Patent Publication Number: US-2002011428-A1

Title: Multistage moving-bed hydroprocessing reactor with separate catalyst addition and withdrawal systems for each stage, and method for hydroprocessing a hydrocarbon feed stream

Description:
[0001] This application is a continuation-in-part of co-pending application entitled CATALYST, METHOD AND APPARATUS FOR AN ON-STREAM PARTICLE REPLACEMENT SYSTEM FOR COUNTERCURRENT CONTACT OF A GAS AND LIQUID FEED STREAM WITH A PACKED BED, Ser. No. 08/497,638, filed Jun. 30, 1995; which application is a continuation of application Ser. No. 08/235,043, filed Apr. 29, 1994; which application is a continuation in part of Ser. No. 08/215,254, filed Mar. 21, 1994, now U.S. Pat. No. 5,409,598, dated Apr. 25, 1995; which application is a divisional of Ser. No. 08/014,847, filed Feb. 8, 1993. now U.S. Pat. No. 5,409,598, dated Apr. 12, 1994; which application is a continuation application of Ser. No. 07/727,656, filed Jul. 9, 1991; which application is a divisional application of application Ser. No. 07/381,948, filed Jul. 19, 1989, now U.S. Pat. No. 5,076,908 dated Dec. 31, 1991. Benefit of the earliest filing date is claimed, especially with respect to all common subject matter. 
    
    
     
       BACKGROUND OF THE INVENTION  
       [0002] 1. Field of the Invention  
       [0003] The present invention broadly relates to a multistage moving bed hydroprocessing reactor with separate catalyst addition and withdrawal systems for each stage.  
       [0004] More particularly, the present invention relates to, a method of, and apparatus for hydroprocessing where two or more different (i.e. different in physical and/or catalytic properties) and distinct moving bed catalysts are employed in a single onstream reactor for any hydroprocessing application, and where catalyst particles are added and withdrawn from each of the two or more different and distinct moving bed catalyst while the single onstream reactor is hydroprocessing a hydrocarbon feed stream. The present invention provides a method and a reactor which combines the advantages of a moving-bed reactor with the advantages of layering different catalyst within a single reactor, preferably and optionally while maintaining continuous or intermittent replacement of catalyst for plug-like flow of each of the catalyst bed through the reactor. Such plug flow with high processing rates is obtained by selecting the size, shape and density of the catalyst particles to prevent ebullation at the design flow rate so as to maximize the amount of catalyst in the vessel during normal operation and during catalyst transfer. Catalysts are selected by measuring bed expansion with hydrocarbon, hydrogen and catalyst at the design pressures and flow velocities within the available reaction volume of the reactor.  
       [0005] 2. Description of the Prior Art  
       [0006] Hydroprocessing or hydrotreatment to remove undesirable components from hydrocarbon feed streams is a well known method of catalytically treating such heavy hydrocarbons to increase their commercial value. “Heavy” hydrocarbon liquid streams, and particularly reduced crude oils, petroleum residua, tar sand bitumen, shale oil or liquified coal or reclaimed oil, generally contain product contaminants, such as sulfur, and/or nitrogen, metals and organo-metallic compounds which tend to deactivate catalyst particles during contact by the feed stream and hydrogen under hydroprocessing conditions. Such hydroprocessing conditions are normally in the range of 212 degree(s) F to 1200 degree(s) F (100 degree(s) to 650 degree(s) C.) at pressures of from 20 to 300 atmospheres. Generally such hydroprocessing is in the presence of catalyst containing group VI or VIII metals such as platinum, molybdenum, tungsten, nickel, cobalt, etc., in combination with various other metallic element particles of alumina, silica, magnesia and so forth having a high surface to volume ratio. More specifically, catalyst utilized for hydrodemetallation, hydrodesulfurization, hydrodenitrification, hydrocracking etc., of heavy oils and the like are generally made up of a carrier or base material; such as alumina, silica, silica-alumina, or possibly, crystalline aluminosilicate, with one more promoters or catalytically active metal(s) (or compound(s)) plus trace materials. Typical catalytically active metals utilized are cobalt, molybdenum, nickel and tungsten; however, other metals or compounds could be selected dependent on the application.  
       [0007] Because these reactions must be carried out by contact of a hydrogen-containing gas with the hydrocarbon feed stream at elevated temperatures and pressures, the major costs of such processing are essentially investment in vessels and associated furnaces, heat exchangers, pumps, piping and valves capable of such service and the replacement cost of catalyst contaminated in such service. Commercial hydroprocessing of relatively low cost feed stocks such as reduced crude oils containing pollutant compounds, requires a flow rate on the order of a few thousand up to one hundred thousand barrels per day, with concurrent flow of hydrogen at up to 10,000 standard cubic feet per barrel of the liquid feed. Vessels capable of containing such a reaction process are accordingly cost-intensive both due to the need to contain and withstand corrosion and metal embrittlement by the hydrogen and sulfur compounds, while carrying out the desired reactions, such as demetallation, denitrification, desulfurization, and cracking at elevated pressure and temperatures. Pumps, piping and valves for handling fluid streams containing hydrogen at such pressures and temperatures are also costly, because at such pressures seals must remain hydrogen impervious over extended service periods of many months and years. It is also cost-intensive to insure that all additional reactor vessels (e.g. fixed bed reactors, etc.) are obtained in order to catalytically process hydrocarbon feed streams from an initial reactor vessel where certain catalytic hydroprocessing (e.g. hydrodemetallation) is/are to be performed to one or more other reactor vessel(s) where additional catalytic hydroprocessing (e.g. hydrodenitrification) is/are to be performed.  
       [0008] Further, hydroprocessing catalyst for such one or more reactor vessel(s), which typically contains catalytically active metals such as titanium, cobalt, nickel, tungsten, molybdenum, etc., may involve a catalyst inventory of 500,000 pounds or more at a cost of $2 to $4/lb. Accordingly, for economic feasibility in commercial operations, the process must handle high flow rates and the one or more reactor vessel(s) should be filled with as much catalyst inventory as possible to maximize catalyst activity and run length. Additionally, the down-time for replacement or renewal of catalyst must be as short as possible. Further, the economics of the process will generally depend upon the versatility of the system to handle feed streams of varying amounts of contaminants such as sulfur, nitrogen, metals and/or organic-metallic compounds, such as those found in a wide variety of the more plentiful (and hence cheaper) reduced crude oils, residua, or liquified coal, tar sand bitumen or shale oils, as well as used oils, and the like.  
       [0009] It is known to use a series of individual reactor vessels stacked one above the other, with fluid flow either co-current or counterflow to catalyst. In such a process, catalyst moves by gravity from the upper vessel to a lower vessel by periodically shutting off, or closing, valves between the individual vessels. In a counterflow system, this permits removal of catalyst from the lowermost or first stage vessel, where the most contaminated, or raw, feed stock, originally contacts the catalyst. In this way, most of the major contaminating components in the hydrocarbon stream are removed before the hydrocarbon material reaches major conversion steps of the process performed in higher vessels of the stacked series. Thus, most of the deactivating components of the feed stream are removed before it reaches the least contaminated catalyst added to the topmost vessel. However, such systems require valves suitable for closing off catalyst flow against catalyst trapped in the line. Hence, valve life is relatively short and down-time for replacement or repair of the valves is relatively costly. Also, such series of individual reactor vessels are costly since each respective reactor vessel must be purchased separately.  
       [0010] As particularly described in U.S. Pat. No. 5,076,908 to Stangeland et al, a substantially packed-bed type reactor system is an upflow type reactor system including multiple reaction zones of packed catalyst particles having little or no movement during normal operating conditions of no catalyst addition or withdrawal. In the substantially packed-bed type reactor system of Stangeland et al., when catalyst is withdrawn from the reactor during normal catalyst replacement, the catalyst flows in a downwardly direction under essentially plug flow or in an essentially plug flow fashion, with a minimum of mixing with catalyst in layers which are adjacent either above or below the catalyst layer under observation.  
       [0011] As particularly distinguished from prior known methods of on-stream catalyst replacement in hydroprocessing, the method and apparatus in U.S. Pat. No. 5,076,908 to Stangeland et al. more specifically provides a system wherein plug flow of the catalyst bed is maintained over a wide range of counterflow rates of a hydrocarbon feed stream and hydrogen gas throughout the volume of the substantially packed catalyst bed. Such packed catalyst bed flow maintains substantially maximum volume and density of catalyst within a given reactor vessel&#39;s design volume by controlling the size, shape and density of the catalyst so that the bed is not substantially expanded at the design rate of fluid flow therethrough. The proper size, shape and density are determined by applying coefficients gained during extensive studying of bed expansion with hydrocarbon, hydrogen and catalyst at the design pressures and flow velocities as particularly described below. However, Stangeland et al. does not teach or suggest a method and/or a reactor for hydroprocessing a hydrocarbon feed stream through a single onstream reactor having two or more different (i.e. different in physical and/or catalytic properties) and distinct moving bed catalysts, where catalyst particles are added and withdrawn from each of the two or more different and distinct moving bed catalysts while the single onstream reactor is hydroprocessing the hydrocarbon feed stream.  
       [0012] The prior art does not disclose or suggest two or more different and distinct moving bed catalysts in a single onstream reactor; nor does the prior art disclosed or suggest the above enumerated and pertinent features of either the total system or significant portions of such a system in U.S. Pat. No. 5,076,908 to Stangeland et al, as disclosed by the following patents:  
       [0013] Jacquin et al. U.S. Pat. No. 4,312,741, is directed toward a method of on-stream catalyst replacement in a hydroprocessing system by controlling the feed of hydrogen gas at one or more levels. Catalyst, as an ebullated bed counterflows through the reactor but is slowed at each of several levels by horizontally constricted areas which increase the hydrogen and hydrocarbon flow rates to sufficiently locally slow downward flow of catalyst. While local recycling thus occurs at each such stage, rapid through-flow of fresh catalyst, with resultant mixing with deactivated or contaminated catalyst, is suppressed. The ebullating bed aids simple gravity withdrawal of catalyst from the vessel. Improvement of the disclosed system over multiple vessels with valves between stages is suggested to avoid the risk of rapid wear and deterioration of valve seals by catalyst abrasion.  
       [0014] Kodera et al. U.S. Pat. No. 3,716,478, discloses low linear velocity of a mixed feed of liquid and H 2  gas to avoid expansion (or contraction) of catalyst bed. By low linear velocity of fluid upflow, gas bubbles are controlled by flow through the packed bed, but the bed is fluidized by forming the bottom with a small cross-sectional area adjacent the withdrawal tube. This assists discharge of catalyst without backmixing of contaminated catalyst with fresh catalyst at the top of the single vessel. The range of the bed level in the vessel is from 0.9 to 1.1 of the allowable bed volume (±10%) due to fluid flow through the bed. A particular limitation of the system is that flow of the fluids undergoing catalytic reaction is restricted to a rate that will not exceed such limits, but must be adequate to ebullate the bed adjacent the catalyst withdrawal tube. Alternatively, injection of auxiliary fluid from a slidable pipe section is required. The patentees particularly specify that the diameter of the lower end of the vessel is smaller to increase turbulence and ebullation of catalyst adjacent the inlet to the catalyst withdrawal line. Fluidization of catalyst is accordingly indicated to be essential to the process. However the disclosed gas flow rates are well below commercial flow rates and there is no suggestion of temperatures or pressures used in the tests or the size, density or shape of the catalyst.  
       [0015] Bischoff et al, U.S. Pat. No. 4,571,326, is directed to apparatus for withdrawing catalyst through the center of a catalyst bed counterflowing to a liquid hydrocarbon and gas feed stream. The system is particularly directed to arrangements for assuring uniform distribution of hydrogen gas with the liquid feed across the cross-sectional area of the bed. Such uniform distribution appears to be created because the bed is ebullating under the disclosed conditions of flow. Accordingly, considerable reactor space is used to initially mix the gas and hydrocarbon liquid feeds in the lower end of the vessel before flowing to other bottom feed distributors. The feeds are further mixed at a higher level by such distributor means in the form of “Sulzer Plates” or a “honeycomb” of hexagonal tubes beneath a truncated, conical, or pyramidal-shaped funnel screen. The arrangement may include an open ramp area parallel to the underside of the screen between the tube or plate ends. Further, to maintain gas distribution along the length of the catalyst bed, quench gas is supplied through upflowing jets in star-shaped or annular headers extending across middle portions of the vessel. The arrangement for withdrawal of spent catalyst requires ebullation of at least the lower portion of the bed. As noted above, added vessel space for uniform mixing of hydrogen and feed before introducing the fluids into an ebullated bed, as well as an ebullating bed, increases the required size of the hydroprocessing vessel, increases catalyst attrition, increases catalyst bed mixing and substantially increases initial, and continuing operating costs of the system.  
       [0016] Bischoff et al. U.S. Pat. No. 4,639,354, more fully describes a method of hydroprocessing, similar to U.S. Pat. No. 4,571,326, wherein similar apparatus obtains uniform ebullation through the vertical height of a catalyst bed, including a quench gas step.  
       [0017] Meaux U.S. Pat. No. 3,336,217, is particularly directed to a catalyst withdrawal method from an ebullating bed reactor. In the system, catalyst accumulating at the bottom of a vessel and supported on a flat bubble-tray may be withdrawn through an inverted J-tube having a particular ratio of the volume of the short leg of the J-tube to the longer leg. The diameter of the J-tube is suited only to flow of catalyst from a body of catalyst ebullated by the upflowing hydrocarbon feed and gas.  
       [0018] U.S. Pat. Nos. 4,444,653 and 4,392,943, both to Euzen, et al., disclose removal systems for catalyst replacement in an ebullating bed. In these patents, the fluid charge including hydrocarbon containing gas is introduced by various arrangements of downwardly directed jets acting laterally against or directly onto the conical upper surface of the bed support screen or screens. Alternatively, the feed is introduced through a conical screen after passing through a distributor arrangement of tortuous paths or a multiplicity of separate tubes to mix the gas and liquid feed over the conical screen. Such arrangements use a considerable volume of the pressure vessel to assure such mixing.  
       [0019] U.S. Pat. Nos. 3,730,880 and 3,880,569, both to Van der Toorn, et al., disclose a series of catalytic reactors wherein catalyst moves downwardly by gravity from vessel to vessel through check valves. As noted above, such valves require opening and closing to regulate the rate of flow, or to start and stop catalyst transfer, with catalyst in the valve flow path. Feed of process fluids is either co-current or countercurrent through the catalyst bed.  
       [0020] Van ZijllLanghaut et al. U.S. Pat. No. 4,259,294, is directed to a system for on-stream catalyst replacement by entrainment of the catalyst in oil pumped as a slurry either to withdraw catalyst from or to supply fresh catalyst to, a reactor vessel. Reacting feed is suggested to be either co-current or countercurrent with catalyst flow through the reactor. Valves capable of closing with catalyst in the line, or after back-flow of slurry oil, are required to seal off the catalyst containing vessel at operating temperatures and pressures from the receiving reactor vessel, or isolate the catalyst receiving lock hopper from the withdrawal section of the vessel.  
       [0021] Carson U.S. Pat. No. 3,470,900, and Sikama U.S. Pat. No. 4,167,474, respectively illustrate multiple single bed reactors and multi-bed reactors in which catalyst is replaced either continuously or periodically. The feed and catalyst flow co-currently and/or radially. Catalyst is regenerated and returned to the reactor, or disposed of. No catalyst withdrawal system is disclosed apart from either the configuration of the internal bed support or the shape of the vessel bottom to assist gravity discharge of catalyst.  
       [0022] One of the basic principles and teachings of Stangeland et al. in U.S. Pat. No. 5,076,908, is that by specifically selecting the size, shape, and density of the catalyst pellets, combined with appropriate control of process liquid and gas velocities, random motion and backmixing of the catalyst can be minimized, and plugflow characteristics of the catalyst downward and the liquid and gas flows upward, maximized. Stangeland et al. economically utilizes space within a hydroprocessing vessel over a wide range of processing rates without substantial random motion or ebullation of a packed bed of catalyst during high counterflow rates of the hydrocarbon feed and a hydrogen containing gas through the packed bed, while maintaining continuous or intermittent replacement of catalyst for plug-like flow of the bed through the vessel. Such plug flow with high processing rates is obtained by Stangeland et al. by selecting the size, shape and density of the catalyst particles to prevent ebullation and bed expansion at the design flow rate so as to maximize the amount of catalyst in the vessel during normal operation and during catalyst transfer. Catalysts are selected utilizing data gained while studying catalyst bed expansion, such as in a large pilot plant run, with hydrocarbon, hydrogen and catalyst at the design pressures and flow velocities within the available reaction volume of the vessel. Catalyst is removed from the bed by Stangeland et al. through laminar flow of the catalyst particles in a liquid slurry system in which the liquid flow line is uniform in diameter, and substantially larger than the catalyst particles, throughout the flow path between the reactor vessel and a pressurizable vessel including passageways through the flow control valves.  
       [0023] However, the method and reactor disclosed by Stangeland et al. in U.S. Pat. No. 5,076,908, as well as the method(s) and reactor(s) taught by the above-identified prior art patents relating to U.S. Pat. No. 5,076,908 to Stangeland et al, all suggest that additional reactor vessels are needed to further process hydrocarbon products produced by a reactor vessel containing a single catalyst bed. The reactor vessel(s) of the prior art allow the employment of one catalyst bed in a single reactor. If two or more catalyst beds are needed and they are not to be mixed together, then two or more separate moving-bed reactors placed in series are required. Separate reactors for separate purposes are expensive. Therefore, what is needed and what has been invented is a method and a reactor that is capable of containing two or more separate and distinct moving catalyst beds wherein each of separate and distinct moving catalyst has a different catalytic purpose (e.g. one moving catalyst bed would be for hydrodemetallation; a second moving catalyst bed would be for hydrodenitrification; a third moving catalyst bed would be for hydrodesulfurization; etc.).  
