Abstract:
A method increases the feed throughput for a process of dehydrogenating light hydrocarbons without loss of conversion or selectivity by increasing the catalyst volume in only the final reactor of at least three reaction zones. The catalyst volume of the final reactor may be increased relative to the other reactors by extending the inner and outer screens that define a radial flow bed therein. Maintaining a low LHSV by increasing the catalyst volume in only the final reactor greatly reduces the expense of improving the capacity and yield of such a process. This method provides the most benefit to moving bed reactor systems since modifications are limited to the last reactor. A further benefit is derived from the simplified method of raising only a section of the last reactor to increase the catalyst volume.

Description:
CROSS-REFERENCE TO RELATED APPLICATION 
     This application claims priority from Provisional Application Serial No. 60/112,588 filed Dec. 17, 1998, the contents of which are hereby incorporated by reference. 
    
    
     FIELD OF THE INVENTION 
     The field of art to which this invention pertains is the dehydrogenation of hydrocarbons in multiple reaction zones that use on stream catalyst replacement. 
     BACKGROUND OF THE INVENTION 
     The dehydrogenation of hydrocarbons is an important commercial hydrocarbon conversion process because of the existing and growing demand for dehydrogenated hydrocarbons for the manufacture of various chemical products such as detergents, high octane gasolines, oxygenated gasoline blending components, pharmaceutical products, plastics, synthetic rubbers, and other products which are well known to those skilled in the art. One example of this process is the dehydrogenation of isobutane to produce isobutylene which can be polymerized to provide tackifying agents for adhesives, viscosity-index additives for motor oils, and impact-resistant and antioxidant additives for plastics. Another example of the growing demand for isobutylene is the production of oxygen-containing gasoline blending components which are being mandated by the government in order to reduce air pollution from automotive emissions. 
     Those skilled in the art of hydrocarbon conversion processing are well versed in the production of olefins by means of catalytic dehydrogenation of paraffinic hydrocarbons. In addition, many patents have issued which teach and discuss the dehydrogenation of hydrocarbons in general. For example, U.S. Pat. No. 4,430,517 (Imai et al) discusses a dehydrogenation process and catalyst for use therein. 
     Most catalysts for the dehydrogenation of hydrocarbons are susceptible to deactivation over time. Deactivation will typically occur because of an accumulation of deposits that block active pore sites or catalytic sites on the catalyst surface. Where the accumulation of coke deposits causes the deactivation, reconditioning the catalyst to remove coke deposits restores the activity of the catalyst. Coke is normally removed from the catalyst by contact of the coke-containing catalyst with an oxygen-containing gas at a high enough temperature to combust or remove the coke in a regeneration process. In a moving bed process, the regeneration process is carried out by removing catalyst from the vessel in which the hydrocarbon conversion is taking place and transporting the catalyst to a separate regeneration zone for coke removal. Arrangements for continuously or semi-continuously removing catalyst particles from a bed in a reaction zone for coke removal in a regeneration zone are well known. U.S. Pat. No. 3,652,231 describes a continuous catalyst regeneration process which is used in conjunction with the catalytic reforming of hydrocarbons, the teachings of which are hereby incorporated by reference. In the reaction zone of U.S. Pat. No. 3,652,231, the catalyst is transferred under gravity flow by removing catalyst from the bottom of the reaction zone and adding catalyst to the top while reactants flow cross currently through a radial flow bed. 
     In the past, flow-related phenomena have limited mass flow and fluid velocity through the radial flow beds. One phenomenon, known as “pinning,” inhibits catalyst transfer in many reactor arrangements. Pinning occurs when the flow of fluid at sufficient velocity blocks the downward movement of catalyst. Pinning is a function of the gas composition, the gas velocity, the physical characteristics of the catalyst, and the physical characteristics of the flow channel through which the catalyst must move. As the gas flows through the channels that retain the catalyst, the gas impacts the catalyst particles and raises intergranular friction between the particles. When the vertical component of the frictional forces between the particles overcomes the force of gravity on the particles, the particles become pinned. As the flow path length of gas through the catalyst particles becomes longer, the forces on the particles progressively increase from the outlet to the inlet of the flow channel. 
