalyst which performs well in a fixed bed hydrotreater can be expected to perform well in the process of the present invention. Although a smaller catalyst particle size enhances the aromatic saturation rates, particularly those of polyaromatics, and this benefits the process of the present invention, all kinetics in the following two illustrative Examples have been based upon those obtainable with a 1/16 inch (1.6 millimeters) diameter nickelmolybdenum on alumina catalyst which is commercially available and is used to determine the kinetics of the downflow fixed bed reactors. This should demonstrate that the considerable capital investment and utility advantages which the present invention possesses over existing reactor systems is independent of any particle size effects.

The two examples in which the process of the present invention is compared with conventional existing processing demonstrate that, with the same reactor catalyst inventory and hydrogen partial pressure, substantial economic advantages occur. The two examples both represent distillate feedstocks derived from Kuwait crude having two widely different boiling ranges.

The first example consists in the hydrogenation of a Kerosene feedstock having a normal boiling range of from 180"C. (356"F) to 2650C. (509"F), an aromatic content of 25 percent by volume and a sulfur content of 0.51 weight percent. This feedstock is hydrogenated in a four-bed hydrotreater using a hydrogen makeup with 0.975 mol fraction of hydrogen and 0.025 mol fraction of methane, to produce a 140 C (284 F) plus product with a 2 percent aromatic content. This may be accomplished with a 1.0 weight hourly space velocity and a hydrogen partial pressure of 1,600 psia for the reactor outlet. Chemical hydrogen consumption is the equivalent of 600 standard cubic feet per barrel.Characteristics of the 140"C plus product are 38 millimeter smokepoint, a freezing point which complies with commercial avtur specifications, and 15 parts per million (ppm) by weight of sulfur. The start of run conditions for a commercial downflow fixed bed hydrotreater are temperatures at 338"C (640"F) for the feed inlet and 371"C (700"F) for the bed outlet.

Since, for the purpose of kinetic calculations, the same catalyst is employed in the process of the present invention to produce the same aromatic content product at the same reactor outlet hydrogen partial pressure, and also the maximum temperature never exceeds the outlet temperature used in the downflow reactor system at the start of the run, the chemical hydrogen consumption and product characteristics are independent of the reactor system.

Three simplifications are made to facilitate the comparison of the processes in the two examples discussed herein.

The first simplification is that all hydrogen quench requirements are calculated using 0.99 mol fraction hydrogen and 0.01 mol fraction methane for the recycle gas. This provides a common basis for both quench requirements and also for the kinetics. It furthermore results in approximately the same design pressure for the reactors, high pressure heat-exchangers, fired trim heater for the inlet fee , reactor product coolers, recycle compressor and after compressor cooler, as well at the same hydrogen makeup compression requirements because of nearly identical hydrogen solution losses in the liquid phase product stream.

In the process of the present invention, because of the comparatively low volume of the hydrogen recycle and the comparatively small proportionate amount of methane plus components in the liquid phase (due to the favourable 1/v ratio at the high pressure separator), cleanup of the hydrogen recycle is economically justifiable.This is due to the fact that the total investment for the high pressure amine absorber (amine regeneration is unaffected as this is related to the total hydrogen sulfide in the reactor product), and for the high pressure adsorbers to remove methane, plus from recycle, including regeneration equipment, is lower than the incremental investment cost for the high pressure reactor section equipment which accommodates the high pressure associated with the methane plus components in the liquid phase product with equilibrium hydrogen recycle circulated for the same hydrogen partial pressure at the reactor outlet.

The high pressure separator and recycle gas compressor aftercooler outlet are assumed to be at 40"C to provide a common basis for cooling requirements and for flash calculations.

The high pressure amine absorber and absorber equipment to clean up the hydrogen recycle are excluded from the comparison in order to preclude any misinterpretation that the cleanup of the hydrogen recycle is biased towards the process of the present invention, as the order of hydrogen recycle to be treated would be at least three times that of the process of the present invention. Therefore, the higher reactor pressure of the equilibrium recycle system would probably be more economic for conventional downflow reactor systems, particularly as increased regeneration volumes would be required for the high pressure absorbers. Removal of hydrogen sulfide should be practised for higher sulfor content feedstocks, since higher concentrations of hydrogen sulfide component in the reaction zone inhibit the aromatic saturation reaction.