       SUMMARY OF THE INVENTION  
       [0024] The present invention accomplishes its desired objects by providing a reactor comprising a reactor vessel having an internal cylindrical wall with an internal reactor cylindrical wall diameter; and a first catalyst bed support means coupled to and supported by the internal cylindrical wall of the reactor vessel such as to essentially entirely span across the internal reactor cylindrical wall diameter for supporting a first catalyst moving bed in a first hydroconversion reaction zone having a first physical and/or catalytic properties and for allowing a hydrocarbon feed stream to pass therethrough while preventing catalyst particles from a first catalyst moving bed to exit therethrough. The reactor also comprises a first catalyst outlet means, communicating through the internal cylindrical wall of the reactor vessel with a first catalyst moving bed supported by the first catalyst bed support means, for withdrawing a first generally spent catalyst from a first hydroconversion reaction zone; and a first catalyst inlet means, communicating through the internal cylindrical wall of the reactor vessel with a first catalyst moving bed supported by the first catalyst bed support means, for introducing a first generally fresh catalyst into a first hydroconversion reaction zone. A second catalyst bed support means is coupled to and supported by the internal cylindrical wall of the reactor vessel such as to essentially entirely span across the internal reactor cylindrical wall diameter for supporting a second catalyst moving bed in a second hydroconversion reaction zone having a second physical and/or catalytic properties and for allowing a hydrocarbon feed stream to pass therethrough while preventing catalyst particles from a second catalyst moving bed to exit therethrough. The reactor further includes a second catalyst outlet means, communicating through the internal cylindrical wall of the reactor vessel with a second catalyst moving bed supported by the second catalyst bed support means, for withdrawing a second generally spent catalyst from a second hydroconversion reaction zone; and a second catalyst inlet means, communicating through the internal cylindrical wall of the reactor vessel with a second catalyst bed supported by the second catalyst bed support means, for introducing a second generally fresh catalyst into a second hydroconversion reaction zone.  
       [0025] The reactor additionally comprises a first catalyst moving bed supported by the first catalyst bed support means and having a first physical and/or catalytic properties; and a second catalyst moving bed supported by the second catalyst bed support means and having a second physical and/or catalytic properties. The first physical and/or catalytic properties differ from the second physical and/or catalytic properties. A first conduit means, communicating through the internal cylindrical wall of the reactor vessel is provided for introducing a hydrocarbon feed stream into the reactor vessel; and a second conduit means, also communicating through the internal cylindrical wall of the reactor vessel is provided for passing a hydrocarbon product out of the reactor vessel. The first catalyst bed support means allows an untreated hydrocarbon feed stream from the first conduit means to pass therethrough while preventing catalyst particles from the first catalyst moving bed to exit therethrough; and the second catalyst bed support means allows an initially treated hydrocarbon feed stream from the first catalyst moving bed to pass therethrough while preventing catalyst particles from the second catalyst moving bed to exit therethrough. Alternatively, the second catalyst bed support means receives an untreated hydrocarbon feed stream and allows an initially treated hydrocarbon feed stream from the second catalyst moving bed to pass therethrough while preventing catalyst particles from the second catalyst moving bed to exit therethrough; and the first catalyst bed support means allows a treated hydrocarbon feed stream from the first catalyst moving bed to pass therethrough while preventing catalyst particles from the first catalyst moving bed to exit therethrough.  
       [0026] The present invention further accomplishes its desired objects by broadly providing a method for hydroprocessing a hydrocarbon feed stream comprising the steps of:  
       [0027] (a) disposing a first catalyst bed in a first hydroconversion reaction zone of a reactor vessel, wherein the first catalyst bed contains a first set of catalytic properties;  
       [0028] (b) disposing a second catalyst bed in a second hydroconversion reaction zone of the reactor vessel of step (a), wherein the second catalyst bed contains a second set of catalytic properties;  
       [0029] (c) flowing through the first catalyst bed in the first hydroconversion reaction zone a hydrocarbon feed stream to produce an initially-treated hydrocarbon feed stream; and  
       [0030] (d) flowing through the second catalyst bed in the second hydroconversion reaction zone the initially-treated hydrocarbon feed stream to produce a treated hydrocarbon feed stream.  
       [0031] The method additionally comprises withdrawing at least partially spent first catalyst from the first catalyst bed simultaneously with the flowing step (c); and adding fresh first catalyst to the first catalyst bed simultaneously with the flowing step (c). The method further provides for withdrawing at least partially spent second catalyst from the second catalyst bed simultaneously with the flowing step (c); and adding fresh second catalyst to the second catalyst bed simultaneously with the flowing step (c). The flowing step (d) comprises flowing the initially-treated hydrocarbon feed stream through the second catalyst bed without the initially-treated hydrocarbon feed stream having left or exited the reactor vessel of the step (a).  
       [0032] The method additionally includes demetallizing the hydrocarbon feed stream of the step (c) in the first catalyst bed; and hydrotreating the initially-treated hydrocarbon feed stream of the step (d) in the second catalyst bed to remove at least one chemical element from the initially-treated hydrocarbon feed stream. The at least one chemical element is selected from the group consisting of nitrogen, sulfur, and mixtures thereof.  
       [0033] The present invention yet further accomplishes its desired objects by providing a reactor comprising a reactor vessel having an internal cylindrical wall with an internal reactor cylindrical wall diameter; a first distributor plate assembly coupled to and supported by the internal cylindrical wall of the reactor vessel; and a first catalyst bed support assembly coupled to and supported by the internal cylindrical wall of the reactor vessel such as to define a first plenum chamber between the first distributor plate assembly and the first catalyst bed support assembly and for supporting a first catalyst moving bed in a first hydroconversion reaction zone having a first set of physical and/or catalytic properties and for allowing a hydrocarbon feed stream to pass therethrough from the first plenum chamber while preventing catalyst particles from a first catalyst moving bed to exit therethrough. The reactor also comprises a first catalyst outlet means, communicating through the internal cylindrical wall of the reactor vessel with a first catalyst moving bed supported by the first catalyst bed support assembly, for withdrawing a first generally spent catalyst from a first hydroconversion reaction zone; and a first catalyst inlet means, communicating through the internal cylindrical wall of the reactor vessel with a first catalyst moving bed supported by the first catalyst bed support assembly, for introducing a first generally fresh catalyst into a first hydroconversion reaction zone. A second distributor plate assembly is coupled to and supported by the internal cylindrical wall of the reactor vessel; and a second catalyst bed support assembly is also coupled to and supported by the internal cylindrical wall of the reactor vessel such as to define a second plenum chamber between the second distributor assembly and the second catalyst bed support assembly and for supporting a second catalyst moving bed in a second hydroconversion reaction zone having a second set of physical and/or catalytic properties and for allowing a hydrocarbon feed stream to pass therethrough from the second plenum chamber while preventing catalyst particles from a second catalyst moving bed to exit therethrough. The reactor further includes a second catalyst outlet means, communicating through the internal cylindrical wall of the reactor vessel with a second catalyst moving bed supported by the second catalyst bed support assembly, for withdrawing a second generally spent catalyst from a second hydroconversion reaction zone; and a second catalyst inlet means, communicating through the internal cylindrical wall of the reactor vessel with a second catalyst bed supported by the second catalyst bed support assembly, for introducing a second generally fresh catalyst into a second hydroconversion reaction zone.  
       [0034] The reactor additionally comprises a first catalyst moving bed supported by the first catalyst bed support assembly and having a first set of physical and/or catalytic properties; and a second catalyst moving bed supported by the second catalyst bed support assembly and having a second physical and/or catalytic properties. The first set of physical and/or catalytic properties differ from the second set of physical and/or catalytic properties. A first conduit means, communicating through the internal cylindrical wall of the reactor vessel, for introducing a hydrocarbon feed stream into the reactor vessel; and a second conduit means, communicating through the internal cylindrical wall of the reactor vessel, for passing a hydrocarbon product out of the reactor vessel. The first catalyst bed support assembly allows an untreated hydrocarbon feed stream from the first conduit means to pass therethrough while preventing catalyst particles from the first catalyst moving bed to exit therethrough; and the second catalyst bed support assembly allows an initially treated hydrocarbon feed stream from the first catalyst moving bed to pass therethrough while preventing catalyst particles from the second catalyst moving bed to exit therethrough. Alternatively, the second catalyst bed support assembly receives an untreated hydrocarbon feed stream and allows an initially-treated hydrocarbon feed stream from the second catalyst moving bed to pass therethrough while preventing catalyst particles from the second catalyst moving bed to exit therethrough; and the first catalyst bed support assembly allows a treated hydrocarbon feed stream from the first catalyst moving bed to pass therethrough while preventing catalyst particles from the first catalyst moving bed to exit therethrough.  
       [0035] In a preferred embodiment of the invention, the catalyst comprises a plurality of catalytic particulates having a mean diameter ranging from about 35 Tyler mesh to about 3 Tyler mesh; and a size distribution such that at least about 90% by weight of the catalytic particulates have an aspect ratio less than about 2.0 and a diameter ranging from R 1  to about R 2 , wherein:  
       [0036] (1) R 1  has a value ranging from about {fraction (1/64)} inch to about ¼ inch,  
       [0037] (2) R 2  has a value ranging from about {fraction (1/64)} inch to about ¼ inch, and  
       [0038] (3) a value of a ratio R 2 /R 1  ranges from about 1.0 to about 1.4.  
       [0039] The catalyst may be employed in any hydrogenation process. Preferably, the catalyst is for producing a plug-flowing substantially packed bed of hydroprocessing catalyst during hydroprocessing by contacting a substantially packed bed of hydroprocessing catalyst with an upflowing hydrocarbon feed stream. More particularly, when the catalytic particulates are disposed in a hydrocarbon reaction zone, a substantially packed bed of hydroprocessing catalyst is produced; and when a hydrocarbon feed stream flows upwardly through the substantially packed bed of hydroprocessing catalyst, plug-flowing commences when a volume of the catalytic particulates is withdrawn from a bottom of the hydrocarbon reaction zone. As used herein “catalyst” includes other particles which interact with a feed stream, such as sorbents, or other fluid contact bodies. The catalyst is disposed in a reaction zone and a hydrocarbon feed stream is flowed upwardly through the catalyst for hydroprocessing the hydrocarbon feed stream.  
       [0040] The catalytic particulates have a size distribution such that a maximum of about 2.0% by weight of said catalytic particulates have a diameter less than R 1 . The catalytic particulates also have a size distribution such that a maximum of about 0.4% by weight of the catalytic particulates have a diameter less than R 3 , wherein R 3  is less than R 1  and the value of the ratio R 1 /R 3  is about 1.4. The catalytic particulates have a maximum attrition of about 1.0% by weight of the catalytic particulates through a diameter having a value of R 1 ; and the catalytic particulates have a maximum attrition of about 0.4% by weight of the catalytic particulates through a diameter having a value of R 3 , wherein R 3  is less than R 1  and the value of the ratio R 1 /R 3  is about 1.4.  
       [0041] From the foregoing summary it will be apparent that several significant factors contribute directly to the present invention accomplishing its desired objects, and to the efficient use of a given process reactor vessel to assure non-ebullating, plug-like flow of a body of catalyst particles therethrough while being contacted by a counter-flowing hydrocarbon fluid stream of gas and liquid at maximum space-velocity. Among such significant factors are: (i) the size, volume and density characteristics of such catalyst particles at preselectable flow velocities and pressures of the hydrocarbon fluid stream; (ii) control of catalyst bed ebullation and/or levitation during hydrocarbon fluid and hydrogen flows; (iii) laminar flow of the catalyst particles during movement into and out of the catalyst moving bed for replacement (or regeneration or rejuvenation) without bed ebullation or levitation; (iv) and for a preferred catalyst support means, concentric annular feed of alternate rings of the gas and liquid components of the hydrocarbon feed uniformly into the full moving catalyst bed, which is capable of recovering promptly from upset or pressure changes in the reactor vessel to restore such alternate rings of gas and liquid over process runs of extended length (e.g. several thousand hours); and (v) redistribution of the gas components along the axial length of the moving bed.  
       [0042] It is another object of the present invention to broadly provide a method for hydroprocessing a hydrocarbon feed stream through two or more essentially downwardly plug-flowing substantially packed bed of hydroprocessing catalyst within two or more hydroconversion reaction zones within a single onstream reactor.  
       [0043] These, together with the various ancillary objects and features which will become apparent to those skilled in the art as the following description proceeds, are attained by this invention, a preferred embodiment as shown with reference to the accompanying drawings, by way of example only, wherein: 
     
    
    
     BRIEF DESCRIPTION OF THE DRAWINGS  
     [0044]FIG. 1 is a schematic view of a typical hydroprocessing vessel to which the present invention is particularly directed for on-stream catalyst replacement during continuous plug-like flow of catalyst through up-flowing liquid hydrocarbon feed and gas streams;  
     [0045]FIG. 2 is a bottom plan view of the concentric and radial catalyst bed support means for a truncated conical or pyramidal screen, taken in the direction of arrows and along the plane of line  2 - 2  in FIG. 1.;  
     [0046]FIG. 3 is an elevational cross-sectional view of the support means and screen taken in the direction of arrows and along the plane of line  3 - 3  in FIG. 2;  
     [0047]FIG. 4 is a partial elevational view of an alternate form of a laminar flow arrangement for withdrawing deactivated catalyst particles from the reactor bed;  
     [0048]FIG. 5 is a cross-sectional plan view of the reactor vessel taken in the direction of arrows and along the plane of line  5 - 5  in FIG. 1 showing a preferred form of gas redistribution and quench system over a central portion of the catalyst bed;  
     [0049]FIG. 6 is a perspective view, partially in cross-section, of one of the quench or redistribution shed units shown in FIG. 5;  
     [0050]FIG. 7 is a perspective view of a preferred arrangement of two tiers of shed units of FIG. 5 at a given level in the catalyst bed;  
     [0051]FIG. 8 is a partial cross-sectional view illustrating a catalytic bed with a plurality of superimposed layers with respect to each other before commencement of a plug-flow;  
     [0052]FIG. 9 is a partial cross-sectional view illustrating a catalytic bed which is moving downwardly in a plug-flow fashion;  
     [0053]FIG. 10 is a partial cross-sectional view of the reactor and a partial perspective view of another embodiment of the catalytic support means;  
     [0054]FIG. 11 is a partial cross-sectional view of the reactor and the catalytic support means of FIG. 10 which includes a plurality of annular mixture zones under the substantially packed bed of hydroprocessing catalyst with each annular mixture zone containing a liquid hydrocarbon component and a hydrogen-containing gas component and wherein the annular mixture zones are concentric with respect to each other and are coaxial with respect to the reactor and the substantially packed bed of hydroprocessing catalyst;  
     [0055]FIG. 12 is the partial cross-sectional view of the reactor and support means in FIG. 11 with the inert pellets, and illustrating ribs or spokes secured to an imperforate center plate and supporting a plurality of segmented plates;  
     [0056]FIG. 13 is another cross-sectional view of the reactor and support means as similarly illustrated in FIG. 12 with a bed of inert pellets having a liquid hydrocarbon component and a hydrogen-containing gas component flowing around the inert pellets for entering the annular mixture zones;  
     [0057]FIG. 14 is a schematic diagram of another embodiment of the present invention having a moving-bed reactor containing two separate hydroconversion reaction zones;  
     [0058]FIG. 15 is schematic diagram for another embodiment of the present invention depicted in FIG. 14; and  
     [0059]FIG. 16 is another schematic diagram for another embodiment of the present invention depicted in FIGS. 14 and 15. 
    
    
     DETAILED DESCRIPTION OF THE INVENTION INCLUDING PREFERRED EMBODIMENTS OF THE INVENTION  
     [0060] Referring in detail now to the drawings, and initially more particularly to FIG. 1, a hydroprocessing system is shown embodying the method of the present invention to increase substantially both the continued catalytic activity of a single volume or single bed of catalyst  10  and the efficient use of a single reactor vessel of a given reactor volume, such as reactor vessel  11 . Vessel  11 , as indicated by the thickness of its cylindrical side wall  12  and domed closure heads, or ends,  13  and  14 , is designed to react a hydrogen containing gas mixed with a liquid hydrocarbon stream at a pressure of up to about 300 atmospheres (about 4500 lbs per square inch) and up to about 650° C. (about 1200° F.). Such reaction gas and a feed stream of hydrocarbon liquids are preferably premixed and introduced as a single stream through bottom head  13  by line  16 .  
     [0061] To assure maximum catalytic benefit during the hydroprocessing of the hydrocarbon feed stream and the hydrogen-containing gas, it is essential that vessel  11  contain as much catalyst as possible within the design volume of vessel  11 . Accordingly as indicated, support means  17  for catalyst bed  10  is placed as low as possible in vessel  11  while assuring full and adequate dispersion of the hydrogen phase within the liquid hydrocarbon stream. At the same time, the upper limit of bed  10  is near the top of domed head  14 , while providing an adequate space  21  for disengaging any entrained catalyst from the resulting products withdrawn through center pipe  18 . To insure that catalyst is not entrained into product fluids exiting through center pipe  18 , a screen  15  may be installed in space  21  above a bed surface  20  defining the top of the catalyst bed  10 . Fresh catalyst is then added to bed surface  20  through pipe  19  extending through screen  15 . Desirably, the upper level or top of the catalyst bed  10 , designated as the bed surface  20 , is preferably controlled on a continuous basis by gamma ray absorption measurement made possible by a gamma ray source  22  and gamma ray detector  24  positioned in close proximity to the bed surface  20  of catalyst bed  10 . Such a gamma ray source may be in the form of radioactive isotopes, such as Cesium  137 , disposed inside the reactor in a specially designed well. Alternatively the source can be an electrically controllable source, such as a thermal neutron activated gamma ray generator. Detector  24  may be in the form of an ionization tube, Geiger-Mueller tube or a scintillation detector. Suitable sources and detectors are manufactured by Ronan Engineering Co., Texas Nuclear and other vendors. By detecting the level of surface  20 , it is possible, in accordance with the invention, to insure that the catalyst inventory is maintained at the optimum level and that the reactor is never overfilled.  
