Avinash GuptaChevron Lummus GlobalBloomfield, New Jersey
INTRODUCTION
Ongoing trends in the petroleum refining industry have resulted in the need to upgrade bot-tom-of-the-barrel heavy oils that otherwise are difficult to transport and market due to theirhigh viscosity and high levels of contaminants, such as sulfur, metals, asphaltenes, carbonresidues, and solid particles. Petroleum refiners find it necessary to process heavier crudesthat require deep residual conversions to produce clean, high-quality finished products.The LC-Fining residual hydroconversion process was developed to specifically targethydrocracking the world’s most difficult, heavy, lower-value hydrocarbon streams (petro-leum residuals, heavy oils from tar sands, shale oils, solvent-refined coal extracts, etc.) atconversion levels of 80 percent and higher.
Increasing demand for light and middle distillates, as well as changing environmentalregulations and specifications for fuel oil production, has further increased the need formore efficient residuum upgrading processes. The LC-Fining process, when coupled withan integrated, fixed-bed, wide-cut hydrotreater/hydrocracker, produces high-quality fin-ished products without significant quantities of undesirable by-products.
Earlier heavy vacuum residual technologies (carbon rejection or hydrogen additiontype) were generally limited to distillate yields of 40 to 60 percent. The remaining uncon-verted bottoms were used as coke, low-BTU gas, or residual fuel oil.
A major process route for coping with these challenges is residue hydrocracking. Thisprocess is characterized by both thermal cracking and hydrogenation reactions wherebythe heavy, hydrogen-deficient components in the feed are converted to lighter products.The LC-Fining process is a commercially proven hydrocracking process for the upgradingof residues.
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Residual upgrading process requirements should include the ability to (1) handle highheats of reaction without wasting reactor volume, (2) handle extraneous material withoutplugging, (3) provide uniform distribution of reactants and efficient contacting, and (4)operate over extended periods without shutdown. The nature of the LC-Fining processmakes it ideally suited for the conversion of residues to lighter, more valuable products.The process can be tailored for the feedstocks, the degree of conversion, and the productqualities required, especially the production of high-quality residual fuel oil with low sul-fur content and good pipeline stability, or high-quality synthetic crude oils.
The LC-Fining process is based on technology initially developed and commerciallydemonstrated by Cities Service, and subsequently improved and refined by ABB LummusGlobal, BP (formerly Amoco Oil Company), and ChevronTexaco Corp. (formerlyChevron). With this process, heavy oil feeds—including gas oils, petroleum atmosphericand vacuum residue, coal liquids, asphalt, bitumen from tar sands, and shale oil—arehydrogenated and converted to a wide spectrum of lighter, more valuable products such asnaphtha, light and middle distillates, and atmospheric and vacuum gas oils. Residual prod-ucts can be used as fuel oil, synthetic crude, or feedstock for a coker, visbreaker, solventdeasphalter, or residual catalytic cracker. Operating conditions and catalyst type and activ-ity can be varied to achieve the desired conversion, Conradson carbon reduction (CCR),desulfurization, and demetallization of residual oil feeds.
DEVELOPMENT AND COMMERCIAL HISTORY
From 1957 to 1975, Cities Service Research and Development Company participated withHydrocarbon Research Institute (HRI) to pilot-test and develop an ebullated bed hydro-conversion process (H-Oil). During this period, research and development programs werecontinually carried out in several pilot units. Based on the pilot tests, the first commercialunit was designed and operated at Lake Charles, Louisiana, in 1963.
In 1975, ABB Lummus Global (Lummus) joined together with Cities Service tolicense, market, design, and generally improve upon the technology from Cities Service.Pilot-plant facilities for this technology, called LC-Fining, were built in New Jersey.Lummus carried out comprehensive process pilot-plant studies and mechanical designdevelopments and offered initial operation and process simulation services.
The first license was sold to Amoco Oil Company in 1981. Amoco operated the com-mercial plant in Texas City, conducted extensive pilot-plant and catalyst developmentwork, and eventually became a joint licensor with Lummus in 1984. The Amoco unitstarted up quickly and performed well right from the beginning: All design targets weremet or exceeded. To maximize profits, throughout its operation, the Amoco unitprocessed the optimum-priced crudes available based on the characteristics of the resid-ual bottoms to be processed in the LC-Fining unit (i.e., high sulfur/high metals contentfeedstocks, including blends with over 40 percent Mexican Maya). Amoco installed itsown LC-Fining pilot-plant facilities in 1980. Lummus was given access to much of theinformation from these pilot units as well as that from Syncrude Canada’s pilot unit,which operated from 1988 to 1998.
Over many years, Lummus developed a large pilot-plant and commercial units data-base on various residual feeds from different geographic locations with a wide range ofcontaminant levels (metals, sulfur, nitrogen, CCR, asphaltenes, etc.). Some of theseresidual feeds included the world’s most difficult, very heavy, lower-value hydrocarbonstreams.