The second simplification to be made is that longer term conditions (approximately two year's onstream time) for the process of the present invention are compared with start of the run conditions for the downflow fixed bed reactor system. This simplification results in the same reactor outlet temperature for the two systems being compared, and provides a common basis for the quench quality calculated. In practice, the conventional fixed bed reactor system gradually increases the temperature othe bed outlets to compensate for a gradual decline in activity. For aromatic saturation, a low aromatic content of the product limits the possible higher temperature because thermodynamic equilibrium becomes controlling with respect to the maximum temperature limit. Normal average run conditions for the downflow fixed bed are, typically, 10 to 15"C. higher than start of the run conditions for both the feed entering the reactor system and for the reactor product leaving the reactor system. Although the higher outlet temperature from the reactor entering the reactor effluent-reactor feed exchangers recovers part of the heat in the form of a higher feed inlet temperature to the fired trim heater, the net overall effect, because of the finite heat-exchange surface provided, is an increase in the fired heater duty and trim cooling duty in cooling the reactor product. Reactor quench requirements are not affected since equipment must still be capable of furnishing requirements for the limiting conditions which is the start of run (SOR) condition.The decreased quench requirement as a result of increased heat in a given quantity being heated to a higher outlet temperature therefore does not effect the design requirements of the fixed bed reactor system. From a temperature-material viewpoint a downflow fixed bed reactor system must have equipment designed for the end of run conditions, and must also provide sufficient equipment, such as the reactor feed trim fired heater, the reactor feed-reactor effluent air cooler and the reactor effluent trim cooler for this end of run condition, as this is limiting from a design viewpoint.

For the same reactor pressure and reactor catalyst inventory, the process of the present invention utilizes the inherently better average bed temperature characteristics to reduce utilities. Normally constant bed outlet temperatures are maintained for the stages other than the feed entry stage. The feed entry stage is gradually increased to compensate for catalyst deactivation with the following estimated temperatures for two years onstream of the kerosene hydrotreater and the gas oil hydrotreater (the second example discussed below):: Kerosene Hydrotreater Gas oil Hydrotreater Volume percent Volume percent Controlled aromatics in 140"C Controlled aromatics in outlet, plus product at outlet 1600C plus "c, ("F) outlet "e ("F) product at outlet Feed inlet 355 (671) 14 359 (678) 27 stage Second stage 364 (687) 7 369 (696) 15 Third stage 369 (696) 4 375 (707) 9 Fourth stage 371 (700) 2 378 (712) 6 A lower maximum temperature where the more reactive components are present in the reaction zone reduces the carbonaceous deposition rate as well as allowing a more selective hydrogenation pattern.A lower temperature at the feed inlet also increases the hydrogen partial pressure for a given reactor pressure, which further moderates the carbonaceous deposition and thereby enhances aromatic saturation rates.

Operating with the fourth stage continuously at start of run temperature from the initial onstream condition, in the process of the present invention, achieves a two-fold purpose.

The additional preheat ensures the operability of the heat-exchange system to provide all the necessary heat without the start-up heater being required for anything other than the increasing the reactor feed inlet to a sufficient kindling temperature, and this means that the reactor system is self-sustaining in generating the necessary heat of reaction for control. If hydrogen availability means that the chemical hydrogen should not exceed the design limits, the fourth stage catalyst inventory will initially be less than normal when commencing entirely with original fresh catalyst. This enables regulation of the atomatic saturation by control of the feed entry stage only, with good bypass control. An initial feed entry stage temperature of 350"C is visualized.Fresh sulfided catalyst is then added to the fourth stage in batches as required to maintain the product quality. When full design reactor inventory of the fourth stage has been obtained, the feed entry stage temperature is increased gradually as required to maintain quality. Having the fourth stage operate continuously at the 371"C temperature provides, and assists in controlling, the necessary chemical hydrogen consumption until the higher than normal fresh sulfided catalyst activity of the initial inventory has been dissipated.This is because the controlled reactor outlet temperature provides a thermodynamic limitation to the amount of aromatic saturation which is possible even with an hyperactive fresh sulfided catalyst, e.g., the lower thermodynamic limit is in the region of 11H2 volume percent aromatics for the kerosene hydrotreater, and in the region of 4 volume percent aromatics for the gas oil hydrotreater.