     [0062] Overfilling the reactor increases the chance that catalyst particles will be crushed in the isolation valves in the transfer lines when they are closed, at the end of each transfer. Bed level control is also needed to confirm that ebullation of the bed is minimized and that undesirable excursions from the design flow rate for hydrogen and hydrocarbon feed flowing upwardly through bed  10  are avoided for the selected catalyst. To this end, the size, shape, and density of catalyst particles supplied to the bed are selected in accordance with the designed maximum rate of flow of the feed streams to prevent such ebullation. Such control assures that bed  10  progressively moves down through vessel  11  in layers as by a plug flow. A “plug flow” of the catalyst bed  10  is illustrated in FIGS. 8 and 9 and may be best described as when a lowermost volumetric layer A is removed, the next volumetric layer B flows downwardly to replace the lowermost volumetric layer B and assumes a new position as a lowermost volumetric layer B. The removed lowermost volumetric layer A is replaced with an upper volumetric layer J. The procedure is again repeated (as best shown by the dotted line representations in FIG. 9) by removing the lowermost volumetric layer B and causing the next volumetric layer C to flow downwardly in a plug-like fashion to replace the lowermost volumetric layer B and assume a new position as a lowermost volumetric layer C. The removed lowermost volumetric layer B is replaced with an upper volumetric layer K. The procedure may be continually repeated to define a downwardly plug-flowing catalyst bed  10  which is moving in direction of arrow W in FIG. 9.  
     [0063] The procedure to determine whether or not a catalyst bed  10  is plug-flowing may be by any suitable procedure. For example, in a preferred embodiment of the present invention wherein metals (e.g. vanadium) are being removed from a hydrocarbon feed stream, the catalyst bed  10  is plug-flowing if a catalytic sample (e.g. 15 catalytic particulates) from withdrawn catalyst is analyzed and it is found through elemental metal analysis that the catalytic sample has a uniform high metal load, preferably at least about 1.5 times more than the average metal load of the catalyst bed  10 , and more preferably at least about 2.0 times more than the average metal load of the catalyst bed  10 . Those possessing the ordinary skill in the art can determine the average load of the catalyst bed  10  from the total amount of metals removed from the hydrocarbon feed stream, the weight of the catalytic bed  10 , etc.  
     [0064] It is to be understood that whenever the specification or the claims states or mentions any type of catalyst movement or catalyst bed  10  movement (e.g. “removing”, “moving”, “supplying”, “replacing”, “delivering”, “flow”, “flowing”, “transfer”, “transferring”, “addition”, “adding”, “admixing”, etc.) for any type or mixture of catalyst without stating or mentioning the basis, the basis for such type of catalyst or catalyst bed movement may be on any type of basis, such as “intermittent basis”, “periodic basis”, “continuous basis”, “semi-continuous basis”, etc. Thus, by way of example only, removal of lowermost volumetric catalytic layers and addition of upper volumetric catalytic layers may be on a “periodic basis”, “a continuous basis”, or even “a one time basis”, all without affecting the spirit and scope of the present invention(s). It is to be also understood that the “removal” or “withdrawal” of catalyst and the “addition” or “replacement” of catalyst are mutually exclusive of each other and may be performed simultaneously or at different times without affecting the spirit and scope and of the present invention(s). Preferably, the “addition” or “replacement” of catalyst is performed after the “removal” or “withdrawal” of catalyst and after the catalyst bed  10  has moved downwardly into a non-moving state or non-moving posture. Catalysts are selected utilizing data acquired by measuring bed expansion, such as in a large pilot plant run, with hydrocarbon, hydrogen and catalyst.  
     [0065] To further assure that plug flow continues throughout the full length of the bed, and particularly at the bottom portion, bed support means  17  is particularly characterized by the truncated polygonal or conical configuration of support means  17 .  
     [0066] As shown in the preferred embodiment of FIGS. 2 and 3, and as best seen in FIG. 2, support  17  includes a series of annular polygons, approaching the form of annular rings, formed by a plurality of segment plates  27  between radial ribs or spokes  26  extending from imperforate center plate  25  to sidewall  12  of vessel  11 . As shown in FIG. 3, spokes  26  may be any suitable geometric shape, such as rod-like (see FIG. 3) or substantially flat plates (see FIG. 10), which divide the circumference of the vessel into many segments (eight in this case) and similarly support the ends of outer octagonal ring  23  of support means  17  formed by annular or circumferential plates  27 . In each case, radial ribs or spokes  26 , and annular segment plates  27  form a plurality of concentric rings, or annular polygons which support conical, or pyramidal, perforated plate or screen  28 . Thus screen  28  is permeable to both gas and liquid rising from the lower portion of vessel  11 .  
     [0067] In one preferred embodiment of the particular merit of the concentric annular polygons as illustrated in FIG. 3, the interconnected plates  27  form a plurality of ring-like structures extending generally axially parallel to the sidewall  12  with the radial ribs or spokes  26  radially extending towards the sidewall  12  of reactor vessel  11 . The mixture of the hydrocarbon liquid feed and hydrogen gas that is to enter the catalyst bed  10  separates by gravity into radially alternate gas and liquid rings, made up of adjacent segments between each pair or radial spokes  26 . Thus, both phases flow upwardly through alternate concentric annular passages under screen  28 . The preferential separation of gas from liquid in each ring includes an annular cap segment of gas overlying an adjacent lower annular segment filled with liquid. Hence, both fluids have equal, and angularly adjacent, access to the bed through screen  28 . The plurality of alternate annular rings of hydrogen gas and hydrocarbon liquid assure even and equal feed of both phases across the full cross-sectional area of screen  28  into bed  10 . Among other factors, we have particularly found that this configuration insures even and equal distribution across the full cross-sectional area of the catalyst bed. Such equal distribution across the full diameter of the bed  10 , permits a quiescent flow section to form directly above center plate  25  which truncates conical bed support means  17 . This decreases substantially potential local ebullation or eddy currents from being induced in the catalyst bed at the point of catalyst withdrawal through inlet  30  of inverted J-tube  29  to assure localized laminar flow of catalyst and liquid from within bed  10 .  
     [0068] Uniform feed of the mixture of the hydrocarbon feed stream and hydrogen is particularly facilitated to the inlet side of plates  27  of support means  17  through plenum or inlet chamber  33  enclosed between support  17  and circular plate member  31 , which extends across the full cross-sectional area of vessel  11 . The circular plate member  31  defines a grid-like structure for supporting a permeable screen  6  having one or more openings, as best shown in FIGS. 11, 12 and  13 . As further best shown in FIGS. 11, 12 and  13 , the permeable screen  6  supports a bed  3  of a plurality of inert pellets  4  (e.g. alumina pellets) which are sized not to pass through the openings in the permeable screen  6 , to prevent eddy currents in the plenum chamber  33 , and to keep bubbles of hydrogen-containing gas diffused within the hydrocarbon feed streams. Plate  31  includes a multiplicity of similar large diameter tubes  32  forming openings through plate  31 . Each tube is several inches in diameter and extends axially to a similar depth, say on the order of  4  to  6  inches, below plate  31 . Tubes  32  provide equal access to the mixture of hydrogen and hydrocarbon feed stream into plenum chamber  33 . Even distribution of the incoming feed stream into bottom header  35  from feed line  16  may also be assisted by deflector plate  34  to assure that oversized bubbles of hydrogen that may be contained in the feed stream will be equally distributed across the full cross-sectional area of plate  31  and equally distributed to each of tubes  32  for flow into plenum chamber  33 . The length of tubes  32  may be selected to form a suitable gas head under plate  31  to suppress surges in the feed streams entering header  35 .  
     [0069] As noted above, the vertical, transverse width or axial length of plates  27  which set off each individual annular and radial segment, provide equal access to both hydrogen and liquid feed into catalyst bed  10 , and stepped under screen  28  so that they effectively form rings of gas and hydrocarbon feed alternately across the full diameter at the inlet side of catalyst bed  10 . In this way, no single area of the inlet to catalyst bed  10  becomes a segregated or preferential, flow path for either gas or the liquid. Further, if pressure surges result in full wetting of screen  28  by the liquid phase, recovery of gas flow is assisted by the aerial breadth of each segment between plates  27  and radial plates  26 .  
     [0070] In another preferred embodiment of the particular merit of the concentric annular polygons as illustrated in FIGS. 10, 11,  12  and  13 , there is seen a liquid hydrocarbon component LH and a hydrogen-containing gas component HG (hydrogen-containing gas bubbles) entering as an LH—HG mixture into the plenum chamber  33  from tubes  32 . The LH—HG mixture is introduced into the plenum chamber  33 . In this preferred embodiment of the present invention, the annular or circumferential plates  27  are secured to and are supported by the radial ribs or spokes  26 , each of which has a vertical or transverse width that is essentially equal to the vertical or transverse width of the annular or circumferential plates  27 . The radial ribs or spokes  26  also function as a means for reducing a size of hydrogen-containing gas bubbles, especially over-size hydrogen-containing gas bubbles from the hydrogen-containing gas component HG. Those skilled in the art will readily recognize that the number of radial ribs or spokes  26  employed will depend on a number of factors, such as the anticipated number of over-size hydrogen-containing gas bubbles in the upwardly flowing hydrocarbon feed stream, the weight of the catalyst bed  10 , etc. The interconnected plates  27  and radial ribs or spokes  26  form a web or web-like structure defining a plurality of annular mixture zones, generally illustrated as MZ in FIGS. 10 and 11. The annular mixture zones MZ are essentially continuous or are generally endless annular mixture zones MZ, and may contain or be subdivided into any reasonable desired number of mixture zones (or sub-mixture zones), such as MZ 1 , MZ 2 , MZ 3 , MZ 4 , MZ 5 , and MZ 6  in FIGS. 10 and 11. Each of the individual mixture zones MZ 1 , MZ 2 , MZ 3 , MZ 4 , MZ 5 , and MZ 6  is for all practical purposes an annularly continuous or endless mixture zone of uniform thickness, excepting a periodic interruption by radially ribs  26 , which are relatively narrow vis-a-vis the spaced distance between any pair of contiguous ribs  26 - 26 . As evident in FIGS. 10, 11,  12  and  13 , concentric with mixture zone MZ 1  and as a partial bottom to same is imperforate center plate  25 , which is preferably spaced from and off of the plate  31  and the screen  6  such that inert pellets  4  may be supported by the screen  6  and the plate  31  immediately underneath the imperforate center plate  25 . Mixture zone MZ 1  is essentially a cylindrical annular mixture zone with an open top and boundaries defined by the space between a plurality of interengaged and coupled plates  27   1 s and the perimeter of the imperforate center plate  25 .  
     [0071] The plurality of annular mixture zones MZ (or the annularly continuous or endless mixture zones MZ 2 s, MZ 3 s, MZ 4 s, MZ 5 s, and MZ 6 s) under the catalyst bed  10  are concentric with respect to each other and are coaxial with respect to the reactor vessel  11  and the catalyst bed  10 . The plates  27  may be radially spaced from each other at any suitable distance (preferably of uniform distance) to assist in accomplishing the desired objects of the present invention; however, preferably the plates  27  are radially spaced from each other at a generally uniform thickness or distance that ranges from about 1 inch to about 4 feet, more preferably from about 6 inches to about 3 feet, most preferably from about 1 foot to about 2 feet. The radially spaced relationship between and among the plates  27  generally defines a uniform thickness for each of the mixture zones (i.e. MZ 2 s, MZ 3 s, etc.). It is to be understood that while the plurality of annular mixture zones MZ is represented in FIGS. 2, 3,  10 ,  11 ,  12  and  13  as being a plurality of non-circular geometric-shaped zones (e.g. octagonal in FIG. 2), the spirit and scope of the present invention includes that the plurality mixture zones MZ may comprise any geometric-shaped zones including not only polygonal-shaped zones, but also a plurality of concentric circular mixture zones, etc., all of which would also be concentric with respect to each other and coaxial with respect to the reactor vessel  11  and/or the catalyst bed  10  (or the hydroconversion reaction zone).  
     [0072] Therefore, the plates  27  function to form generally uniform thick and essentially circular bands of concentric hydrocarbon feed streams that are also coaxial with respect to the catalyst bed  10 . By way of example only and as best shown in FIGS. 2 and 10, angular mixture zone MZ 2  is defined by the eight (8) interengaged or intercoupled plates  27   1 s and the eight (8) interengaged or intercoupled plates  27   2 s. The eight (8) plates  27   1 s and the eight (8) plates  27   2 s each form an annulate boundary for the essentially circular band of hydrocarbon feed stream in mixture zone MZ 2 . Because the spacing or distance between plates  27   1 s and  27   2 s is generally circumferentially uniform, the thickness or size of the essentially circular band of hydrocarbon feed stream in mixture zone MZ 2  is essentially uniform transversely and/or equal in transverse or horizontal cross section. Similarly, mixture zone MZ 6  is defined by the eight (8) interengaged or intercoupled plates  27   5 s and the eight (8) interengaged or intercoupled plates  27   6 s, the combination of which form annulate boundaries for the essentially circular band of hydrocarbon feed stream in mixture zone MZ 6 . As was previously similarly indicated for plates  27   1 s and  27   2 s, because the spacing or distance between plates  27   5 s and  27   6 s is generally circumferentially uniform, the thickness or size of the circular band of hydrocarbon feed stream in mixture zone MZ 6  is essentially uniform transversely and/or equal in transverse or horizontal cross section. Plates  27   2 ,  27   3 ,  27   4 , and  27   5  similarly functionally interengage and intercouple to define annulate boundaries for mixture zones MZ 3 , MZ 4 , and MZ 5 . As indicated and as best shown in FIG. 2, ribs  26  extend radially from imperforate center plate  25  and planarly represent visually pie-shaped segments. Between any pair of contiguous ribs  26 - 26 , the lengths of the respective plates  27  increase from plate  27   1  to or towards plate  27   6  while the widths are essentially the same as best shown in FIG. 3. Thus, plate  27   2  is longer than plate  27   1  while possessing the identical approximate width. Likewise: plate  27   3  is longer than plate  27   2 , plate  27   4  is longer than  27   3 , plate  27   5  is longer than plate  27   4 , and plate  27   6  is longer than plate  27   5 , while all the plates  27  simultaneously have generally the same width or the same longitudinal extension below the screen  28  (see FIG. 3). Thus, the vertical dimensions or the widths of the plates  27  (i.e. the structural extensions of the plates  27  that are generally parallel to the longitudinal axis of the reactor vessel  11  and/or the catalyst bed  10  therein) are generally equal. All plates  27  are preferably spaced such that the hydrocarbon feed stream flows parallel to the longitudinal axis of the catalyst bed  10  before contacting and entering the same. Both the upper edges and lower edges of plates  27   1 s,  27   2 s,  27   3 s,  27   4 s,  27   5 s, and  27   6 s are all at a different level or height, as best shown in FIGS. 3 and 11. The mixture zones MZ differ from a plurality of tubes, conduits, or pipe-like passages for introducing an essentially complete or essentially total integral cylindrical hydrocarbon feed stream into the catalytic bed  10 . As best shown in FIGS. 3 and 11, the upper and lower edges of plates  27   1 s are at a different level or height than the upper and lower edges of plates  27   2 s which are at a different level or height than the upper and lower edges of plates  27   3 s. Similarly, the upper and lower edges of plates  27   3 s are at a different level or height than the upper and lower edges of plates  27   4 s which are at a different level or height than the upper and lower edges of plates  27   5 s. The upper and lower edges of the latter are at a different level or height than the upper and lower edges of plates  27   6 s.  
     [0073] After the LH—HG mixture enters and flows through the screen  6  into the plenum chamber  33 , the flowing LH—HG mixture enters into each of the generally continuous annular mixture zones MZ 2 s, MZ 3 s, etc. for dividing or separating the flowing LH—HG mixture into a plurality of flowing generally continuous annular LH—HG mixtures, which have been designated LH—HG 2 , LH—HG 3 , LH—HG 4 , LH—HG 5  and LH—HG 6  in FIG. 11. As was previously indicated, the mixture zone MZ 1  is also basically an annular or cylindrical shaped mixture zone defined by the space between the perimeter of the imperforate center plate  25  and intercoupled segmented plates  27   1 s and receives hydrocarbon feed stream (i.e. hydrocarbon liquid feed and/or hydrogen gas) in and through the space by which it is being defined. In a preferred embodiment of the present invention and as best shown in FIG. 13, before the flowing LH—HG mixture enters into each of the generally continuous annular mixture zones MZ 1 s, MZ 2 s, MZ 3 s, etc. the LH—HG mixture flows around the plurality of inert pellets  4  in zig-zag fashions for reducing the possibility of eddy currents and for keeping bubbles of hydrogen gas diffused within the liquid hydrocarbon and preventing agglomeration of same into larger size bubbles. The hydrocarbon feed stream entering into mixture zone MZ 1  is designated LH—HG 1 . The plurality of LH—HG mixtures (i.e. LH—HG 1 , LH—HG 2 , etc.) pass through the screen  28  and respectively enter into the catalyst bed  10  from each of the mixture zones (i.e. MZ 1 s, MZ 2 s, MZ 3 s, etc.) at a flow rate such as not to ebullate, levitate or expand the catalyst bed  10  upwardly and/or towards the screen  15  and the domed head  14  by more than 10% by length beyond substantially the full axial length of the bed catalyst  10  in a packed bed state, such as the packed bed state reflected in FIG. 8. The plurality of generally continuous annular LH—HG mixtures flow upwardly through screen  28  and into the catalyst bed  10 . The catalyst bed  10  in the present invention preferably comprises catalyst particles which are substantially the same and/or uniform size, shape, and density and which are selected in accordance with the average optimum velocity of the hydrocarbon feed stream (i.e. a mixture of a liquid hydrocarbon component LH and a hydrogen-containing gas component HG, or the continuous annular LH—HG mixtures) flowing into the plenum chamber  33  and subsequently into and through the plurality of mixture zones MZ 2 s, MZ 3 s, etc. The rates of flow of the plurality of the respective LH—HG mixtures (i.e. LH—HG 1 , LH—HG 2 , etc.) from the respective mixture zones MZ 1 s, MZ 2 s, etc., and thus also the flow rates of the hydrocarbon feed stream into plenum chamber  33  from and through line  16 , are all to be controlled in an amount and to an extent sufficient to maintain expansion or levitation of the catalyst bed  10  to less than 10% by length over or beyond substantially the full axial length of the bed  10  in a packed bed state. More preferably, the expansion of the substantially packed bed of catalyst is limited to less than 5%, most preferably less than 2% or even less than 1%, by length over or beyond substantially the full axial length of the bed  10  in a packed bed state. Ideally the expansion of the substantially packed bed of catalyst is limited to essentially 0% by length.  