     Another flow-limiting phenomena is called “void blowing”. Void blowing occurs when the gas velocity displaces from a surface of the catalyst bed across which it flows from the screen or other retaining elements, thereby creating a void space. The rapid circulation or churning of catalyst over the free surface of the void space can abrade or break catalyst particles, resulting in the production of reduced size particles or fines. This fine material can plug the catalyst bed, exacerbate the void blowing problem, or accumulate in other process piping or equipment in a manner that interferes with the continued effective operation of the process. 
     As technology has improved and problems such as pinning and void blowing have been better understood, it has been possible to increase the fluid velocity through the radial flow beds of existing reaction zones that were previously limited by these fluid flow phenomena. Increasing the velocity or throughput is of course desirable because it permits an increase in capacity with only minor operating changes to the existing equipment. However, it has been found that further increases in the capacity of many existing dehydrogenation units cannot only be obtained with decreases in conversion or selectivity of the products produced. The limitations in conversion and/or selectivity are related to the reduced time that results from the higher fluid velocity through the catalyst beds. The catalyst contact time is typically expressed in terms of the liquid hourly space velocity (LHSV) of the feed through the catalyst bed. Higher LHSV&#39;s may be compensated for, to some extent, by an increased reaction temperature which raises the catalyst activity. Although compensating for higher throughput by increasing the reactor temperatures may maintain conversion levels, it is typically at the expense of lower selectivity and increased coke production on the catalyst. 
     The most direct way to overcome the problems of space velocity limitations is to add more catalyst to the process. Increasing the catalyst volume is readily accomplished in the design stage for a new unit. Unfortunately for existing units, adding additional catalyst could require expensive modification or replacement of all of the reactors and the associated piping for the delivery of reactants and the transfer of catalyst between the reactors. 
     Accordingly, it is an objective of this invention to increase the throughput of existing dehydrogenation reactors with only limited modifications to the reaction zone and the associated piping. 
     BRIEF SUMMARY OF THE INVENTION 
     It has now been surprisingly discovered that it is possible to increase throughput in a dehydrogenation process without sacrificing the conversion or selectivity by only increasing the catalyst volume in the last reaction zone and that it is possible to increase the catalyst volume of the last reaction zone in a simple and relatively inexpensive manner. Adding catalyst to the last reaction zone effectively decreases the overall LHSV for the feed stream at higher throughput rates. As a result, conversion is maintained without increasing the temperature in any of the individual reactors and thereby avoiding any fall off in selectivity or increase in coke production. Achieving these results by only requiring a simple modification of the last reaction zone makes the operation of such dehydrogenation processes at higher throughput highly cost effective. The method applies to multiple reactors wherein the last reactor is in a side-by-side arrangement to the upstream reactors. Furthermore, the modification of the last reaction zone may be simplified by the arrangement of many reactors which often include a flange connection at a location that may be beneficially used to extend the catalyst volume of the radial flow bed. 
     One embodiment of the present invention may be characterized as a method of increasing the throughput of a process for the dehydrogenation of hydrocarbons. The process passes a combined dehydrogenation feed into contact with a catalyst bed in a first reactor, passes the effluent from the first reactor into contact with a catalyst bed in at least one intermediate reactor, and passes an intermediate effluent from the last of the at least one intermediate reactor into contact with a final catalyst bed in a final reactor. The process effects catalyst circulation in at least the last reactor by adding catalyst to the top of the final catalyst bed and withdraws catalyst from the bottom of the final catalyst bed in a final reactor while contacting the intermediate effluent in the final catalyst bed of the final reactor. The final reactor is located to the side of the intermediate reactor from which it receives the intermediate effluent and the catalyst bed of the final reactor retains catalyst between an inner and outer screen in an arrangement for the radial flow of reactants. The method includes splitting the vessel wall of the last reactor at a disconnection point above the middle of the inner and outer screens to define a lower reactor section that retains the inner and outer screens and a disconnected upper reactor section that retains equipment for supplying catalyst to the final catalyst bed. The disconnected upper section is at least temporarily removed. The inner and outer screens are extended by a distance that will increase the catalyst volume in the process by at least 5%. A new section of vessel wall installed above the disconnection point encloses the extended inner and outer screens with a closing upper section of the final reactor that retains equipment for supplying catalyst to the final catalyst bed. 