The ability in the process of the present invention to add catalyst and to withdraw catalyst onstream while continuing to make specification product, provides an insurance feature for obtaining the desired end result which a fixed bed reactor does not posess. For instance, let us assume that a different crude must be processed by the refinery which results in a feedstock of the same boiling range but with a considerably higher catalyst activity decline.

When the product specification cannot be met, one may reduce the feed rate to approximately 70 percent of the design rate, withdraw the entire catalyst batch from the fourth stage and replace it with fresh sulfided catalyst. Full design feed rate may then be resumed with specification product made during the replacement. The withdrawn catalyst batch may then be regenerated to restore fresh sulfided activity by oxidation of the carbonaceous deposits. Such catalyst is then presulfided using part of the letdown gas stream containing hydrogen sulfide, and is utilized to replace the catalyst withdrawn when similar withdrawals are required to maintain catalyst activity.Apart from mishaps, such as the introduction of contaminated feed, one would expect that onstream regenerations would be confined to the third or fourth stages, as the activity of these two stages is controlling for an aromatic specification. In the case of a contaminated feed being accidentally introduced and seriously impairing the activity in the feed entry stage, one may also recover, e.g., by operating at a reduced rate (instead of being forced to shutdown as in the case of a downflow fixed bed reactor system), and replacing the contaminated feed entry bed with sulfided non-contaminated catalyst.

The third simplification in the comparison is that, arbitrarily, the reactor effluent is considered to be exchanged only with the reactor feed inlet. The hydrocarbon feed is assumed to enter the reactor system at 600C. Air cooling is assumed to cool the reactor effluent leaving the reator feed-reactor effluent exchanger to 600C. The reactor effluent then enters a water cooler which cools the reactor effluent to 40"C before it enters the high pressure separator. This simplification favours the conventional process since a considerable portion of the heat available in the reactor product stream and which is removed by the reactor product aircooler, may be utilized economically in separating the reactor product.

The total of the reactor feed-reactor effluent exchanger surface and the recovered heat transfer surface also using the reactor effluent for the process of the present invention would not be greater than 70 percent of the conventional reactor feed effluent exchanger surface.

Heat of vaporization and condensation have been neglected in the calculated heat duties.

This again favours the conventional process as over four times the weight of original feed material is vaporized in the conventional process as compared with the process of the present invention, at the same reactor outlet pressure and temperature. Furthermore, in the conventional process, the greater portion of the vaporization heat input occurs between the feed trim heater and the outlet of the feed entry stage, thereby significantly increasing, in practice, the fired heater inlet trim duty. On the other hand, most of the vaporization in the process of the present invention occurs from the outlet of the feed entry stage to the final reactor outlet because of the progressive stage temperature increase, and also the considerable change in liquid to vapour ratio from that of the feed entry stage outlet to that of the final reactor outlet.Vaporization in this section of the reactor reduces the quench requirements by the heat equivalent. Quench requirements for the process of the present invention may therefore, be expected to be significantly lowered in practice, whereas the conventional process undergoes only a slight reduction. This is because the feed entry stage outlet temperature in the conventional process is already at reactor outlet temperature with an appreciable gas phase present, and, accordingly, the liquid to vapour ratio change between the feed entry and the final reactor outlet is only one-fourth that of the process of the present invention.

A high hydrogen recycle in contrast to a low hydrogen recycle is disadvantageous in heat-exchange since any heat vaporization is recovered as heat of condensation at a considerably lower temperature. A significant increase may therefore, be expected in practice for the aircooling and water cooling duties in the conventional process as opposed with a limited increase for the process of the present invention, which process, of course, uses a low hydrogen recycle.