     [0074] The flow rate of the hydrocarbon feed stream through line  16  is to be at a rate not substantially greater than the optimum rate of flow. The optimum rate of process fluid flow through the substantially packed bed of catalyst will vary from process unit to process unit based on several factors including oil and hydrogen feed characteristics, catalyst specifications, process objectives, etc. Based on catalyst particles having substantially the same and/or uniform size, shape and density, the flow rate of the hydrocarbon feed stream preferably ranges from about 0.01 ft/sec to about 10.00 ft/sec and more preferably from about 0.01 ft/sec to about 1.00 ft/sec. Similarly and/or likewise and further based on the catalyst particles having substantially the same and/or uniform size, shape, and density, the flow rate of the continuous annular LH—HG mixtures (i.e. the summation of the flow rates for LH—HG 1  through LH—HG 6  from mixture zones MZ 1 s through MZ 2 s respectively in FIG. 11) is to be at a rate not substantially greater than the optimum rate of flow, preferably ranging from about 0.01 ft/sec to about 10.00 ft/sec, and more preferably from about 0.01 ft/sec to about 1.00 ft/sec. The specific flow rate would depend as indicated on a number of variables, such as the particular application (e.g. demetallation or desulfurization etc.) of the hydroprocessing process. The specific flow rates however would be at any suitable rate controlled in an amount and to an extent sufficient to limit expansion of the substantially packed bed of catalyst to less than 10% by length over or beyond a substantially packed bed of hydroprocessing catalyst in a packed bed state.  
     [0075] In a preferred embodiment of the invention and for such a flow rate for the hydrocarbon feed stream and for such a flow rate for the continuous annular LH—HG mixtures, the catalyst particles preferably have the substantially same and/or uniform size, shape and density in order to obtain over the desired demetallization and/or desulfurization and/or denitrification of the liquid hydrocarbon component LH in the hydrocarbon feed stream (i.e. LH—HC mixture) into produced hydrogen upgraded product fluids that are being withdrawn from the reactor vessel  11  through the center pipe  18 . At the above indicated flow rates for the hydrocarbon feed stream flowing through line  16 , and for the flow rates for the generally continuous annular LH—HG mixtures (i.e. LH—HG 1 , LH—HG 2 , etc.), the produced upgraded product fluids are being preferably withdrawn through the center pipe  18  from the reactor vessel  11  at a rate ranging from about 0.01 ft/sec to about 10.00 ft/sec and more preferably from about 0.01 ft/sec to about 1.00 ft/sec. The withdrawal rate(s) of the produced upgraded product fluids is not to be greater than the optimum rate of flow and will also vary from process unit to process unit based on several factors including oil and hydrogen feed characteristics, catalyst specifications, process objectives, etc. The specific withdrawal rate(s) would be any suitable withdrawal rate, controlled in an amount and to an extent sufficient to prevent and/or limit expansion of the substantially packed bed of catalyst to less than 10% (more preferably less than 5%, most preferably less than 2% or even less than 1%) by length over or beyond substantially the full axial length of the bed  10  in a packed bed state.  
     [0076] The arrangement in inlet distributor  31  for uniformly distributing hydrogen gas and liquid hydrocarbon feed as shown in FIG. 3 may be modified by lengthening or shortening tubes  32 , forming uniformly distributed cylindrical passageways into plenum chamber  33 . A particular advantage of using tubes  32 , as compared to merely perforations or holes of adequate diameter, lies in the formation of a gas pocket under plate  31  in the areas around the individual tubes  32 . We have found that this is desirable because such a gas pocket trapped beneath tray or plate  31  provides pressure surge dampening, which may result from flow changes of the mixture of hydrogen and liquid being supplied to the reactor vessel. However, the length of the tubes  32  is maintained as short as reasonably possible to so function. Again, this is because of the desirability of utilizing as little as possible of all processing space available in vessel  11  for anything but contacting the feed streams with conversion catalyst. A particular advantage to using tubes  32 , as compared to a combination of tubes and perforations, is that the designed flow distribution pattern is maintained over a wider range of flow rates. With tubes and perforations, gas normally flows up the perforations and liquid flows up the tubes. However, gas will find new flow paths through the tubes if the gas flow increases or the perforations become plugged, resulting in undesigned and potentially undesirable flow patterns.  
     [0077] To further assist in maintenance of plug-like flow of catalyst bed  10  throughout its axial length, there is additionally provided in a preferred form or embodiment of the invention a plurality of axially spaced apart hydrogen gas redistribution or hydrogen gas-quenching stages  39  within bed  10 . In the arrangement of FIG. 1, the location of one of the gas redistribution stages  39  is illustrated by the single inverted angle member  40  extending transverse to the axis of bed  10 . The details of quench system  39  are best seen in FIGS.  5  to  7  where a plurality of inverted V-shaped sheds  40  (i.e. inverted angle members  40 ) are equally distributed over at least one transverse row extending generally across the cross-sectional area of vessel  11 . As shown in FIG. 6 and in FIG. 7, a gas injection line  42  feeds an elongated tube  41  extending through each individual shed  40  from a header  44  and branch lines  45  supplying the individual tubes  41 . Desirably, but not necessarily, a second tier of sheds  40  is axially spaced above the first tier, with the sheds  40  in each tier being positioned at 90 degree(s) to the other tier, as shown in FIG. 7. Construction of an individual shed  40  is best seen in FIG. 6, wherein distribution pipe  41  includes a plurality of discharge holes  48 , desirably proportioned to give equal distribution of hydrogen gas along the full length of tube  41 . Desirably, holes  48  are on the top side of tube  41  so that gas leaving the tube is forced to flow downwardly within shed  40  to join gas rising from bed  10  under the area enclosed by the V-sides  49  of shed  40 . Preferably, the full length of each skirt formed by sides  49  includes equally spaced slots  50  to exhaust both rising gas from bed  10  and quench gas entering from line  42 . A particular value of the present arrangement is that gas which may have become channeled in a portion of the bed below the quench system can be redistributed across the full cross-sectional area of the bed to further avoid generation of local hot spots, eddy currents, or ebullation, within the upper portion of bed  10 .  
     [0078] In accordance with another significant aspect of the present invention, FIG. 1 shows a catalyst replacement system, which in general comprises a series of lock chambers for transferring fresh catalyst into bed  10  through a pair of pressure lock chambers, including charging vessel  60  and supply vessel  70 . A similar series of lock chambers, including discharge vessel  80  and disposal vessel  90 , transfer catalyst out of bed  10 . If necessary, a single pair of vessels could be used to charge and discharge the catalyst, although the piping and sequencing procedure would be more complex. In both cases, transfer flow is specifically designed to be as a liquid slurry and laminar to avoid undue abrasion of catalyst particles going into reactor vessel  11  and to avoid abrupt agitation of the overlying bed of particles, with consequent ebullation and eddying of catalyst or fines in bed  10 , when catalyst is withdrawn through inlet  30  of J-tube  29  at the bottom of reactor vessel  11 .  
     [0079] To achieve laminar flow for supply of catalyst from charging vessel  60  to the top of reactor vessel  11  or for catalyst removal from the bottom of bed  10  to discharge vessel  80 , it is essential that the pressure differential between reactor vessel  11  and vessels  60  or  80 , be accurately controlled as by detecting the pressure differences between supply line  61  or discharge line  82  and reactor vessel  11 . The pressure difference is best zero when shut-off valves  64  or  84  are first opened or closed. The pressure differences between vessel  11  and line  61  is measured by gage  63  and pressure detectors  62  and  65 . Differential pressure gage  83  and detectors  81  and  85  serve a similar function to control transfer of catalyst through valve  84  from the bottom of reactor vessel  11  to discharge vessel  80 .  
     [0080] With reference particularly to supply of catalyst from vessel  60 , it will be understood, of course, that the vessel  60  is capable of being brought to a slightly higher pressure than the operating pressure of reactor vessel  11 , and closely controlled to assure that catalyst supplied to vessel  60  from supply vessel  70  is by laminar flow. For this purpose, as indicated, vessels  70  and  60  are at atmospheric pressure, catalyst is first introduced into supply vessel  70  by way of funnel  100  through line  101  and valve  102 , and nitrogen is preferably flushed through supply vessel  70  through line  106  and/or line  71  to eliminate air and moisture that may be present on the catalyst. Either before or after catalyst is introduced, vessel  70  is charged with a distillate hydrocarbon stream, preferably gas oil, to provide the necessary slurrying liquid for mixing and transporting catalyst. This may either be through funnel  100 , valve  102 , and line  101 , or through line  104 , valve  105  and line  106 . Valve  102  is then closed and the catalyst is then preferably heated to dehydrate and eliminate water from the catalyst. It is to be understood that whenever the specification or the claims states, mentions, or implies “mixing” or “admixing” or “commingling”, or any of the like, including of any type(s) of catalyst, such stated, mentioned, or implied verbiage means within the spirit and scope of the present invention any type of “mixing” or “admixing” or “commingling”, or any of the like, including any incidental mixing or any otherwise non-thorough/non-homogeneous mixing. Preferably, however, any type of “mixing” or “admixing” or commingling”, or any of the like, will be essentially thorough and/or essentially homogeneous.  
     [0081] An important requirement is that before transferring liquid to the charging vessel  60 , the pressure in supply vessel  70  must be equalized to that in charging vessel  60 , assuming, of course, that isolation valve  64  between vessel  60  and the reactor vessel  11  is closed, and also that valves  67 ,  68  and  78  are closed. With valves  64 ,  67 ,  68 ,  78  and  102  closed and pressure equalized between the vessels  60  and  70 , transfer valve  75  may be opened to provide the same diameter path for the catalyst slurry to flow throughout the path from J-tube  71  to vessel  60 . The transfer is closely controlled by regulating the nitrogen gas flow rate and pressure introduced from line  104  through valve  105 . The pressure and flow rate are just sufficient to assure the desired laminar flow of catalyst into inlet  72  of J-tube  71  and thus upwardly through line  76  and into charging vessel  60 , which forms a catalyst charging vessel. Laminar flow to transfer catalyst through J-tube  71  is entirely in the liquid phase, with the catalyst as a slurry in the gas oil. Transfer of all catalyst is assisted by the funnel shape of bottom  79  of vessel  70 , and the position of intake  72  to J-tube  71  at the apex of bottom  79 . If all the catalyst in vessel  70  is transferred to vessel  60 , flush oil from vessel  70  will naturally clear all the catalyst out of line  76 . However, to assure that all such catalyst has passed through valve  75  (so that valve  75  need not close on hard, abrasive catalyst with potential danger of scoring the valve  75  or the valve seat therein) additional flush fluid is preferably introduced from line  77  through valve  78  to clear line  76 , either back into vessel  70 , or forward into vessel  60 .  
     [0082] With catalyst thus loaded into vessel  60 , a similar procedure is used for transferring catalyst under laminar flow conditions as a liquid slurry into reactor vessel  11  through supply pipe  61  for distribution to the top  20  of bed  10 . If desired, of course, a deflector plate (not shown) may be used to distribute catalyst evenly across the top of catalyst bed  20 . However, we have found that such a distribution aid is not required. In the transfer of catalyst from the charging vessel  60  to reactor vessel  11 , it will be understood that the pressure in vessel  60  is brought to the pressure of reactor vessel  11 . This is done by injecting process hydrogen through valve  67 . The oil should be heated to a temperature as close as possible to the temperature of reactants in vessel  11 , without vaporizing the oil. We have found this to be particularly important to minimize any disturbance of the hydroprocessing process when fresh catalyst is added to an onstream reactor vessel, such as reactor vessel  11 . Once these requirements are met, valve  64  should be opened for transfer. The actual laminar transfer of the liquid slurry is controlled by valve  67  throttling the flow and pressure of hydrogen admitted from line  66 . After transfer of the catalyst, valve  68  in flush line  69  is opened briefly to assure that any catalyst left in lines  61  and  19  is cleared before valve  64  is closed, for the reasons noted before. Excess hydrogen pressure in vessel  60  may be relieved in a controlled manner via a suitable bleed line running back to the common hydrogen source (not shown) of the hydroprocessing system.  
     [0083] Substantially continuous or intermittent transfer of deactivated catalyst for regeneration or disposal from the bottom of catalyst bed  10  in reactor vessel  11  to discharge vessel  80  is controlled in the same manner. As in all transfer of catalyst throughout the system of the present invention depicted in FIG. 1, the flow path from inlet  30  of J-tube  29 , through line  82 , including the bore of valve  84 , is uniform in cross-sectional area and diameter. Similarly, transfer from discharge vessel  80  to disposal vessel  90  is through inlet  89  of J-tube  86  to discharge outlet  98  of line  92 , including valve  94 , into vessel  90 . Deactivated catalyst is transferred laminarly from the bottom of the catalyst bed  10  as a slurry in the hydrocarbon feed stream which, as previously mentioned, comprises the liquid hydrocarbon feed stream or a mixture of hydrocarbon liquid feed and hydrogen-containing gas. Typically, the catalyst is transferred essentially in the liquid hydrocarbon feed stream (i.e. the liquid component of the hydrocarbon feed stream).  
     [0084] In general the diameter of these laminar flow passageways are at least five times, and may be as high as fifty or more times, the diameter of the individual particles to be passed therethrough. In this connection to avoid jamming or obstruction, the inlets  72 ,  109 ,  30 ,  89  and  99  into their respective tubes  71 ,  108 ,  29 ,  86  and  96  are not flared or otherwise restricted, or perforated, so that all flow is solely and directly through the full and equal bore of such inlets. In the case of catalyst removal from reactor vessel  11 , inlet  30  of tube  29  is positioned at and over unperforated center plate  25  of catalyst support screen means  17 , so that it is out of the direct flow of any hydrogen gas stream rising through the innermost annular passageway formed by walls  27  and radial ribs or spokes  26 . This assures that flow into entry  30  is substantially a liquid only slurry mixture with catalyst particles. Such a mixture at laminar flow conditions produces maximum carrying capacity of the fluid. Additionally, the external dimensions of the circular bend or arc portion of the J-section of the tube  29  is several times the diameter of inlet  30  and the connected flow path, including the downwardly directed portion. The portion of tube  29  above inlet  30  is many times shorter and smaller in volume than the remainder of J-tube  29 , down to, and including, control valve  84 . A particular advantage of keeping this portion of tube  29  small is to avoid the necessity of forcing substantial amounts of catalyst back into the bed  11  against the gravity head of catalyst bed  10  when that portion of the line is cleared at the end of each transfer.  
     [0085] Desirably, during periods when the catalyst is not being transferred, a small amount of hydrogen may be continually bled through valve  88  into bed  10  through J-tube  29  to assure that catalyst particles do not clog entry  30 . This avoids potential build up of coke at entry  30  of pipe  29 . Such an arrangement assures that catalyst can be withdrawn by laminar flow without artificially fluidizing or levitating bed  10  directly adjacent to J-tube entry  30 .  
     [0086] Because gravity drainage of catalyst by an opening through the center of the catalyst support screen means  17  is not required in the present arrangement, as in the prior art, it is possible to operate the entire system without use of solids handling valves. Accordingly, each of the transfer valves in the present arrangement are preferably conventional ball valves formed with a single through bore in a rotatable ball. Specifically, we have found that conventional valves used to feed and control flow of hydrocarbons, catalyst and hydrogen, into and out of the vessel  11 , must seal against high pressure differentials between the vessel and the transfer vessels. For this service, a solid satellite, spherical-ball valve having a through bore of the same diameter as the inlet and outlet lines to the valve and metal-to-metal seals, provides superior service when used in the catalyst transfer lines for carrying out the method of the present invention. Further, their commercial cost and ready availability for such severity of service makes them most useful economically, both for initial installation and for service replacement. Valves manufactured by The Kaymr and Mogas Companies, called full-port valves are particularly useful in the present embodiment. Further, the arrangement permits transfer of catalyst almost exclusively in a liquid phase which substantially reduces abrasion or comminution of catalyst particles during transfer. Additionally, exclusion of entrained gas substantially improves the efficiency of liquid transfer of catalyst particles and further reduces potential damage to the catalyst.  
     [0087]FIG. 4 illustrates a partial view of the bottom of pyramidal catalyst bed support means  17  showing an alternate system for transferring catalyst in a laminarly flowing liquid. In this embodiment, an L-valve is formed by vertical tube  54  and horizontal tube  52  for withdrawing catalyst particles from the bottom of bed  10 . As shown, intake  56  is preferably directly above the central, non-perforated, section  25  of the truncated pyramid formed by support means  17 . While such an arrangement is less preferred than that shown in the embodiment of FIG. 1, such an arrangement is made suitable by the fact that the slurry of liquid and catalyst can be made to flow only under laminar flow conditions. With either the J-tube of FIG. 1, or the L-valve of FIG. 4, arrangements, the pressure in discharge vessel  80  is brought up to equal that in reactor vessel  11 . Valve  84  is opened and catalyst flow is controlled, as seen in FIG. 1, by regulating flow through valve  93 . Such flow decreases the gas-pressure in discharge vessel  80  and line  82  sufficiently to induce a laminar flow of catalyst particles from vessel  11  when transfer valve  84  is opened. After valve  84  has been flushed with vacuum gas oil through valve  88  and line  87  and then closed. The pressure in vessel  80  is then reduced to a lower pressure (about 50 psig or less). The residuum is drained from discharge vessel  80  through drain line  120 , below J-tube  86  and conical screen  121 . Flush oil is then sent in through valve  93  to wash residuum off the catalyst and to cool the catalyst. The discharge vessel  80  can be drained and filled as many times as needed. The pressure in disposal vessel  90  is made equal to that in vessel  80  and valve  94  is opened. The flow and pressure are then controlled through valve  110  to induce laminar flow of catalyst through J-tube  86  and into disposal vessel  90 . Valve  94  is flushed with flush oil through valve  107  and closed. The contents of the disposal vessel  90  is preferably washed and cooled with flush oil which is then drained through drain line  122  below conical screen  123 . The spent catalyst contents of the disposal vessel  90  is then washed with water if desired through valve  110 . The disposal vessel  90  should be purged of any hydrogen by sending in nitrogen gas also through valve  110 . Finally, disposal vessel  90  is nearly depressurized and the catalyst is dumped using water as the carrier fluid through J-tube  96  by nitrogen flow through valve  110  to control the rate of catalyst flow in discharge pipe  124 .  