     Extending the inner and outer screens increases the volume of the catalyst with a corresponding decrease in the LHSV of the process. The catalyst volume will usually be increased by an amount that will keep the LHSV of the process below about 4 hrs −1 . Increasing the catalyst volume in the last reaction will also result in the local LHSV of the first reactor and any intermediate reactor being greater than the LHSV of the final reactor. The LHSV of the final reactor, once the screens have been extended, will usually be at least 5 hrs −1  less than the LHSV of the other reactors. Of course, any decrease in the LHSV will be limited by the amount of catalyst that can be added to the modified reactor. Preferably, the final reactor will increase the catalyst volume by 10% and, more preferably, the catalyst volume will increase by 15% or more. Structural limitations such as foundation support may control the amount of catalyst that may be added to the final reaction zone by extension of the inner and outer screens. 
     The splitting of the last reactor may take place at any convenient disconnection point. A disconnection point at or above the inner and outer screens is preferred in order to minimize the disturbance of existing screen supports and related instrumentation or other miscellaneous piping that may be connected to the reactor vessel. The upper section of the reactor that is removed for access to the screens may be reused or replaced with a new reactor section that contains an additional section of external vessel wall. Preferably, the upper reaction section will be reused along with the equipment for the delivery of catalyst that is typically included therewith. 
     It is particularly convenient to reuse the upper section when the existing reactor has an upper section that is flanged to the lower section. In accordance with such an arrangement, this invention comprises a method of increasing the capacity of a process for the dehydrogenation of hydrocarbons. The process passes a combined dehydrogenation feed into contact with a first catalyst bed in a first reactor at a first LHSV through the first catalyst bed, passes the effluent from the first reactor into contact with an intermediate catalyst bed in at least one intermediate reactor at an intermediate LHSV through the intermediate catalyst bed, and passes an intermediate effluent from the last of the at least one intermediate reactor into contact with a final catalyst bed in a final reactor at a final LHSV. The process adds catalyst to the top of the final catalyst bed and withdraws catalyst from the bottom of the final catalyst bed in the final reactor while contacting the intermediate effluent in the final catalyst bed of the final reactor. The final reactor is located to the side of the intermediate reactor from which it receives the intermediate effluent and the catalyst bed of the final reactor retains catalyst between an inner and outer screen in an arrangement for the radial flow of reactants. The method splits the vessel wall of the last reactor above the top of the inner and outer screens to create an upper reactor section that retains catalyst supply equipment for supplying catalyst to the final reactor bed and a lower reactor section that retains the inner and outer screens. The upper section is temporarily removed and the inner and outer screens are extended by a distance that will retain sufficient catalyst to reduce the LHSV of the last reactor below the lowest LHSV of the first reactor and all of the intermediate reactors. The method also inserts an intermediate section of vessel wall between the upper and lower reactor sections in a sealed arrangement to enclose the extended inner and outer screens and to position the catalyst supply equipment for delivery of catalyst to the final reactor bed. The split is preferably made across a flanged connection and the new section of vessel wall has flanged ends for attachment to the existing flanges of the upper and lower reactor sections. 
     Other embodiments of the present invention encompass further details of preferred process conditions and reactor modification methods. 
    
    
     BRIEF DESCRIPTION OF THE DRAWINGS 
     FIG. 1 is a schematic drawing of a series of dehydrogenation reactors. 
     FIG. 2 is a schematic drawing of a single dehydrogenation reactor. 
     FIG. 3 is a schematic drawing showing a modified dehydrogenation reactor. 
    
    
     DETAILED DESCRIPTION OF THE INVENTION 
     The dehydrogenation of paraffinic hydrocarbons is well known to those skilled in the art of hydrocarbon processing. In the dehydrogenation process, fresh hydrocarbon feed is combined with recycle hydrogen and unconverted hydrocarbons. Dehydrogenatable hydrocarbons for this invention will preferably include isoalkanes having 3 or 5 carbon atoms. A suitable feed of dehydrogenatable hydrocarbons will often contain light hydrocarbons (i.e., those having less carbon atoms than the primary feed components) which, for the purpose of this invention, serve as contaminants. In most cases, olefins are excluded from the dehydrogenation zone recycle in order to avoid the formation of dienes which produce unwanted by-products in many of the olefin conversion processes. Along with the dehydrogenatable hydrocarbons, the feed to the dehydrogenation zone of the present invention comprises an H 2 -rich stream, preferably containing at least 75 mol-% H 2 . The H 2  acts to suppress the formation of hydrocarbonaceous deposits on the surface of the catalyst, more typically known as coke, and can act to suppress undesirable thermal cracking. Because H 2  is generated in the dehydrogenation reaction and comprises a portion of the effluent, the H 2 -rich stream introduced into the reaction zone generally comprises recycle H 2  derived from separation of the dehydrogenation zone effluent. Alternately, the H 2  may be supplied from suitable sources other than the dehydrogenation zone effluent. 