If both reactor systems operated with the entire methane and higher boiling products, including hydrogen sulfide dissolved in the hydrocarbon liquid phase product, the comparison is even more favorable to the process of the present invention, because of the considerably higher liquid to vapour ratio prevailing in the high-pressure separator. A higher pressure would be required for the conventional process as compared with the process of the present invention to maintain the same hydrogen partial pressure when operating at the same reactor outlet temperature because of the additional methane introduced into the reactor system with the conventional process as compared with the process of the present invention, under the equilibrium solution conditions, is particularly detrimental to the reactor hydrogen partial pressures. The presence of additional methane also unfavourably affects the vapour-liquid equilibrium ratios, for a given set of process conditions, in that the volatility of the higher boiling components is increased. The methane which would be introduced into the reactor system if the hydrogen recycle were not cleaned up in the conventional process would be of the order of 4.2 times that of the process of the invention, e.g., for the gas oil hydrotreater, the mol fraction hydrogen for the conventional process hydrogen recycle would be 0.826 as compared with 0.832 for the process of the invention, whereas the methane concentration would be 0.132 mol fraction for the conventional process versus 0.128 for the process of the present invention. Thus, all the simplifications made favour the conventional process.

Unit requirements per barrel Process of Conventional of feedstock for Kerosene Invention Reactor Hydrotreater Comparison Process Total hydrogen recycle to 1,340 5,190 be compressed, SCF Hydrogen recycle compressor, .011 .043 BHP Hydrogen recycle compressor 270 1,100 aftercooler duty, Btu Hydrogen makeup to feed inlet, 675 675 SCF Hydrogen recycle to feed inlet, None 2,825 SCF Hydrogen recycle to interstage 1,314 2,330 quench, SCF Reactor feed-reactor effluent 80,480 116,640 exchanger duty, Btu Reactor feed-reactor effluent 95 66 exchanger MTD, "C Reactor feed fired heater duty, Startup 5,980 Btu only Reactor effluent temperature 179 151 entering aircooler, "C Reactor effluent aircooler duty, 42,540 43,260 Btu Reactor effluent trim water cooling 6,490 9,080 duty, Btu It can be calculated that the reactor feed-reactor effluent exchanger surface area for the process of the present invention is only 0.57 that of the conventional process, even after allowing for an additional 20 percent to ensure operating control without a fired heater after start-up. From a capital investment viewpoint, this item represents an order of magnitude similar to the reactor. High pressure interconnecting piping between the high pressure equipment also constitutes a costly item for high-pressure hydrogenation installations.

Because the gas circulation is the largest single factor in the sizing of these lines and the hydrogen recycle for the process of the invention is only 0.258 of that of the conventional process, the interconnecting piping line sizes may be expected to be significantly reduced when using the process of the invention.

The second example is a straight-run gas oil with a normal boiling range of 345"C (653"F) to 440C (884"F) having an aromatic content of 40 percent by volume, a sulfur content of 2.65 weight percent and a nitrogen content of 590 ppm. This feedstock is hydrogenated in a four-bed hydrotreater to 6 weight percent aromatics in the 1600C (320"F) plus product with a similar hydrogen supply and recycle to that described above. This may be accomplished with an 1.0 weight hourly space velocity and an hydrogen partial pressure of 1,800 psia for the reactor outlet. Chemical hydrogen consumption is the equivalent of 1,020 standard cubic feet per barrel. Product characteristics are a 75 diesel index, a sulfur of 23 ppm, and a nitrogen content of 2 ppm.The associated 375"C (707"F) plus product has a viscosity of 110 after dewaxing to OOF. pour point. Start of the run condition for the downflow fixed bed hydrotreater are a reactor feed inlet of 340"C. (644"F) with 3780C. (712"F) bed outlet temperatures.