     [0088] Continuing to refer to the drawings for other preferred embodiments of the present invention, a method is provided for maximally occupying a reactor volume with a substantially packed bed of hydroprocessing catalyst (e.g. catalyst bed  10 ) during hydroprocessing by contacting the substantially packed bed of hydroprocessing catalyst with an upflowing hydrocarbon feed stream having a liquid component and a hydrogen-containing gas component. The substantially packed bed of hydroprocessing catalyst, preferably comprising the catalyst of the present invention as more particularly described below under the subtitle “The Catalyst”, is disposed in a reactor zone (or reaction zone or zone for reaction) contained within a reactor volume (e.g. the entire internal volumetric space available within the reactor vessel  11 ) such that the substantially packed bed of hydroprocessing catalyst occupies from about 50% by volume to about 98% by volume of the reactor volume; more preferably from about 50% by volume to about 98% by volume; most preferably at least about 90% by volume or from about 90% by volume to about 95% by volume of the reactor volume. Stated alternatively, hydroprocessing catalyst is disposed or otherwise positioned within a reactor volume such that hydroprocessing catalyst occupies from about 50% by volume to about 98% by volume of the reactor volume; more preferably from about 50% by volume to about 98% by volume; most preferably at least about 90% by volume or from about 90% by volume to about 95% by volume of the reactor volume. “Reactor volume” (or the entire internal volumetric space available within the reactor vessel  11 ) means or may be generally defined as the volumetric space within the reactor vessel  11  (or any similar hydroprocessing reactor vessel), including the summation or addition of the following internal volumes: (i) an internal volume within the reactor vessel  11  represented by a volume (or internal cylindrical volume or main body volume of the reactor vessel  11 ) spanning or extending from an upper tangent line  180  (see FIG. 8) to a lower tangent line  182  and generally illustrated as arrow TL in FIG. 8; and (ii) an internal volume within the upper dome closure end  14  (or hemispherical head) of the reactor vessel  11 ; and (iii) an internal volume within the lower dome closure end  13  (or hemispherical bottom) of the reactor vessel  11 . A “tangent line” is known to those skilled in the art as a plane (i.e. horizontal plane) taken generally along the junctures of the sidewall  12  (which is essentially a straight upright wall) of the reactor vessel  11  with the upper and lower dome closure ends  14  and  13  respectively.  
     [0089] A hydroprocessing feed stream including a liquid component and a hydrogen-containing gas component upflows into the substantially packed bed of hydroprocessing catalyst at a rate of flow such that expansion of the substantially packed bed of hydroprocessing catalyst is limited to less than 10% by length beyond a substantially full axial length of the substantially packed bed of hydroprocessing catalyst in a packed bed state. A volume of the hydroprocessing catalyst is withdrawn from the reactor zone to commence essentially plug-flowing downwardly of the substantially packed bed of hydroprocessing catalyst within the reactor zone; and hydroprocessing replacement catalyst is added to the essentially plug-flowing downwardly, substantially packed bed of hydroprocessing catalyst at a rate to substantially replace the volume of the withdrawn hydroprocessing catalyst. The procedure may be repeated as many times as desired, even continuously repeated during continual hydroprocessing.  
     [0090] Another method is provided for hydroprocessing a hydrocarbon feed stream that is upflowing through a hydroconversion reaction zone having a substantially packed bed of catalyst which comprises forming a plurality of annular mixture zones under a hydroconversion reaction zone having a substantially packed bed of hydroprocessing catalyst such that each of the annular mixture zones contains a hydrocarbon feed stream having a liquid component and a hydrogen-containing gas component and wherein the annular mixture zones are concentric with respect to each other and are coaxial with respect to the hydroconversion reaction zone. The hydrocarbon feed stream from each of the annular mixture zones is introduced into the substantially packed bed of hydroprocessing catalyst to commence upflowing of the hydrocarbon feed stream from each of the annular mixture zones through the substantially packed bed of the catalyst.  
     [0091] Considering the range of hydroconversion systems and/or hydroconversion reaction zones which could benefit from the preferred embodiments of the present invention, one skilled in the art will appreciate the variety of catalysts, having a variety of physical properties and elemental compositions, which could be used in such a range of systems. It is within the spirit and scope of the present invention to encompass these systems employing catalysts having a size, shape and density which vary widely from system to system. However, it is important for the present preferred embodiment that the catalyst particles be of uniform and/or same size, and shape (same density when in fresh catalyst state) within a single hydroconversion reaction zone of a hydroconversion system, in order to achieve the desired catalyst and hydrocarbon flow patterns within the hydroconversion reaction zone. A detailed description of the preferred catalyst is presented below under the subtitle “THE CATALYST”. It is to be understood that whenever the specification or the claims states, mentions, or implies “fresh catalyst”, such stated, mentioned, or implied “fresh catalyst” means within the spirit and scope of the present invention any type of catalyst having any usable catalyst life or activity (e.g. regenerated catalyst, rejuvenated catalyst, partially fouled catalyst obtained from any source, etc.). Preferably, “fresh catalyst” means a type of catalyst that has never been used before and is obtained directly from a manufacturer with the lowest desired density and the highest desired catalyst life or activity.  
     [0092] A hydroconversion system and/or a hydroconversion reaction zone of a present preferred embodiment of the present invention contains a catalyst which is described in detail below under the subtitle “The Catalyst”, and may also be operated as a fixed bed (i.e. a catalyst bed which does not expand), a moving bed, an ebullated bed, an expanded bed or a fluidized bed configuration. A moving bed system is preferred.  
     [0093] By “moving bed”, as used herein, is meant a reaction zone configuration in which a catalyst is added at one end of a catalyst bed in an intermittent or substantially continuous manner and is withdrawn at the other end in an intermittent or substantially continuous manner. A “moving bed” also includes a “plug-flow” or “plug flowing” catalyst bed  10  or substantially packed bed of catalyst. As previously indicated, when any type of catalyst or catalyst bed  10  movement is mentioned, stated, or implied, the spirit and scope of the present invention includes such type of movement on any type of basis or in any manner (e.g. “periodic”, “fully continuous”, “non-continuous” etc.) without the necessity of having to specifically mention the type of basis or manner. Preferably, catalyst is added at the top of the reaction zone and withdrawn at the bottom. In the type of moving bed to which the present preferred embodiment is directed, the catalyst particles in the bed are substantially in contact with one another and plug-flow downwardly. The catalyst bed is not significantly expanded when process fluids (e.g., liquid and gas) passes through it. It has essentially the character of a fixed bed except for maybe a slight expansion upwardly and for the addition and removal of catalyst. As the term is used herein, a “moving bed” is not the same as a “fluidized bed”, “ebullating bed” or “expanded bed”. In fluidized beds, the flow rate of a single phase fluid, relative to the particles of the catalyst, is fast enough so that the catalyst behaves like a fluid, with particles circulating throughout the bed or even being carried out of the bed with the products. Ebullating or expanded beds are very similar to fluidized beds, except that the relative rate of flow of two phase fluids (e.g., liquid and gas) is regulated to expand the catalyst bed in random motion between 110% and 140% of the height of the catalyst in a “slumped” or packed state. The typical ebullating bed reactor will have a mass of solid particles whose gross volume in the reaction vessel is at least 10 percent larger when feed is flowing through it, as compared to the stationary mass with no feed flowing through it. Although the particles in the bed do not necessarily circulate as if they were fluids, they are separated from one another and go through random motion.  
     [0094] Several advantages ensue from use of a moving bed reactor. By establishing and maintaining appropriate gas and liquid velocities in packed bed type reactors, just below the threshold of inertia that would cause the catalyst bed to fluidize and/or “channel” and/or lift the catalyst into random motion, the uniform catalyst characteristics described above will allow the catalyst to migrate downward through the reactors in a predictable plugflow manner, as catalyst batches are withdrawn from the reactor bottom. And further, by maintaining plug flow catalyst movement downward within the reactors (e.g. reactor  11 ), the catalyst within the reactors can be maintained in layers having differing activity levels and reaction rates. The number of catalyst layers depend on the frequency of catalyst addition and withdrawal, and the amount added and withdrawn in any given period of time. Typically, however, the number of different aged catalyst layers within the reactor (e.g. reactor  11 ) will be in the range from 10 to 60.  
     [0095] Intermittent or continuous catalyst additions and withdrawal may be used. Catalyst replacement rates can range from several percent of the charge per day to several percent of the charge per week, depending on the reactor size, catalyst metals loading capacity, feed rate, and feed composition and processing objectives. Fresh catalyst is introduced into the downstream end of the catalyst bed (e.g. catalyst bed  10 ), and a corresponding volume of deactivated catalyst is removed from the upstream end of the catalyst bed, at a rate which is sufficient to maintain the actual overall average level of catalytic upgrading activity of the bed as a whole at or above the selected minimum average activity level. By “upstream” end of the catalyst bed (e.g. catalyst bed  10 ), as used herein, is meant the end of the moving bed into which the heavy hydrocarbonaceous feed is introduced. By “downstream” end of the catalyst bed is meant the end of the bed from which the process effluent is recovered. In a normal gravity flow system, the catalyst is added and effluent removed at the top of the vessel (the downstream end). Spent catalyst is withdrawn and feed introduced at the bottom (the upstream end).  
     [0096] In a particularly important application of the present invention, catalyst is continuously added at the top of the reactor (e.g. reactor  11 ) to the slowly moving bed (e.g. bed  10 ), and spent (and deactivated catalyst) catalyst is continuously withdrawn from bottom of the slowly moving bed. The deactivated catalyst is removed from the reactor (e.g. reactor  11 ) after it has been deactivated to a substantially lower level of activity than an acceptable minimum average level of activity of the overall catalyst bed. This allows more efficient and complete use of the catalyst activity, e.g. its metals capacity, for such feed upgrading functions as demetallation. As previously indicated, spent (and deactivated) catalyst is withdrawn from the bottom of a reactor in a hydrocarbon liquid. One of the features of the present invention is that the hydrocarbon liquid that is withdrawing and transporting catalyst is the liquid hydrocarbon component LH which is intended to flow upwardly through the bed of catalyst but has not. Thus, one of the features of the present invention is that the hydrocarbon liquid for transporting spent (and deactivated) catalyst is an unconverted liquid hydrocarbon component LH or a partially converted liquid hydrocarbon component LH or a mixture of both; and the transporting hydrocarbon liquid (i.e. the liquid hydrocarbon component LH) has not passed entirely upwardly through the catalyst bed.  
     [0097] The product from the method of the present invention exits a reactor (e.g. reactor vessel  11 ) and is normally subjected to further conventional refinery processing. All or part of the product can be passed to a conventional, fixed bed upgrading operation, such as a hydrodesulfurization operation. Part of the product stream can be recycled, either for further catalytic treatment or as a diluent. Treatment of heavy feeds by catalytic demetallation according to the present process followed by fixed bed desulfurization is particularly effective, but all or part of a demetallized product from the countercurrent demetallation reaction zone can also be processed in a countercurrent moving bed desulfurization reaction zone.  
     [0098] The present preferred embodiments of the present invention are applicable to hydroconversion reaction zones for hydrocracking, hydrodemetallization, hydrotreating, hydrodesulfurization, hydrodenitrification, hydrofinishing and the like, all of which catalytically upgrade a heavy hydrocarbonaceous oil that represents the liquid hydrocarbon stream or liquid hydrocarbon feed stream (i.e. the liquid hydrocarbon component LH). By “heavy” liquid hydrocarbon stream, as used herein and as previously indicated, is meant liquid hydrocarbon stream at least 50 volume percent of which boils above about 2040 C and which preferably contains a substantial fraction boiling above about 343° C. and particularly preferably above about 510° C. Preferred liquid hydrocarbon streams are residual fractions and synthetic crudes. They can be derived from crude petroleum, from coal, from oil shale, from tar sand bitumen, from heavy tar oils, and from other synthetic sources. The present invention is advantageously employed to refine highly refractory and contaminated liquid hydrocarbon streams. The liquid hydrocarbon stream may be substantially free from finely divided solids such as shale fines, sand or the like. Alternatively, the liquid hydrocarbon stream may contain a substantial concentration (e.g. about 1 weight percent or more) of finely divided solids. As previously indicated, the liquid hydrocarbon stream (i.e. the liquid hydrocarbon component LH) is preferably premixed with any type of hydrogen-containing gas (i.e. the liquid hydrocarbon component HG) which is preferably hydrogen, before being introduced into the reactor vessel  11  as a single stream or hydrocarbon stream. The mixing ratios of the liquid hydrocarbon stream to the hydrocarbon containing gas may be any suitable ratio, well known to those artisans possessing the ordinary skill in the art.  
     [0099] Typically, a heavy hydrocarbonaceous oil employed as a hydrocarbon feed stream in the present invention contains undesirable metals. Undesirable metals which are often present in hydrocarbonaceous feeds notably include nickel, vanadium, arsenic, and iron. These metals deactivate conventional, fixed bed catalysts (such as fixed bed hydroprocessing catalysts) and also rapidly and irreversibly deactivate catalysts when high metals level feed are charged directly to conventional units. These metals are often present as organo-metallic compounds. Thus, the use of the terminology “iron, nickel, arsenic or vanadium compounds” means, those metals in any state in which they may be present in the hydrocarbon feed stream in the process of the present invention, either as metal particles, inorganic metal compounds, or an organo-metallic compounds. Where amounts of metals are referred to herein, the amounts are given by weight based on the metal itself. For maximum efficiency in such a countercurrent demetallation process, the hydrocarbon feed stream should have levels of undesirable metals greater than about 150 ppm by weight of the hydrocarbon feed stream, preferably greater than about 200 ppm by weight of the hydrocarbon feed stream, and more preferably greater than about 400 ppm by weight. Although nickel, vanadium, arsenic, and iron are the usual metal contaminants, other undesired metals, such as sodium and calcium, may also contribute to the metals content of the hydrocarbon feed stream for purposes of catalytic demetallation upgrading processing.  
     [0100] Catalytic upgrading conditions (e.g. catalytic desulfurization conditions, catalytic hydrogenation conditions such as designed for asphaltenes saturation, catalytic denitrification conditions and catalytic hydrocracking conditions, etc.) employed in the hydroconversions reaction zones within the reactor vessel  11  for preferred embodiments of the present invention all include a reaction temperature generally in the range of from about 230° C. to about 480° C., a pressure generally in the range of from about 30 to about 300 atmospheres, a hydrogen rate ranging from about 1000 to about 10,000 standard cubic feet per barrel of feed, and a liquid hourly space velocity (LHSV) in the range of from about 0.20 hr−1 to about 10 hr−1. For feed demetallation upgrading, the temperatures and pressures within the reaction zone can be those typical for conventional demetallation processing. The pressure is typically above about 500 psig (514.7 psia; 35.5 bar). The temperature is typically greater than about 315° C., and preferably above 371° C. Generally, the higher the temperature, the faster the metals are removed; but the higher the temperature, the less efficiently the metals loading capacity of the demetallation catalyst is used. While demetallation reaction can be conducted in the absence of added hydrogen, hydrogen is generally used and therefore requires full and equal distribution into the moving bed along with any gases evolving from the feed.  
     [0101] In carrying out a process of the preferred embodiments of the present invention, a minimum average level of catalytic feed upgrading activity for the countercurrently moving catalyst bed (e.g. catalyst bed  10 ) as a whole is selected for the particular catalytic upgrading reaction. For a moving bed (e.g. catalyst bed  10 ) in a demetallation reaction system, for example, the minimum average upgrading activity level for the catalyst bed is one which removes the necessary amount of metals from the hydrocarbon feed stream when it passes through the moving bed at demetallation conditions. Similarly, for a desulfurization reaction system, the moving catalyst bed (e.g. catalyst bed  10 ) removes the necessary amount of sulfur from the hydrocarbon feed stream when it passes through the moving bed at desulfurization conditions. Thus, as will be apparent to those skilled artisans, the minimum average upgrading activity level for a particular reaction system will depend on the desired degree of a contaminant, such as metals, sulfur, nitrogen, asphaltenes, etc., which the refiner desires to remove from the heavy oil feed. The degree of demetallation or desulfurization (or etc.) will typically be set by economics and the downstream processing that the heavy feed will undergo. Further, according to preferred embodiments of the present invention, the actual average level of catalytic upgrading activity for the catalyst bed (e.g. catalyst bed  10 ) as a whole is measured. Measurement of the actual average level of upgrading is made by determining the extent to which the hydrocarbon feed stream is being upgraded in the countercurrent moving bed system. For example, when upgrading involves demetallation, demetallation activity is measured by a determination of the residual concentration of metals remaining in the liquid effluent stream from the moving bed system. When upgrading involves desulfurization, desulfurization activity is, analogously, measured by a determination of the residual concentration of sulfur remaining in the liquid effluent from the reaction system. Overall catalyst bed upgrading activity for other upgrading reactions is measured in a similar manner by determining the residual amount of the containment which is to be removed by the process. In the present preferred embodiments, the rate at which catalyst is removed from the reaction zone, and the rate of catalyst replacement to the reaction zone, is established by a number of economic and operating factors, which include maintaining a desired average level of catalytic upgrading activity.  