     The combined stream of hydrogen and hydrocarbons is passed through a suitable bed of dehydrogenation catalyst maintained at the proper dehydrogenation conditions such as temperature, pressure and space velocity, and the effluent from the catalytic reaction zone is processed further to yield a stream of olefinic hydrocarbons. The dehydrogenation reactors of this invention preferably comprise at least a last radial flow reactor through which the catalytic composite gravitates downwardly to allow a substantially continuous replacement of the catalyst with fresh and/or regenerated catalyst. A detailed description of the moving bed reactors herein contemplated may be obtained by reference to U.S. Pat. No. 3,978,150. The dehydrogenation reaction is a highly endothermic reaction which is typically effected at low (near atmospheric) pressure conditions. The precise dehydrogenation temperature and pressure employed in the dehydrogenation reaction zone will depend on a variety of factors such as the composition of the paraffinic hydrocarbon feedstock, the activity of the selected catalyst, and the hydrocarbon conversion rate. In general, dehydrogenation conditions include a pressure of from about 0 to about 35 bars and a temperature of from about 480° C. (900° F.) to about 760° C. (1400° F.). A suitable hydrocarbon feedstock is charged to the reaction zone and contacted with the catalyst contained therein at an LHSV of from about 1 to about 10. Hydrogen, principally recycle hydrogen, is suitably admixed with the hydrocarbon feedstock in a mole ratio of from about 0. 1 to about 10. Preferred dehydrogenation conditions, particularly with respect to C 4 -C 5  paraffinic hydrocarbon feedstocks, include a pressure of from about 0 to about 5 bars and a temperature of from about 540° C. (1000° F.) to about 705° C. (1300° F.), a hydrogen-to-hydrocarbon mole ratio of from about 0.1 to about 2, and, after employment of this invention, an LHSV of less than 4. 
     The dehydrogenation zone of this invention may use any suitable dehydrogenation catalyst. Generally, the preferred catalyst comprises a platinum group component, an alkali metal component, and a porous inorganic carrier material. The catalyst may also contain promoter metals which advantageously improve the performance of the catalyst. It is preferable that the porous carrier material of the dehydrogenation catalyst be an absorptive high surface area support having a surface area of about 25 to about 500 m 2 /g. The porous carrier material should be relatively refractory to the conditions utilized in the reaction zone and may be chosen from those carrier materials which have traditionally been utilized in dual function hydrocarbon conversion catalysts. A porous carrier material may, therefore, be chosen from an activated carbon, coke or charcoal, silica or silica gel, clays and silicates including those synthetically prepared and naturally occurring, which may or may not be acid-treated as, for example, attapulgus clay, diatomaceous earth, kieselguhr, bauxite; refractory inorganic oxides such as alumina, titanium dioxide, zirconium dioxides, magnesia, silica alumina, alumina boria, etc.; crystalline alumina silicates such as naturally occurring or synthetically prepared mordenite or a combination of one or more of these materials. The preferred porous carrier material is a refractory inorganic oxide, with the best results being obtained with an alumina carrier material. The aluminas, such as theta alumina, give the best results in general. The preferred catalyst will have a theta alumina carrier which is in the form of spherical particles having relatively small diameters on the order of about {fraction (1/16)} -inch. Of the platinum group metals, which include palladium, rhodium, ruthenium, osmium and iridium, the use of platinum is preferred. The platinum group component may exist within the final catalyst composite as a compound such as an oxide, sulfide, halide, oxysulfide, etc., or an elemental metal or in combination with one or more other ingredients of the catalyst. The platinum group component generally comprises from about 0.01 to about 2 wt-% of the final catalytic composite with a preferred platinum content of about 0.1 and 1 wt-%. When present, the preferred alkali metal is normally either potassium or lithium, depending on the feed hydrocarbon. The concentration of the alkali metal may range from about 0.1 to 5 wt-%. The preferred promoter metal is tin and normally constitutes about 0.01 to about 1 wt-%. It is preferred that the atomic ratio of tin to platinum be between 1:1 and about 6:1. The tin component may be incorporated into the catalytic composite in any suitable manner known to effectively disperse this component in a very uniform manner throughout the carrier material. A more detailed description of the preparation of the carrier material and the addition of the platinum component and the tin component to the carrier material may be obtained by reference to U.S. Pat. No. 3,745,112. 