Unit requirements per barrel Process of Conventional of feedstock for Gas Oil invention Reactor Hydrotreater Comparison Process Total hydrogen recycle to 2,385 9,170 be compressed, SCF Hydrogen recycle compressor, .019 .077 BHP Hydrogen recycle compressor, 490 1,940 aftercooler duty, Btu Hydrogen makeup to feed inlet, 1,038 1,038 SCF Hydrogen recycle to feed None 4,949 inlet, SCF Hydrogen recycle to interstage 2,307 3,381 quench, SCF Reactor feed-reactor effluent 75,470 144,670 exchanger duty, Btu Reactor feed-reactor effluent 132 77 exchanger MTD, "C Reactor feed hired heater Startup 7,430 duty, Btu only Reactor effluent temperature 226 176 entering aircooler, Reactor effluent aircooler 71,160 74,400 duty, Btu Reactor effluent trim water 7,520 12,040 cooling duty, Btu Hydrogen recycle to be compressed is 0.26 of the conventional process. Reactor feed-reactor effluent exchanger surface, including control (20 percent additional heat removal), is only 0.29 that of a conventional process. From the two tables, it may be seen that the process of the present invention has substantial capital investment and utility advantages over conventional processing.

The process of the present invention should have a particular advantage when synthetic crude materials become available in the future and are to be upgraded into aviation turbine and automotive diesel fuels which comply with existing specifications. Coal liquids, for example, typically contain 75 volume percent aromatics but, based upon the foregoing, it should be possible to make an acceptable cetane number in a more economical manner using the process of the present invention.

It may also be economical to consider upgrading gas oils for olefin manufacture by pyrolysis because the olefin yield is proportional to the saturates in the feedstock and the high boiling tar by-product is a function of the aromatics in the feedstock. The gas oil example indicates that the sulfur, nitrogen, and aromatic contents of the hydrogenated product can be lower than the typical naphtha feedstocks.

One final advantage that the process of the present invention provides over conventional processes is that the user has flexibility in case excess hydrogen is available and lighter distillates are required. Part of the catalyst in the fourth stage may be withdrawn and replaced with a hydrocracking catalyst. For example, 18 volume percent of gas oil consists of tri- and higher aromatics. These cyclics may be selectively converted into lower boiling distillates with a limited amount of hydrocracking catalyst because these components are preferentially hydrocracked first. The quantity of hydrocarbon catalyst in the 4th stage may be varied to match either the desired conversion or the H2 availability. If conditions change, one may withdraw the fourth stage and replace with normal catalyst so that the original design objectives can be resumed.These changes in yield pattern may be made onstream by temporarily reducing the feed rate to 70 percent while withdrawing the fourth stage catalyst.

WHAT I CLAIM IS: 1. A process for the hydrogenation of distillates which comprises the steps of: i) passing the hydrocarbon feed and hydrogen through a plurality of fluidised catalyst bed reactor stages operated in a progressive temperature pattern with the lowest bed outlet temperature at the feed inlet and the highest bed outlet temperature at the final reactor outlet, said plurality of fluidised catalyst bed reactor stages having a liquid phase mentioned at a near hydrogen-saturated state, and the temperature of the inlet feed in the first reactor stage being regulated to control the outlet stream of said first reactor stage; and ii) passing the products from the outlet of the final reactor stage to a separator stage.

2. A process as claimed in claim 1, wherein the hydrogen makeup constitutes the entire inlet hydrogen requirements and the temperature of the inlet feed is regulated by bypassing part of the feed around a feed inlet-reactor product heat exchanger.

3. A process as claimed in claim 1 or claim 2, wherein part of the reaction heat is removed by introducing cold liquid feed as a liquid quench medium in the second stage of the plurality of reactor stages.

4. A process as claimed in claim 1 or claim 2, wherein part of the reaction heat is removed by direct heat exchange.

5. A process as claimed in claim 4, wherein the direct heat exchange is performed by generating high pressure steam.

6. A process as claimed in any of claims 1 to 5 wherein hydrogen recycle is injected interstage to remove the heat of reaction and to maintain the bed outlets of the stages following the feed bed entry stage.

7. A process as claimed in claim 2, wherein a fired heater is used to heat up the reactor feed, until the reactor outlet temperature becomes sufficient to furnish the necessary heating of the feed stream by heat exchange, and is then by-passed and shutdown.

8. A process as claimed in any of claims 1 to 7, wherein the superficial gas phase velocity is 6cm/sec or below.

9. A process as claimed in any of claims 1 to 8, wherein the plurality of fluidised catalyst bed reactor stages are all sequentially stacked in a single reactor.