     [0102] Referring in detail now to FIGS.  14 - 16  for another embodiment of the present invention and wherein similar parts of the present invention depicted in FIGS.  14 - 16  are identified by like reference numerals that depict the same similar parts of the invention depicted in FIGS.  1 - 13 , there is seen a pair of moving catalyst bed  10 - 10  contained in the reactor vessel  11 . Each moving catalyst bed  10  is in communication with the pipes  19  wherein fresh catalyst is added onto the bed surface  20 . Each moving catalyst bed  10  is also in communication with the discharge line  82  wherethrough at least partially spent catalyst is withdrawn. In FIG. 14, pipe  19  passes through domed head  14  for adding fresh catalyst to the top moving catalyst bed  10 . The center pipe  18  in FIG. 14 receives a hydrocarbon feed which flows downwardly through the top or upper moving catalyst bed  10 , and subsequently downwardly through the bottom moving catalyst bed  10  for discharge through line  16  which is functioning as a discharge line instead of a feed line. Similarly, center pipe  18  in FIG. 14 is functioning as a feed line instead of a discharge line.  
     [0103] Referring particularly to FIGS. 15 and 16 center pipe  18  is functioning as a discharge line wherethrough treated hydrocarbon streams exit the reactor vessel  11 . Pipe  19  passes through center pipe  18  in a coaxial manner in order to deposit fresh catalyst upon the bed surface  20  of the top catalyst bed  10 . In FIG. 15 another pipe  19  passes through the domed head  14 , through the entire top moving catalyst bed  10 , through the catalyst support means  17 , and then through the circular plate member  31  (i.e. a distributor plate assembly with depending hollow tubes  32 ) for depositing catalyst upon the bed surface  20  of the bottom or lower moving catalyst bed  10 . In FIG. 16 the second pipe  19  passes through the cylindrical side wall  12  of the reactor  11 .  
     [0104] The discharge line  82  has various embodiments. In FIG. 14 the discharge line passes through the bottom domed head  13  for removing at least partially spent catalyst from the bottom or lower moving catalyst bed  10 . In FIGS. 15 and 16 the discharge line  82  passes generally coaxially into line  16  and through the circular plate member  31  and through the imperforate center plate  25  of the catalyst support means  17  for removing at least partially spent catalyst from the bottom or lower moving catalyst bed  10 . In FIG. 15 a second discharge line  82  passes through the bottom domed head  13 , through the bottom or lower moving catalyst bed  10 , and terminates into the J-tube  29  after passing through the circular plate member  31  and the catalyst support means  17  for removing at least partially spent catalyst from the top or upper moving catalyst bed  10 . In FIG. 16 the discharge line  82  passes through the cylindrical side wall  12 , through the circular plate member  31 , and then through the imperforate center plate  25  of the catalyst support means  17 , which supports the top or upper moving catalyst bed  10 , for withdrawing at least partially spent catalyst from the top moving catalyst bed  10 . In FIG. 14 the discharge line  82  also passes through the cylindrical side wall  12  for withdrawing partially spent catalyst from the top moving catalyst bed  10 .  
     [0105]FIG. 14 is a general schematic diagram of the present invention, generally illustrating the reactor vessel  11  having two separate hydroconversion reaction zones including a pair of the moving catalyst beds  10 - 10 . Each of the hydroconversion reaction zones has an inlet (i.e. pipe  19 ) and an outlet (i.e. discharge line  82 ) for respectively charging the respective moving catalyst beds  10  with fresh catalyst and for respectively discharging at least partially spent catalyst from the respective moving catalyst beds  10 . The hydrocarbon feed stream in FIG. 14 enters the reactor vessel  11  through the center pipe  18  and treated hydrocarbon product fluids exit the reactor vessel  11  through line  16 . No catalyst enters or leaves with the hydrocarbon feed stream or with the hydrocarbon product stream; instead catalyst enters (i.e. through pipe  19 ) and leaves (discharge line  82 ) through the separate connections provided for the catalyst. Although FIG. 14 illustrates a cocurrent moving-bed, feed and catalyst downflow, the present invention may also be used for a countercurrent moving-bed, upflow feed and downflow catalyst, as best shown in FIGS. 15 and 16.  
     [0106]FIGS. 15 and 16 illustrate a more preferred embodiment of the present invention. Both FIGS. 15 and 16 illustrate a countercurrent moving catalyst bed reactor system with the two separate moving catalyst beds  10 - 10 . Hydrocarbon feed stream enters the reactor vessel  11  through line  16 , and after hydrotreatment within the two moving catalyst beds  10 - 10 , a hydrocarbon product stream is withdrawn through center pipe  18 . As previously indicated, fresh catalyst is added to both moving catalyst beds  10 - 10  via pipes  19 - 19  and exits the respective moving catalyst beds  10  through discharge lines  82 . The preferred catalyst support means  17  (which includes screen  28 ) has been described in detail above. Screen  28  of the catalyst support means  17  allows gas and liquid to flow through them but does not allow catalyst to exit through them. Therefore, screens  28  of the catalyst support means  17 - 17  keep the catalyst within its appropriate hydroconversion zone.  
     [0107] In FIGS. 15 and 16, each of the catalyst support means  17  is a conical shaped catalyst support means  17  (including conical shaped screens  28 ) for directing catalyst in a plug-flow fashion towards discharge lines  82 . The angle of each of the conical shaped catalyst support means  17  including the screens  28  associated therewith, is preferably between about 10° and about 80°, more preferably between about 25° and about 70°. The circular plate members  31 - 31  including the depending tubes  32  function as a distributor assembly for distributing both gas and liquid into the plenum chambers  3 - 33  in the upflow type reactor vessels  11  in FIGS. 15 and 16.  
     [0108] It is to be understood that the specific designs for reactor internals within the reactor vessels  11  in FIGS.  14 - 16  are not critical to the present invention. Any internals capable of providing multiple hydroconversion reaction zones, each supplied with catalyst displacement and replacement systems would accomplish the purpose of the present invention. Designs similar to FIGS. 15 and 16 could be devised for cocurrent flow of hydrocarbon feed streams through catalyst. Obviously different reactor internals would be required, but the basic overall design would remain the same.  
     [0109] It is also to be understood that while only two hydroconversion reaction zones (i.e. the pair of catalyst beds  10 - 10 ) are shown in FIGS.  14 - 16 , the spirit and scope of the present invention would include more then two hydroconversion zones (i.e. more then two catalyst beds  10 ). Therefore, if one was interested in using three or more moving catalyst beds  10  within the single moving bed reactor vessel  11 , three or more hydroconversion zones could be provided for the catalyst.  
     [0110] Each of the moving catalyst beds  10  would consist of catalyst particles of different catalytic characteristics, such as different physical and/or catalytic properties. Considering the range of hydroconversion reaction capabilities which each of the moving catalyst beds  10  could possess, one skilled in the art will appreciate the variety of catalysts having a variety of physical properties and elemental compositions, which could be used for each of the moving catalyst beds  10 . It is within the spirit and scope of the present invention to encompass hydroconversion reaction systems employing catalysts possessing such shape and density which vary widely from system to system. Each of the moving catalyst beds  10  would typically have a particular hydroconversion reaction (e.g. hydrometallation, hydrodenitrification, hydrodesulfurization; etc.) capability that differs from the particular hydroconversion reaction capability of the other moving catalyst bed(s)  10 . Thus, by way of example only, the catalyst particles in the bottom or lower moving catalyst bed  10  could possess hydrodemetallation capabilities, where as the catalyst particles in the top or upper moving catalyst bed  10  could posses hydrodenitrification and/or hydrosulfurization and/or any hydrotreating capabilities. A broad and general description of any suitable catalyst for any of the two or more moving catalyst beds  10  is described below under the subtitle “THE CATALYST”, especially for a preferred embodiment of the present invention where each of the two or more moving catalyst beds  10  is a “plug-flowing” catalyst bed  10  possessing different catalytic capabilities (i.e. different physical and/or catalytic and/or elemental properties).  
     [0111] Continuing to refer to FIGS.  14 - 16  for a method for hydroprocessing a feed stream, the top catalyst bed  10  and the bottom catalyst bed  10  are respectfully disposed in the respective hydroconversion reaction zones of the reactor vessel  11 . As previously indicated, the top or upper first catalyst bed  10  contains a first set of catalytic properties or characteristics that differ from the catalytic properties and characteristics of the bottom catalyst bed  10 . A hydrocarbon feed stream flows through the bottom catalyst bed  10  to produce an initially-treated hydrocarbon feed stream. From the bottom catalyst bed  10 , the initially-treated hydrocarbon feed stream flows (i.e. flows through distributor plate  31  into plenum chamber  33  and subsequently from the plenum chamber  33  through the catalyst support means  17 ) through the upper catalyst bed  10  to produce a treated hydrocarbon feed stream, more specifically a hydrocarbon product that exits through center pipe  18 . At least partially spent catalyst is withdrawn from both of the catalyst beds  10 - 10  through the discharge lines  82  simultaneously with the flowing of a hydrocarbon stream through the catalyst beds  10 - 10 . Similarly, fresh catalyst is added to both of the catalyst beds  10 - 10  through pipes  19  also simultaneously with the flowing of the hydrocarbon stream through the catalyst beds  10 - 10 .  
     [0112] The initially-treated hydrocarbon feed stream flows from the lower catalyst bed  10  directly into the upper catalyst bed  10  without having left or exited the reactor vessel  11 . Preferably, the lower or bottom catalyst bed  10  demetallizes a hydrocarbon feed stream in producing the initially-treated hydrocarbon feed stream, and the top catalyst bed  10  hydrotreats (i.e. hydrodenitrifies and/or hydrodesulfurizes and/or hydrodemetallizes) the initially-treated hydrocarbon feed stream to remove at least one chemical element from the initially-treated hydrocarbon feed stream. The at least one chemical element would be selected from the group consisting of nitrogen, sulfur, or any other element, and mixtures thereof.  
     The Catalyst  
     [0113] In a preferred embodiment of the invention, the catalyst which is charged to the reactor vessel  11  preferably satisfies the following four main criteria: (i) the catalyst has the appropriate catalytic activity and life for the particular application (e.g. demetallation, denitrification, hydrodesulfurization, etc.); (ii) the catalyst has physical properties which minimize its random motion in the reactor vessel  11 ; (iii) the catalyst has physical properties which minimize catalyst loss both in the catalyst transfer steps and in the reactor vessel  11 ; and (iv) the catalyst is sufficiently uniform in size and shape and density to prevent classification by size in normal operation.  
     [0114] The catalyst in the present invention preferably has the appropriate catalytic activity and life for the specific application (e.g. demetallation, denitrification, hydrodesulfurization, etc.). For example, if the catalyst is to be used for demetallation, it should have sufficient HDM activity and metals loading capacity (i.e. life) to meet the target demetallation without the use of uneconomic amounts of catalyst. The metals loading capacity of the catalyst is preferably greater than about 0.10 grams of metal per cubic centimeter of catalyst bulk volume and is more preferably greater than about 0.20 grams metal per cubic centimeter of catalyst bulk volume. The catalyst properties which most affect catalytic activity and metals loading capability are: pore structure (pore volume and pore size distribution); base material (e.g. alumina versus silica); catalytic metals (amount, distribution, and type (nickel, molybdenum, cobalt, etc.)); surface area; and particle size and shape. If the catalyst is to be used for denitrification and/or desulfurization or any other hydrotreating purposes, the same foregoing principles and criteria would apply.  
     [0115] The catalyst in the present invention also preferably has physical properties which minimize catalyst lifting into random motion in the upflow type reactor vessel  11 . Since one of the benefits of the present invention is the countercurrent contacting that is achieved between the reactants and catalyst, it is preferred to maintain plug flow of the catalyst downwards through the entire length of the reactor vessel  11 . The catalyst properties which are critical to minimizing or preventing catalyst expansion are: catalyst particle density (highest particle density possible is preferred while still meeting catalytic activity and metals loading requirements); particle size (largest size practical is preferred); skeletal density (higher skeletal density is preferred to reduce skeletal buoyancy); and size uniformity. One of the salient features of the present invention is that the catalyst will not expand into random motion in the reactor vessel  11 , but will still move rather easily during flow transportation.  
     [0116] Under actual process conditions within the reactor, significantly smaller catalysts could rise to the top while significantly larger catalysts could migrate to the bottom. This intervenes with optimal plug flow movement of catalyst. For this reason, size specifications for the catalysts of the present invention are narrower than those for fully packed or fixed bed and ebullated bed catalysts.  
     [0117] The catalyst of the present invention should further have physical properties which minimize catalyst loss in the catalyst transfer steps and in the reactor vessel  11 . Breakage or attrition of the catalyst in either the transfer steps or in the reactor vessel  11 , can have significant adverse effects on the performance of the reactor system itself and on any downstream equipment or processing unit. The following catalyst properties are critical to catalyst loss: catalyst attrition (minimum attrition is absolutely required); catalyst crush strength (maximum crush strength is required without producing a catalyst which is very brittle and might suffer from excessive attrition); catalyst size and shape (spherical catalyst are preferred since they move more easily and have no rough or sharp edges to break off); and fines content (minimum fines is an absolute requirement to avoid adverse effects in the reactor vessel  11  and downstream equipment).  
     [0118] The catalyst is sufficiently uniform in size and shape and density to prevent classification by size in normal operation. Generally, narrow specifications are required for the catalyst to prevent classification by size. Specific catalyst size is selected so that it is near the point of being expanded into random motion, but not to the point of expansion into random motion per se or ebullation.  
     [0119] All of the four main criteria for the selection of the catalyst of the present invention are important and are not independent or mutually exclusive of each other. The four main criteria must be balanced against each other to optimize the catalyst for the specific application. For example, to minimize catalyst expansion into random motion we would prefer a large and very dense catalyst. This is contrary to the properties we might want for a residuum demetallation application where we need a small particle with low density diameters. These competing needs must be balanced to ensure minimum catalyst expansion or ebullation while achieving adequate catalytic activity and metals loading capability, minimum attrition and minimum classification by size.  
     [0120] Because there are competing catalyst requirements and because each application is unique, the catalyst for the present invention may be any suitable catalyst that is capable of assisting in the operation of the invention and assisting in accomplishing the desired objects of the invention.  
     [0121] The catalyst of the present invention unexpectedly produces a plug-flowing substantially packed bed (i.e. catalyst bed  10 ) of hydroprocessing catalyst during hydroprocessing by contacting a substantially packed bed of hydroprocessing catalyst with an hydrocarbon feed stream (i.e. a liquid component and a hydrogen-containing gas component) that is upflowing at a rate controlled in an amount and to an extent sufficient to limit expansion of the substantially packed bed of hydroprocessing catalyst to less than 10% by length beyond a substantially full axial length of the substantially packed bed of hydroprocessing catalyst in a packed bed state. More preferably, the expansion of the substantially packed bed of hydroprocessing catalyst is limited to less than 5%, most preferably less than 2% or even less than 1% , by length beyond a substantially full axial length of the substantially packed bed of hydroprocessing catalyst in a packed bed state. The rate of flow of the hydrocarbon feed stream may be any suitable rate controlled in an amount and to an extent sufficient to limit the expansion of the substantially packed bed of hydroprocessing catalyst, preferably the rate of flow is at a rate ranging from about 0.01 ft/sec. to about 10.00 ft/sec.  
     [0122] The catalyst of the present invention more specifically unexpectedly produces a plug-flowing substantially packed bed of hydroprocessing catalyst when a volume of the hydroprocessing catalyst is withdrawn or transferred under preferably laminar flow conditions from the bottom of the substantially packed bed of hydroprocessing catalyst while, and simultaneously to, the substantially packed bed of hydroprocessing catalyst maximally and optimally occupies at least about 50% by volume, preferably from about 80% by volume to about 98% by volume (i.e. the entire internal and/or inside available volume) of the reactor vessel  11 . The substantially packed bed of hydroprocessing catalyst of the present invention maximally and optimally occupies a volume within the reactor vessel  11  that is larger or greater than a volume of a bed of catalyst in an ebullating reactor vessel that has substantially the same entire internal and/or inside available volume as the reactor vessel  11  and wherein the volume of the bed of catalyst in the ebullating reactor vessel is in a “slumped” (or packed) catalyst bed condition or state. Typically, a bed of catalyst in an ebullating reactor vessel in a “slumped” catalyst bed condition occupies approximately up to less than about 75% by volume (maximum) of the entire internal and/or inside available volume of the ebullating reactor vessel. Thus, the substantially packed bed of hydroprocessing catalyst preferably maximally and optimally occupies at least about 75% by volume, preferably from about 80% by volume to about 98% by volume of the entire internal and/or inside available volume of the reactor vessel  11 . Most preferably, the substantially packed bed of hydroprocessing catalyst of the present invention maximally and optimally occupies from about 75% by volume to about 95% by volume of the entire internal and/or inside available volume of the reactor vessel  11 .  