     Operation of the dehydrogenation zone will produce a mixture of hydrogen and hydrocarbons. Normally, a portion of the hydrocarbons will include an equilibrium mixture of the desired isoolefin and its isoalkane precursor. The effluent from the dehydrogenation reaction section passes to a hydrogen recovery section. This separation section removes hydrogen from the effluent and recovers it in high purity for recycle to the dehydrogenation reaction section. Separation steps for the removal of hydrogen will normally include cooling and compressing with subsequent cooling and flashing in a separation vessel. Such methods for the separation of hydrogen and light gases are well known by those skilled in the art. 
     A typical dehydrogenation process passes the combined hydrocarbon and hydrogen feed through a plurality of reactors with interstage heating between the reactors. This invention is generally applicable to dehydrogenation processes having three or more reaction zones. The feed hydrocarbons and hydrogen are initially heated by indirect heat exchange with the effluent from the dehydrogenation zone. Following heating, the feed mixture normally passes through a heater to further increase the temperature of the feed components before it enters the dehydrogenation reactor where it is contacted with the dehydrogenation catalyst. The endothermic reaction reduces the temperature of the reactants which then undergo interstage heating before entering the next reactor. After heat exchange with the feed, the effluent from the last dehydrogenation zone effluent passes to product separation facilities. 
     DETAILED DESCRIPTION OF THE DRAWINGS 
     FIG. 1 shows a simplified flow path for reactants and catalyst in a dehydrogenation process having three reactors. Additional details of heat exchangers, inter-heaters, fluidization piping for catalyst transfer, instrumentation, and other equipment are omitted. Line  10  passes the combined feed to first reactor  12  that contains first catalyst bed  14 . The feed passes in radial flow through bed  14  and exits reactor  12  through line  16 . Line  16  carries the effluent from the first reactor through an interstage heater (not shown) that reheats the effluent from reactor  12  before it enters single intermediate reactor  18  that contains catalyst bed  20 . The effluent from the first reaction zone passes again in radial flow through bed  20  and an intermediate effluent stream exits reactor  18  via line  22 . Line  22  passes the feed though another interstage heater (not shown) and to final reactor  24  for radial flow through final catalyst bed  26 . Line  28  recovers the effluent from catalyst bed  26  for separation in separation facilities (not shown). 
     Looking next at the flow of catalyst, catalyst supply line  30  delivers fresh catalyst to a reduction zone that prepares the freshly regenerated catalyst for passage through reaction zones  12 ,  18 , and  24 . Periodic withdrawal of catalyst from collector  32  prepares reactor  12  for incremental transport of catalyst downwardly through reactor bed  14  via catalyst transfer lines  34  and  36 . Catalyst lifted from collector  32  passes to reactor  18  via line  38  that drops catalyst into surge hopper  40 . Surge hopper  40 , in a manner similar to that described for reactor  12 , retains catalyst in preparation for passage through bed  20  and into collector  42  via catalyst transfer lines  44  and  46 . Similar to that described in connection with reactor  18 , catalyst transfer line  48  transfers catalyst from collector  42  to surge hopper  50  for the supply of catalyst to reactor  24 . Catalyst passes through bed  26  via catalyst transfer lines  52  and  54  before passing into surge hopper  57 . Line  61  withdraws catalyst from surge hopper  57  for regeneration in regeneration facilities (not shown). 
     Additional internal details of reactor  24  are shown in FIG.  2 . Surge hopper  50  is fixed above reactor shell  59  to frusto-conical reducer  55 . A pair of flanges  56  and  58  connect frusto-conical reducer  55  to reactor shell  59 . Upper baffle  60  operates in conjunction with nozzles  62  and  64  to control the movement of catalyst through the surge hopper. Baffle  66  distributes gas entering through nozzle  62  while nozzle  64  lo controls the withdrawal of gas from a lower portion of surge hopper  50 . Catalyst enters the top of surge hopper  50  through upper nozzles  51 . Catalyst exits the bottom of the surge hopper across covered inlets  53  of catalyst transfer lines  52 . Catalyst transfer lines  52  discharge catalyst particles across cover plates  65  into annular bed  26  as catalyst is withdrawn from the bottom of bed  26  via catalyst transfer lines  54 . 