10. A process as claimed in any of claims 1 to 8, wherein each stage of the plurality of fluidised catalyst bed reactor stages is contained in an individual reactor.

11. A process as claimed in any of claims 1 to 10, wherein the catalyst inventory is varied in accordance with temperature and with hydrogen availability and to maintain catalyst activity based on a desired product specification, part of the feed being selectively conducted to a lower boiling range, when needed by varying the catalyst inventory in the final stages.

12. A process as claimed in any of claims 1 to 11, wherein, in the event of combination of the feed, feed entry stage catalyst is partly or wholly replaced with fresh sulphide catalyst while continuing to operate onstream with at least 70 percent feed capacity.

13. A process as claimed in any claims 1 to 12 substantially as herein described with reference to the accompanying drawings.

**WARNING** end of DESC field may overlap start of CLMS **.

Claims (13)

**WARNING** start of CLMS field may overlap end of DESC **. replaced with a hydrocracking catalyst. For example, 18 volume percent of gas oil consists of tri- and higher aromatics. These cyclics may be selectively converted into lower boiling distillates with a limited amount of hydrocracking catalyst because these components are preferentially hydrocracked first. The quantity of hydrocarbon catalyst in the 4th stage may be varied to match either the desired conversion or the H2 availability. If conditions change, one may withdraw the fourth stage and replace with normal catalyst so that the original design objectives can be resumed. These changes in yield pattern may be made onstream by temporarily reducing the feed rate to 70 percent while withdrawing the fourth stage catalyst. WHAT I CLAIM IS:

1. A process for the hydrogenation of distillates which comprises the steps of: i) passing the hydrocarbon feed and hydrogen through a plurality of fluidised catalyst bed reactor stages operated in a progressive temperature pattern with the lowest bed outlet temperature at the feed inlet and the highest bed outlet temperature at the final reactor outlet, said plurality of fluidised catalyst bed reactor stages having a liquid phase mentioned at a near hydrogen-saturated state, and the temperature of the inlet feed in the first reactor stage being regulated to control the outlet stream of said first reactor stage; and ii) passing the products from the outlet of the final reactor stage to a separator stage.

2. A process as claimed in claim 1, wherein the hydrogen makeup constitutes the entire inlet hydrogen requirements and the temperature of the inlet feed is regulated by bypassing part of the feed around a feed inlet-reactor product heat exchanger.

3. A process as claimed in claim 1 or claim 2, wherein part of the reaction heat is removed by introducing cold liquid feed as a liquid quench medium in the second stage of the plurality of reactor stages.

4. A process as claimed in claim 1 or claim 2, wherein part of the reaction heat is removed by direct heat exchange.

5. A process as claimed in claim 4, wherein the direct heat exchange is performed by generating high pressure steam.

6. A process as claimed in any of claims 1 to 5 wherein hydrogen recycle is injected interstage to remove the heat of reaction and to maintain the bed outlets of the stages following the feed bed entry stage.

7. A process as claimed in claim 2, wherein a fired heater is used to heat up the reactor feed, until the reactor outlet temperature becomes sufficient to furnish the necessary heating of the feed stream by heat exchange, and is then by-passed and shutdown.

8. A process as claimed in any of claims 1 to 7, wherein the superficial gas phase velocity is 6cm/sec or below.

9. A process as claimed in any of claims 1 to 8, wherein the plurality of fluidised catalyst bed reactor stages are all sequentially stacked in a single reactor.

10. A process as claimed in any of claims 1 to 8, wherein each stage of the plurality of fluidised catalyst bed reactor stages is contained in an individual reactor.

11. A process as claimed in any of claims 1 to 10, wherein the catalyst inventory is varied in accordance with temperature and with hydrogen availability and to maintain catalyst activity based on a desired product specification, part of the feed being selectively conducted to a lower boiling range, when needed by varying the catalyst inventory in the final stages.

12. A process as claimed in any of claims 1 to 11, wherein, in the event of combination of the feed, feed entry stage catalyst is partly or wholly replaced with fresh sulphide catalyst while continuing to operate onstream with at least 70 percent feed capacity.

13. A process as claimed in any claims 1 to 12 substantially as herein described with reference to the accompanying drawings.