     [0123] The catalyst of the present invention furthermore specifically unexpectedly produces the plug-flowing substantially packed bed of hydroprocessing catalyst when the volume of the hydroprocessing catalyst is withdrawn or transferred in the hydrocarbon feed stream under preferably laminar flow conditions from a central portion or section of the substantially packed bed of hydroprocessing catalyst and at a lowermost or bottommost section thereof and below the entry points of the plurality of annular mixture zones MZ containing the hydrocarbon feed stream (i.e. a liquid component and a hydrocarbon-containing gas component). As previously indicated, when the volume of the hydroprocessing catalyst of the present invention is withdrawn or transferred to commence plug-flow, it is transferred or withdrawn preferably laminarly in the liquid component of the hydrocarbon feed stream and is removed from above and in proximity to an impervious zone (i.e. imperforate center plate  25 ) of the bed support means  17  and substantially out of the flow path of the LH—HG mixtures (i.e. LH—HG 2 , LH—HG 3 , etc.) emanating out of the mixture zones MZ (i.e. MZ 2 s, MZ 3 s, etc.). The particular volume (or amount) of catalyst that is withdrawn at any desired time from the bottom of the substantially packed bed of hydroprocessing catalyst may be any suitable volume or amount which accomplishes the desired objectives of the present invention. Preferably, such as by way of example only, the particular volume or amount of catalyst that is withdrawn at any desired time is a volume or amount ranging from about 0.10% by weight to about 25.00% by weight of the substantially packed bed (i.e. catalyst bed  10 ). The rate of withdrawal of a particular volume (or amount) of catalyst may also be any suitable volume or amount which accomplishes the desired objectives of the present invention, such as a rate of withdrawal where the flow rate of the catalyst (e.g. the catalyst in the hydrocarbon feed stream) ranges from about 0.1 ft/sec. to about 20 ft/sec., more preferably from about 0.1 ft/sec. to about 10 ft/sec., and at a catalyst concentration ranging from about 0.10 lbs catalyst/lb. catalyst slurry (i.e. weight of hydroprocessing catalyst plus weight of hydrocarbon feed stream) to about 0.80 lbs catalyst/lb. catalyst slurry, more preferably from about 0.15 lbs catalyst/lb. catalyst slurry to about 0.60 lbs catalyst/lb. catalyst slurry. As previously indicated, the withdrawn catalyst may be conveniently replaced by introducing a volume of fresh catalyst through the top of the reactor vessel  11  onto the catalyst bed  10 . The replacement or catalyst addition rate may be any suitable replacement or catalyst addition rate which will accomplish the desired objects of the present invention, such as a flow replacement rate of the replacement catalyst (i.e. the replacement catalyst in the hydrocarbon refined stream (e.g. gas oil)) ranging from about 0.1 ft/sec. to about 20 ft/sec., more preferably from about 0.1 ft/sec. to about 10 ft/sec., and at a catalyst replacement concentration ranging from about 0.10 lbs. replacement catalyst/lb. catalyst slurry (i.e. weight of replacement catalyst plus the hydrocarbon refined stream (e.g. gas oil) as the slurrying medium) to about 0.80 lbs replacement catalyst/lb. catalyst slurry, more preferably from about 0.15 lbs catalyst/lb. catalyst slurry to about 0.60 lbs catalyst/lb. catalyst slurry.  
     [0124] In a preferred embodiment of the present invention, the catalyst of the present invention comprises an inorganic support which may include zeolites, inorganic oxides, such as silica, alumina, magnesia, titania and mixtures thereof, or any of the amorphous refractory inorganic oxides of Group II, III or IV elements, or compositions of the inorganic oxides. The inorganic support more preferably comprises a porous carrier material, such as alumina, silica, silica-alumina, or crystalline aluminosilicate. Deposited on and/or in the inorganic support or porous carrier material is one or more metals or compounds of metals, such as oxides, where the metals are selected from the groups Ib, Vb, VIb, VIIb, and VIII of the Periodic System. Typical examples of these metals are iron, cobalt, nickel, tungsten, molybdenum, chromium, vanadium, copper, palladium, and platinum as well as combinations thereof. Preference is given to molybdenum, tungsten, nickel, and cobalt, and combinations thereof. Suitable examples of catalyst of the preferred type comprise nickel-tungsten, nickel-molybdenum, cobalt-molybdenum or nickel-cobalt-molybdenum deposited on and/or in a porous inorganic oxide selected from the group consisting of silica, alumina, magnesia, titania, zirconia, thoria, boria or hafnia or compositions of the inorganic oxides, such as silica-alumina, silica-magnesia, alumina-magnesia and the like.  
     [0125] The catalyst of the present invention may further comprise additives which alter the activity and/or metals (and/or nitrogen and/or sulfur and/or any other elements) loading characteristics of the catalyst, such as but not limited to phosphorus and clays (including pillared clays). Such additives may be present in any suitable quantities, depending on the particular application for the hydroconversion process including the applied catalyst. Typically, such additives would comprise essentially from about zero (0)% by weight to about 10.0% by weight, calculated on the weight of the total catalyst (i.e. inorganic oxide support plus metal oxides).  
     [0126] Although the metal components (i.e. cobalt, molybdenum, etc.) may be present in any suitable amount, the catalyst of the present invention preferably comprises from about 0.1 to about 60 percent by weight of metal component(s) calculated on the weight of the total catalyst (i.e. inorganic oxide support plus metal oxides) or and more preferably of from about 0.2 to about 40 percent by weight of the total catalyst, and most preferably from about 0.5 to about 30 percent by weight of the total catalyst. The metals of Group VIII are generally applied in a minor or lesser quantity ranging from about 0.1 to about 30 percent by weight, more preferably from about 0.1 to about 10 percent by weight; and the metals of Group VIB are generally applied in a major or greater quantity ranging from about 0.5 to about 50 percent by weight, more preferably from about 0.5 to about 30 percent by weight; while as previously mentioned above, the total amount of metal components on the porous inorganic support is preferably up to about 60 percent by weight (more preferably up to about 40 percent by weight) of the total catalyst. The atomic ratio of the Group VII and Group VIB metals may vary within wide ranges, preferably from about 0.01 to about 15, more preferably from about 0.05 to about 10, and most preferably from about 0.1 to about 5. The atomic ratios would depend on the particular hydroprocessing application for the catalyst and/or on the processing objectives.  
     [0127] The groups in the Periodic System referred to above are from the Periodic Table of the Elements as published in Lange&#39;s Handbook of Chemistry (Twelfth Edition) edited by John A. Dean and copyrighted 1979 by McGraw-Hill, Inc., or as published in The Condensed Chemical Dictionary (Tenth Edition) revised by Gessner G. Hawley and copyrighted 1981 by Litton Educational Publishing Inc.  
     [0128] In a more preferred embodiment for the catalyst, the oxidic hydrotreating catalyst or metal oxide component carried by or borne by the inorganic support or porous carrier material is molybdenum oxide (MoO 3 ) or a combination of MoO 3  and nickel oxide (NiO) where the MoO 3  is present in the greater amount. The porous inorganic support is more preferably alumina. The Mo is present on the catalyst inorganic support (alumina) in an amount ranging from about 0.5 to about 50 percent by weight, preferably from about 0.5 to about 30 percent by weight, most preferably from about 1.0 to about 20 percent by weight, based on the combined weight of the inorganic support and metal oxide(s). When nickel (Ni) is present it will be in amounts ranging up to about 30 percent by weight, preferably from about 0.5 to about 20 percent by weight, more preferably from about 0.5 to about 10 percent by weight, based on the combined weight of the catalyst inorganic support and metal oxide(s). The oxidic hydrotreating catalyst or metal oxide component may be prepared by any suitable technique, such as by depositing aqueous solutions of the metal oxide(s) on the porous inorganic support material, followed by drying and calcining. Catalyst preparative techniques in general are conventional and well known and can include impregnation, mulling, co-precipitation and the like, followed by calcination.  
     [0129] The catalyst has a surface area (such as measured by the B.E.T. method) sufficient to achieve the hydroprocessing objectives of the particular application. Surface area is typically from about 50 sq. meters per gram to about 300 sq. meters per gram, more typically from about 75 sq. meters per gram to about 150 sq. meters per gram.  
     [0130] The catalyst mean crush strength should be a minimum of about 5 lbs. Crush strength may be determined on a statistical sample of catalytic particulates. For example, a fixed number (say 30 catalyst particles) are obtained from a statistical lot comprising a plurality of catalyst particles that are to be employed in the hydrogenation process of the present invention. Each catalyst particle is subsequently disposed between two horizontal and parallel steel plates. A force is then applied to the top steel plate until the disposed catalyst particle breaks. The force applied to break the catalyst particle is the crush strength. The test is repeated for the remaining catalyst particles, and a mean crush strength is obtained. Preferably further, no more than about 35% by wt. of the catalyst particles or catalytic particulates have a mean crush strength of less than about 5 lbs.; more preferably, no more than about 15% by wt. of the catalyst particles or catalytic particulates have a mean crush strength of less than about 5 lbs; and most preferably, no more than about 0% by wt.  
     [0131] The catalyst of the present invention comprises a plurality of catalytic particulates having a uniform size, which is preferably spherical with a mean diameter having a value ranging from about 35 Tyler mesh to about 3 Tyler mesh, more preferably ranging from about 20 Tyler mesh to about 4 Tyler mesh, and most preferably from about 14 Tyler mesh to about 5 Tyler mesh. The Tyler mesh designations referred to herein are from a table entitled “Tyler Standard Screen Scale Sieves” in the 1981 Edition of Handbook 53, published by CE Tyler Combustion Engineering, Inc., 50 Washington St., South Norwalk, Conn. 06856.  
     [0132] Likewise, the preferred catalyst particle has a uniformly smooth and rounded surface. Preferred shapes include, for example, spheres, spheroids, egg-shaped particles and the like. More preferably, the catalyst of the present invention, is a rounded particle having an aspect ratio of less than about 2.0, more preferably equal to or less than about 1.5, most preferably about 1.0. More preferably, the catalyst of the present invention has a size distribution such that at least about 90% by weight (more preferably at least about 95% by weight; most preferably at least about 97% by weight) of the catalytic particulates contained within the catalyst bed  10  have an aspect ratio of less than about 2.0, more preferably less than about 1.5, and most preferably about 1.0. As used herein, “aspect ratio” is a geometric term defined by the value of the maximum projection of a catalyst particle divided by the value of the width of the catalyst particle. The “maximum projection” is the maximum possible catalyst particle projection. This is sometimes called the maximum caliper dimension and is the largest dimension in the maximum cross-section of the catalyst particle. The “width” of a catalyst particle is the catalyst particle projection perpendicular to the maximum projection and is the largest dimension of the catalyst particle perpendicular to the maximum projection.  
     [0133] The catalyst should have a particle size distribution such that the catalyst bed  10  expands under the conditions within the reactor vessel  11  to less than 10% by length (more preferably less than 5% and most preferably less than 1% by length) beyond a substantially full axial length of the substantially packed bed of the hydroprocessing catalyst in a packed bed state. In order to maximize reactor throughput, the catalytic particulates have a narrow size distribution. The catalyst employed in the hydrogenation process of the present invention broadly comprises a size range or size distribution such that at least about 90% by weight, preferably at least about 95% by weight, more preferably, at least about 97% by weight, of the catalytic particulates in the catalyst bed  10  have a diameter ranging from R 1  to R 2 , wherein: (i) R 1  has a value ranging from about {fraction (1/64)} inch (i.e. the approximate opening size of a 35 mesh Tyler screen) to about ¼ inch (i.e. the approximate opening size of a  3  mesh Tyler screen); (ii) R 2  also has a value ranging from about {fraction (1/64)} inch (i.e. the approximate opening size of a 35 mesh Tyler screen) to about ¼ inch (i.e. the approximate opening size of a 3 mesh Tyler screen); and (iii) the ratio R 2 /R 1  has a value greater than or equal to about 1 and less than or equal to about 1.4 (or about the square root of 2.0). More preferably, the catalytic particulates in the catalyst bed  10  have a diameter ranging from R 1  to R 2  wherein R 1  and R 2  each has a value ranging from about {fraction (2/64)} inch (i.e. the approximate opening size of a 20 mesh Tyler screen) to about {fraction (12/64)} inch (i.e. the approximate opening size of a 4 mesh Tyler screen), most preferably from about {fraction (3/64)} inch (i.e. the approximate opening size of a 14 mesh Tyler screen) to about {fraction (9/64)} inch (i.e. the approximate opening size of a 5 mesh Tyler screen), and wherein the ratio R 2 /R 1  has a value ranging from about 1.00 to about 1.4 (or about the square root of 2.0).  
     [0134] The catalyst employed in the hydrogenation process of the present invention also broadly comprises a size range or size distribution such that a maximum of about 2.0% by weight (more preferably a maximum of about 1.0% by weight and most preferably a maximum of about 0.5% by weight or less) of the catalyst particles or catalytic particulates has a diameter less than R 1 . The catalyst also has a size range or size distribution such that a maximum of about 0.4% by weight (more preferably a maximum of about 0.2% by weight and most preferably a maximum of about 0.1% by weight or less) of the catalyst particles or catalytic particulates have a diameter less than R 3 , wherein R 3  is less than R 1  and the value of the ratio R 1 /R 3  is about 1.4 (or about the square root of 2.0). The catalyst particles or catalytic particulates of the catalyst preferably have a maximum attrition of about 1.0% by weight (more preferably a maximum of about 0.5% by weight and most preferably a maximum of about 0.25% by weight or less) of the catalyst particles or catalytic particulates through a diameter (i.e., a Tyler screen opening) having a value of R 1 , and a further maximum attrition of about 0.4% by weight (more preferably a maximum attrition of about 0.2% by weight and most preferably a maximum attrition of about 0.1% by weight or less) of the catalyst particles or catalytic particulates through a diameter (i.e., again a Tyler screen opening) having a value of R 3  wherein R 3  again (as stated above) is less than R 1  and the value of the ratio of R 1 /R 3  is about 1.4 (or about the square root of 2.0). [Note that the attrition procedure is specified in ASTM D 4058-87. However, in the standard method, the fines are removed through a 850T ( ˜ 20 mesh) screen. In the present method, the screen is an opening equal to the minimum catalyst size desired for the particular application, as more specifically defined by the value of R 1  and R 3 .] Thus, by way of example only, for a catalyst with a specified size range of about 10 to about 12 Tyler mesh, one would specify up to about 2.0% by wt. fines (more preferably up to about 1.0% by wt.) MAX through 12 Tyler mesh and up to about 0.4% by wt. (more preferably up to about 0.2% by wt.) MAX through 14 Tyler mesh. Similarly, for a catalyst with a specified size range of about 6 to about 8 Tyler mesh, one would specify up to about 2.0% by wt. fines (more preferably up to about 1.0% by wt. fines) MAX through 8 Tyler mesh and up to about 0.4% by wt. fines (more preferably up about 0.2% by wt. fines) MAX through 10 Tyler mesh. For the catalyst with the specified size range of about 10 to about 12 mesh, one would specify an attrition of up to about 1.0% by wt. (more preferably up to about 0.5% by wt., most preferably up to about 0.25% by wt.) MAX through 12 Tyler mesh and up to about 0.4% by wt., (more preferably up to about 0.2% by wt., most preferably up to about 0.1% by wt.) MAX through 14 Tyler mesh. Similarly further, for catalyst with the specified size range of about 6 to about 8 Tyler mesh, one would specify an attrition of up to about 1.0% by wt. (more preferably up to about 0.5% by wt., most preferably up to about 0.25% by wt.) MAX through 8 Tyler mesh and up to about 0.4% by wt. (more preferably up to about 0.2% by wt., and most preferably up to about 0.1% by wt.) MAX through 10 Tyler mesh.  
     [0135] The specific particle density of the catalyst particles is determined by the requirements of the hydroconversion process. For the present invention it is preferred that the catalyst particles have a uniform density. By “uniform density” is meant that the density of at least about 70% by weight, preferably at least about 80% by weight, and more preferably at least about 90% by weight, of the individual catalyst particles do not vary by more than about 10% from the mean density of all catalyst particles; and more preferably the individual catalyst particles do not vary by more than about 5% from the mean density of all of the particles. In a preferred embodiment of the present invention the catalyst (i.e. fresh catalyst) has a particle density ranging from about 0.6 g/cc to about 1.5 g/cc, more preferably from about 0.7 g/cc to about 1.2 g/cc, most preferably from about 0.8 g/cc to about 1.1 g/cc. After the catalyst has at least been partially spent, the particle density would range from about 0.6 g/cc to about 3.0 g/cc, more preferably from about 0.7 g/cc to about 3.0 g/cc and most preferably from about 0.8 g/cc to about 3.0 g/cc. The particle size determination will remain substantially the same as defined above. Fines and attrition may increase during hydroprocessing.  
     [0136] While the catalyst of the present invention may be any catalyst as defined above, we have discovered that the more preferred catalyst for optimally accomplishing the objectives of the present invention comprises in combination the following properties: (i) a porous inorganic oxide support; (ii) one or more catalytic metals and/or additional catalytic additives deposited in and/or on the porous inorganic oxide support; (iii) a crush strength at least about 5 pounds force; (iv) a uniform size ranging from about 6 to about 8 Tyler mesh sizes; (v) a fines content up to about 1.0 percent by weight through 8 Tyler mesh and up to about 0.2 percent by weight through 10 Tyler mesh; (vi) an attrition up to about 0.5 percent by weight through 8 Tyler mesh and up to of about 0.2 percent by weight through 10 Tyler mesh; (vii) a generally uniform spherical shape; and (viii) a uniform density ranging from about 0.7 g/cc to about 3.0 g/cc. We have discovered unexpectedly that the more preferred catalyst having or containing the immediate foregoing combination of properties, unexpectedly produces in an optimal fashion the plug-flowing substantially packed bed (i.e. catalytic bed  11 ) of hydroprocessing catalyst which is simultaneously expanding to less than 10 percent by length (more preferably less than 1% by length) beyond a substantially full axial length of the substantially packed bed of hydroprocessing catalyst in a packed bed state while (and simultaneously with) the substantially packed bed of hydroprocessing catalyst maximally and optimally occupying from about 50% percent by volume to about 98% percent by volume (i.e. the entire internal and/or inside available volume or reactor volume) of the reactor vessel  11 .  