     Inner screen  67  and outer screen  68  retain annular catalyst bed  26  and at least partially define boundaries of distribution space  73  and collection space  78 . Outer screen section  68  is supported from bracket  70  through attachment to an upper portion of the vessel wall that forms reactor shell  59 . Inner screen  67  is fixed with respect to outer screen  68  for support thereof. Central distributor plug  71  occupies the central portion of reactor  24  surrounded by inner screen  67  and serves to distribute incoming reactants while minimizing the volume of distribution space  73 . Reactants enter bottom nozzle  74  through closed conduit section  76  that extend to the bottom of inner screen  67 . Conduit  76  may contain expansion elements as necessary to accommodate differential expansion between screens  67  and  68  and reactor shell  59 . Collection space  78  to the outside of screen  68  serves as a collection zone that supplies the reactor effluent for discharge through nozzle  80 . 
     FIG. 3 shows reactor  24  modified in accordance with this invention to add approximately 50% more catalyst to the annular bed of the reactor. To effect this modification, surge hopper  50  was removed by separating flanges  58  and  56 . Surge hopper  50  and reducer  55  were removed together along with transfer lines  52 . 
     With surge hopper  50  and its associated equipment removed, temporarily, the cover plate assemblies at the top of existing screens  67  and  68  were removed to permit extension of the screens by the addition of inner screen extension  67 ′ and outer screen extension  68 ′. Screen extensions  67 ′ and  68 ′ contain cover plate arrangement  65 ′ at the top sections of the screens. Extended distributor plug  71 ′ was inserted into the central space bounded by screen extensions  67 ′ and inner screen  67 . Distributor plug  71 ′ as well as the inner screens and outer screens are supported by support bracket  70 ′ fixed to the wall of vessel extension section  82 . Existing support brackets  70  undergo minor modification to now serve as additional guide brackets. 
     The existing inner screen and outer screen sections may be extended by welding new screen sections of the same diameter to the top of the existing screens. Welding can be facilitated through the use of backing bars or other solid members that may be added to the top of the screens or available as part of the cover plate assembly. Stiffeners and other longitudinal supports may be carried through from the existing screens to the new screens as necessary. 
     Vessel extension  82  has lower flange  83  and upper flange  85 . Flange  58  receives flange  83  for a simple sealed connection of additional vessel section  82 . Flange  85  in turn receives flange  56  of existing hopper  50  and reducer  55  to provide a completed enclosure for screen extensions  67 ′ and  68 ′. As surge hopper  50  is put back in place, catalyst transfer lines  52  are received by new cover plate arrangement  65 ′. 
     The new surge hopper section and the extended screens were provided without disturbing the lower section of reactor  24  in any manner. Only minor modifications were necessary at the top of the reactor to provide additional guides that are useful in controlling the placement and location of the screens within reactor  24 ′. As a result, the additional throughput for the unit is achieved with a relatively quick modification using components that are easily fabricated prior to any field modification. Prior fabrication of the modified internals has a further advantage of minimizing downtime for the process when making changes to the reactor. 
     EXAMPLE 
     A three-reactor dehydrogenation process was evaluated using a simulation of the dehydrogenation process to verify the advantages and performance from the modification of a single third reactor to retain a higher catalyst volume. The simulation was first performed to evaluate the three-reactor system for processing  41 ,420 barrels per day of a dehydrogenation feed through the three reaction zones at an LHSV of 4.8 hrs −1 . The process used a hydrogen-to-hydrocarbon ratio of 0.4 and the predicted operation was based on the performance of a platinum aluminum type catalyst containing a potassium modified and a tin promoter. The inlet temperatures through the three reaction zones were 632°, 642°, and 643° C., respectively. The base case showed a conversion of 48.3 wt-% with a selectivity of 83.5 wt-%. 
     In following the method of this invention, the catalyst volume of the third reactor was increased by 50% through an increase in screen length of from 40 feet to 60 feet. The overall increase in the catalyst volume was approximately 19%. The overall LHSV with the modified reactor dropped to 4.0 hrs −1 . The remaining operating parameters remained the same as for the base case. Evaluation of the performance showed an increase in conversion to 50.4 wt-% and an increase in selectivity to 83.8 wt- %.