     [0137] The particular type of porous base material or inorganic oxide support, the particular type of catalytic metal, the pore structure, the catalyst surface area and catalyst size, would all depend on the intended specific application (e.g. demetallation, desulfurization, denitrification, etc.) of the catalyst. Generally, the more preferred catalyst comprises a porous inorganic oxide support selected from the group consisting alumina, silica, and mixtures thereof, and has a surface area ranging from about 75 square meters per gram to about 150 square meters per gram. The preferred catalyst comprises catalytic metal(s), present as oxide(s) deposited in and/or on the porous inorganic support. Oxide(s) of the catalytic metal(s), or the metallic oxide component of the preferred catalyst, is selected from the group consisting of molybdenum oxide, cobalt oxide, nickel oxide, tungsten oxide, and mixtures thereof, and comprises from about 0.5 to about 50 percent by weight, more preferably from about 0.5 to about 30 percent by weight, of the total catalyst (i.e. inorganic oxide support plus metal oxide(s)). The more preferred catalyst further comprises a general uniform spherical shape having a mean diameter ranging from about 20 Tyler mesh to about 4 Tyler mesh. While a spherical shaped catalyst is the more preferred catalyst, an extrudate may be employed if it is very strong, say having a crush strength over 5 lbs. of force. The absolute size of the catalyst may vary from application to application, but the more preferred catalyst has the narrow size distribution as previously stated above.  
     [0138] From the foregoing discussion it will be clear to the skilled practitioner that, though the catalyst particles of the present process have a uniform size, shape, and density, the chemical and metallurgical nature of the catalyst may change, depending on processing objectives and process conditions selected. For example, a catalyst selected for a demetallation application with minimum hydrocracking desired, could be quite different in nature from a catalyst selected if maximum hydrodesulfurization and/or hydrodenitrification and/or hydrocracking are the processing objectives. The type of catalyst selected in accordance with and having the properties mentioned above, is disposed in any hydroconversion reaction zone. A hydrocarbon feed stream is passed through the catalyst, preferably passed through such as upflow through the catalyst, in order to hydroprocess the hydrocarbon feed stream. More preferably, the catalyst is employed with the various embodiments of the present invention.  
     EXAMPLES  
     [0139] The following examples are exemplary examples of process runs, conducted in accordance with various method steps of the present invention and employing the apparatus in accordance with various preferred embodiments of the present invention. The following set-forth examples are given by way of illustration only and not by any limitation, and set-forth a best mode of the invention as presently contemplated. All parameters such as concentrations, flow rates, mixing proportions, temperatures, pressure, rates, compounds, etc., submitted in these examples are not to be construed to unduly limit the scope of the invention.  
     Example I  
     [0140] In a semi-commercial scale residuum conversion pilot plant operating at 100-200 BPD, the catalyst transfer procedure, as described above, was demonstrated more than 50 times. During each transfer, about 2 cubic feet of catalyst was moved into and out of the reactor vessel running continuously at typical residual desulfurization (RDS) conditions. Transfer rates up to 16 cubic feet per hour of catalyst were accomplished through pipes with an inside diameter 8 times larger than the catalyst diameter. Plug flow movement of the catalyst and the absence of bed ebullation were proven using radioactively tagged catalyst particles incorporated in the test bed.  
     [0141] Among the significant features of the invention specifically demonstrated in such runs were that: (1) ball valves, such as those made by the Kaymr and Mogas companies, can be used to isolate the RDS reactor from the catalyst transfer vessels, and to transfer catalyst particles without using solids handling valves, (2) the catalyst bed level and thus ebullation can be adequately monitored using a gamma-ray source and detector, (3) J-tubes (all with upward flow sections substantially shorter than the downward flow paths) will satisfactorily transfer catalyst particles, without local ebullation, by laminar fluid flow, (4) use of fluid feed inlet distributor means with a conical support and concentric annular segment plates prevents ebullation at the base of the catalyst bed and provides adequate radial distribution of gas and liquid, by forming concentric alternate rings of gas and liquid (5) substantial differences were shown where bed ebullation (expansion) occurred with one catalyst as compared to no significant bed ebullation (expansion) with another catalyst using the same size and shape but with a lower density, and (6) transfer of catalyst into , and out of, a bed traveling downwardly by gravity in a reactor vessel while continuously operating a hydroprocessing system to react a gas containing hydrogen and a feed stream of hydrocarbon liquids flowing as a single stream from the bottom of the bed, will permit countercurrent flow without separation during upward passage through and out of the top of the reactor vessel, and (7) as the result of intermittent catalyst discharge, the catalyst bed moves countercurrently down through the reactor in plug-like flow.  
     Example II  
     [0142] From the foregoing tests in an apparatus under flow conditions described in Example I, the effectiveness of the foregoing mechanical and hydraulic factors were validated for performing hydrotreating processing with hydrocarbon and hydrogen streams counterflowing through a moving bed of catalyst particles, as follows:  
     [0143] In a pilot plant operating at up to 4 BPD hydrocarbon feed and hydrogen at 2200 PSI, catalyst bed expansion measurements were made at commercial-scale flow velocities with beds of catalysts of different sizes, shapes, and densities as indicated in Table I. Each type of catalyst was tested separately. Bed ebullation (expansion) was measured using a gamma-ray source and detector means mounted to detect 10% bed expansion. Table I shows flow velocities required to produce 10% bed expansion with several catalysts at a standard hydrogen recirculation rate of 5000 SCFB. These results confirmed the bed expansion results from the semi-commercial scale plant of Example I.  
     [0144] Table II is a similar set of runs using beds of three of the same catalyst particles as those tested under conditions shown in Table I except that the liquid viscosity, liquid density and pressure of the hydrocarbon feed stream and gas were lower in Table II than Table I to match a different set of commercial operating conditions. From Tables I and II the effect of catalyst particle size, density and shape are clearly indicated for different flow conditions for the liquid and gas components of the feed. The design feed rates for a hydrocarbon treating process were calculated by standard scaling procedures to indicate the values in MBPD (thousands of barrels per day) through a reactor vessel containing a catalyst bed 11.5 feet in diameter.  
     [0145] In general catalyst for commercial use would be selected on the basis of levitation or ebullation at a selected rate which is substantially higher than normal design feed rate, say up to 100% greater. Additionally, these tests indicate that some commercial catalysts will not lift at reasonable design feed rates if the particles have a high degree of uniformity and are sufficiently strong to maintain their integrity during movement into and out of the reactor vessel, without attrition or breakage.  
               TABLE I                          CATALYST BED EXPANSION STUDY TEST RESULTS       2200 PSI Hydrogen and Flush Oil       Liquid Density 51 lb/ft 3 Viscosity 1.1 cp       Gas Density 0.49 lb/ft 3 Viscosity 0.016 cp                             Flow rates for 10% Bed Expansion               @ 5000 SCFB H 2                                                                           Effective   Liquid   Gas               Relative       Skeletal   Particle   Density in Oil   Velocity   Velocity   MBPD in 11.5       Catalyst   Size   Shape   Density   Density   (1)   Ft/Min   Ft/Sec   Ft. ID Reactor                                                         A   1   Cylinder   2.69   1.05   0.55   0.46   0.11   13       B   1.6   Quadralobe   3.55   1.03   0.56   0.60   0.14   17       C   2   Cylinder   3.61   1.60   1.05   0.46   0.11   13       D   3.2   Sphere   2.33   0.60   0.21   0.32   0.07   9       E   3.2   Sphere   3.63   0.83   0.47   1.38   0.33   40       F   3.2   Cylinder   3.58   1.37   0.89   1.38   0.33   40                          
 
     [0146]               TABLE II                          CATALYST BED EXPANSION STUDY TEST RESULTS       With Hydrogen and Hydrocarbon at 1000 PSI       Liquid Density 48 lb/ft 3 Viscosity 0.56 cp       Gas Density 0.23 lb/ft 3 Viscosity 0.013 cp                             Flow rates for 10% Bed Expansion               @ 5000 SCFB H 2                                                                       Effective   Liquid                   Relative       Skeletal   Particle   Density   Velocity   Gas Velocity   MBPD in 11.5       Catalyst   Size   Shape   Density   Density   in Oil (1)   Ft/Min   Ft/Sec   Ft. ID Reactor                                                         C   2   Cylinder   3.61   1.60       0.53   0.13   15       E   3.2   Sphere   3.63   0.83       1.38   0.33   40       F   3.2   Cylinder   3.58   1.37       1.50   0.50   60                    
     Example III  
     [0147] In a 4 foot diameter vessel a “cold model” was operated using up to 8000 BPD water and 275 SCFM air. The features of the inlet liquid and gas distributor as well as the hydrogen gas redistribution and quench stages, described above and shown in the drawings were scaled and tested. Flow measurements and underwater photography proved that distribution of the inlet gas and liquid was uniform across the full cross-sectional area of the catalyst support screen in the vessel. Redistribution of the rising gas through the inverted V-shaped sheds was shown to be surprisingly effective even when gas was intentionally maldistributed below the redistributor stages.  
     Summary of Test Results for Examples I,II and III  
     [0148] Briefly, these test results show that the present invention makes possible substantially continuous flow of uniformly distributed hydrogen and hydrocarbon liquid across a densely packed catalyst bed to fill substantially the entire volume of a reactor vessel by introducing the fluids as alternate annular rings of gas and liquid at a rate insufficient to levitate the bed and with the catalyst selected with a density, shape and size adequate to prevent lifting of the bed at the desired feed rates. (Catalysts are selected by measuring bed expansion in a large pilot plant run with hydrocarbon, hydrogen, and catalyst at the design pressures and flow velocities). At the desired flow rate, such catalyst continually flows in a plug-like manner downwardly through the vessel by introducing fresh catalyst at the top of the bed by laminarly flowing such catalyst in a liquid stream on a periodic or semicontinuous basis. Catalyst is removed by laminarly flowing catalyst particles in a liquid stream out of the bottom of the bed. Intake for such flow is out of direct contact with the stream of gas flowing through the bed and the flow path is substantially constant in cross-sectional area and greater in diameter by several times than the diameter of the catalyst particles.  
     Example IV  
     [0149] A plurality of catalytic particulates were charged into a reaction zone contained within a reactor, such as reactor vessel  11 . The plurality of catalytic particulates formed a catalyst bed (such as catalyst bed  10  in FIGS. 1, 8 and  9 ). The catalyst bed was supported in the reactor by a truncated conical bed support similar to the support that is generally illustrated as  17  in FIGS. 8, 9 and  11 - 13 . An inlet distributor, such as circular plate member  31  in FIGS. 1 and 11 with the multiplicity of tubes  32 , extended across a full cross-sectional area of the reactor underneath the truncated conical bed support to form a plenum or inlet chamber between the inlet distributor and the truncated conical bed support, as generally illustrated in FIGS. 8, 9 and  11 - 13 . The truncated conical bed support for the catalyst bed included a series of annular polygons that included a plurality of segmented plates (such as segmented plates  27  in FIGS. 2 and 3) connected to or formed with radial spoke members such as members  26  FIGS.  10 - 13 . The plurality of segmented plates, each having a thickness of about 10 inches and a width of about 1.5 inch, were secured to 8 radial spoke members. The interengaged segmented plates and radial spoke members formed a web-like structure to produce essentially annularly continuous mixture zones for receiving a flow of hydrocarbon feed stream, and were overlayed with a screen having screen openings with a mean diameter that was smaller than the catalytic particulates. Each mixture zone underneath the screen had a generally circumferentially uniform thickness.  
     [0150] The catalytic particulates comprised an alumina porous carrier material or alumina inorganic support. Deposited on and/or in the alumina porous carrier material was an oxidic hydrotreating catalyst component consisting of NiO and/or MoO 3  and/or P 2 O 5 . The Mo was present on and/or in the alumina porous carrier material in an amount of about 3% by wt., based on the combined weight of the alumina porous carrier material and the oxidic hydrotreating catalyst component(s). The Ni was present on and/or in the alumina porous carrier material in an amount of about 1% by wt., based on the combined weight of the alumina porous carrier material and the oxidic hydrotreating catalyst component(s). The surface area of the catalytic particulates was about 120 sq. meters per gram.  
     [0151] The plurality of catalytic particulates were generally spherical with a mean diameter having a value ranging from about 6 Tyler mesh to about 8 Tyler mesh and an aspect ratio of about 1. The mean crush strength of the catalytic particulates was about 5 lbs. force. The metals loading capacity of the catalyst or plurality of catalytic particulates was about 0.3 grams of metal per cubic centimeter of catalytic particulate bulk volume.  
     [0152] The catalytic particulates had a size distribution such that 98.5% by weight of the catalytic particulates in the catalyst bed had a diameter ranging from R 1  to R 2  wherein: (i) R 1  had a value of about 0.093 inch (i.e. the approximate opening of an 8 mesh Tyler screen); (ii) R 2  had a value of about 0.131 inch (i.e. the approximate opening size of a 6 mesh Tyler screen); and (iii) the ratio R 2 /R 1  had a value equal to about the square root of 2.0 or about 1.414. The size distribution of the catalytic particulates was also such that a maximum of about 1.0% by weight of the catalytic particulates had a diameter less than R 1 . The catalyst further also had a size distribution such that a maximum of about 0.2% by weight of the catalytic particulates had a diameter less than R 3 , wherein R 3  was less than R 1  and the value of the ratio R 1 /R 3  was about the square root of 2.0 or about 1.414.  
     [0153] The catalytic particulates of the catalyst had a maximum attrition of about 0.5% by weight of the catalytic particulates through a diameter (i.e. a Tyler screen opening) having the value of R 1 , and a further maximum attrition of about 0.2% by weight of the catalytic particulates through a diameter (i.e. a Tyler screen opening) having the value of R 3  wherein R 3  again was less than R 1  and the value of the ratio of R 1 /R 3  was about the square root of 2.0 or about 1.414. Stated alternatively, for the catalytic particulates with the specified size range or distribution of about 6 to about 8 Tyler mesh, the specified attrition for the catalytic particulates was up to about 0.5% by weight MAX through 8 Tyler mesh and up to about 0.2% by weight MAX through 10 Tyler mesh.  
     [0154] The catalytic particulates had a maximum fines content of up to about 1.0% by wt. through 8 Tyler mesh and up to about 0.2% by wt. through 10 Tyler mesh. Stated alternatively, for the catalytic particulates with the specified size range or distribution of about 6 to about 8 Tyler mesh, the specified fines content for the catalytic particulates was up to about 1.0% wt. fines MAX through 8 Tyler mesh and up to about 0.2% by wt. fines MAX through 10 Tyler mesh. The catalytic particulates had a uniform density such that mean density of the catalytic particulates were about 0.9 g/cc.  
     [0155] The liquid component of the hydrocarbon feed stream was a heavy atmospheric residuum wherein at least 95% by volume of which boiled above about 343° C. and wherein a substantial fraction (e.g. 50% by volume) boiled above about 510° C. The “heavy” hydrocarbon feed had an undesirable metal content of about 90 ppm by weight of the “heavy” hydrocarbon feed. The hydrogen-containing gas of the hydrocarbon feed stream was essentially 97% pure hydrogen and was mixed with the heavy atmospheric residuum stream in a mixing ratio of 623 liters of hydrogen-containing gas at standard conditions per liter of heavy atmospheric residuum in order to form the hydrocarbon feed stream.  
     [0156] The hydrocarbon feed stream was passed through the inlet distributor and introduced into the plenum chamber of reactor at a flow rate ranging from about 0.1 ft/sec. to about 1.00 ft/sec. The hydroprocessing pressure and temperature within the reactor were about 2300 psig. and about 400° C. respectively. From the plenum chamber of the reactor the hydrocarbon feed stream entered into the annular continuous mixture zones and was uniformly fed through the screen and into the catalyst bed such as not to induce local ebullation or eddy currents in the catalyst bed, especially in proximity to the conical bed support which was overlayed with the screen.  
     [0157] The catalyst bed in the reactor contained a plurality of axially spaced apart hydrogen gas redistribution (or hydrogen gas-quenching) assemblies (see FIGS. 5 and 7 as illustrative of the hydrogen gas-quenching assemblies). As the hydrocarbon feed stream flowed upwardly through the catalyst bed, hydrogen gas was emitted from the hydrogen gas redistribution assemblies, which redistributed any hydrogen-containing gas that had become channeled in a portion of the catalyst bed below (or in close proximity to) the hydrogen gas redistribution assemblies and further avoided generation of local hot spots, eddy currents or ebullation in the upper part (especially above the hydrogen gas redistribution assemblies) of the catalyst bed.  
     [0158] The liquid hydrocarbon feed stream exited the reactor at a withdrawal flow rate of about 3.6 ft/sec. and had been upgraded such that it contained a metal content of about 3 ppm by wt. of the liquid hydrocarbon feed stream. As the hydrocarbon feed stream flowed upwardly through the catalyst bed, a gamma ray source in the catalyst bed in combination with a gamma ray detector on the reactor (see for example gamma ray source  22  in the catalyst bed  10  competing with the gamma ray detector  24  on the reactor vessel  10  in FIG. 1) detected that the catalyst bed expanded less than 10% by length over or beyond substantially the full axial length of the catalyst bed in a packed bed state.  
     [0159] After the reactor was on stream for about 1 weeks, approximately 7.25 cubic meters (or about 3.3% by weight of the catalyst bed) of catalytic particulates were laminarly withdrawn in the hydrocarbon feed stream through a J-tube (such as J-tube  29  in FIG. 1) at a flow rate of about 3.6 ft/sec. The withdrawn catalyst in the hydrocarbon feed stream had a concentration of about 0.5 lbs. catalyst/lb. catalyst slurry (i.e. weight of withdrawn catalyst plus weight of hydrocarbon feed stream). When and/or as the volume of catalytic particulates were withdrawn or transferred from the bottom of the catalyst bed, the catalyst bed (i.e. a substantially packed bed of catalyst) began to plug-flow. The withdrawn catalyst was replaced by introducing a comparable volume of fresh replacement catalyst through the top of the reactor. The fresh replacement catalyst was slurried in a hydrocarbon refined stream (e.g. gas oil) and was introduced into the reactor at a flow catalyst replacement rate of about 3.6 ft/sec., and at a catalyst replacement concentration of about 0.5 lbs. replacement catalyst/lb. catalyst slurry (i.e. weight of replacement catalyst plus the hydrocarbon refined stream (e.g. gas oil) as the slurrying medium).  
     [0160] While the present invention has been described herein with reference to particular embodiments thereof, a latitude of modification, various changes and substitutions are intended in the foregoing disclosure, and it will be appreciated that in some instances some features of the invention will be employed without a corresponding use of other features without departing from the scope of the invention as set forth.