Lummus’s pilot units were used to conduct many programs for design data develop-ment for various potential clients and the U.S. government. At one time, there were three
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EBULLATED BED BOTTOM-OF-THE-BARREL HYDROCONVERSION 8.69
ebullated bed pilot plants operating continuously in addition to mini-stirred autoclave testfacilities. Many of the Middle Eastern, Mexican, Venezuelan, western Canadian, south-eastern Asia, Russian, and U.S. reduced crudes and vacuum bottoms, solvent-refined coalextracts, heavy oils from tar sands, and shale oils were processed in Lummus’s pilot-plantfacilities.
When the H-Oil process was developed, the initial goal was to process Athabasca bitu-men for Cities Service Athabasca (now Syncrude Canada). Many of the pilot programsconducted from 1957 to 1972 were devoted to Athabasca bitumen. In the 1970s and 1980sLummus also conducted numerous pilot runs on bitumen feeds for Syncrude.
In the early 1990s, an extensive pilot program was performed for Alberta Oil SandsTechnology Research Administration (AOSTRA) in order to demonstrate Lummus’s highconversion technology (LC-Fining) with Athabasca bitumen and other Alberta heavycrudes. During this work, the integrated hydrotreater was closely studied and piloted atconditions that could be applied commercially.
In 1990, a 70-day pilot-plant run was conducted on Arabian Heavy vacuum residue inLummus’s research facilities to support AGIP Petroli Raffineria di Milazzo’s (RAM) com-mercial LC-Fining unit design. In this run, a series of test programs were conducted atvarying operating conditions and using various feed diluents and cutter stocks. The pri-mary objective of the run was to establish the reactor operating conditions and proper feeddiluent blends that would permit the maximum level of conversion to be attained, consis-tent with meeting RAM’s low-sulfur fuel oil requirements. The commercial unit, commis-sioned in September 1998, has been running with an on-stream time in excess of 96percent while producing 1 wt % sulfur stable fuel oil.
In 1995, Russian vacuum residual supplied by Slovnaft for its refinery in Bratislava,Slovakia, was processed in Lummus’s pilot-plant facility at conversion levels rangingfrom 60 to 88 vol %. For operations at higher conversion levels, an aromatic solvent(heavy cycle oil) was used with the vacuum residual, using high-HDS and low-sedimentactivity catalysts. This catalyst was superior to the standard LC-Fining catalyst tested inearlier runs in terms of sediment control and hydrodesulfurization (HDS) and CCRactivity. The objective of this pilot run was to establish the reaction yields and productqualities to be used in the design and guarantees for Slovnaft’s commercial LC-Finingunit. The Slovnaft plant is similar to the RAM unit and has been successfully operatingsince 2000.
In 2000, Lummus and Chevron joined forces to jointly develop and market the resid-ual upgrading technologies of both companies—including the LC-Fining process—undera single entity, Chevron Lummus Global LLC (CLG).
In 1999 and 2000, two LC-Fining pilot plant runs were conducted at CLG’s facilitiesto support Shell Canada’s and Petro-Canada’s commercial LC-Fining units designefforts. During the Shell Canada run, a close-coupled, online, integrated, wide cut dis-tillate, fixed-bed hydrotreater was also tested. In addition, off-line hydrotreating testswere performed to replicate the inhibition effects of H2S and NH3 expected in commer-cial operation.
Under a joint cooperative agreement, CLG and AGIP Petroli conducted several pilot-plant runs in the LC-Fining pilot-plant facility at AGIP’s research center related to processoptimization/ development and catalyst screening programs for RAM, Shell Canada, andPetro-Canada. The long-term goals of this joint effort are to extend the database, furtherenhance the correlations and models, test new process designs (e.g., interstage separa-tor/stripper, optimize quantity and interstage injection location of aromatic diluents), andcontinue to screen and evaluate new residual conversion catalysts.
CLG has built and is building several small pilot units at Chevron’s research facilitiesin Richmond, California to support its residual upgrading technologies: ARDS, VRDS,OCR, upflow reactor (UFR), and LC-Fining.
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PROCESS CHEMISTRY
Residual hydrocracking is accomplished at relatively high temperatures and high pressuresin the presence of hydrogen and a residual conversion catalyst to hydrogenate the productsand prevent polymerization of the free radicals as cracking reactions proceed. The catalystconsists of a combination of metals that promote hydrogenation (e.g., cobalt and molyb-denum, or nickel and molybdenum) deposited on an alumina base.
The two most important reactions that take place in residual hydrocracking are thermalcracking to lighter products and catalytic removal of feed contaminants. These reactions gen-erally require operating temperatures between 750 and 850°F, hydrogen partial pressures from1100 to 2200 lb/in2, and space velocities ranging from 0.1 to 0.8 (vol oil/h)/vol of reactor.
Hydrocarbons present in the residual are generally classified as oils, resins, andasphaltenes. Typical residual may contain about 20 percent oils, 65 percent resins, and 15percent asphaltenes. The asphaltenes are the high-molecular-weight material in the resid-ual that typically contains a large concentration of sulfur, nitrogen, metals, Conradson car-bon, and highly condensed polynuclear aromatics.
Nitrogen removal is generally much more difficult than sulfur removal. Some nitrogencompounds in the cracking reactions are merely converted to lower-boiling-range nitrogencompounds rather than being converted to NH3.
The highest concentration of metals (V and Ni) resides in the asphaltene fraction withsome in the resin fraction. The oil fraction tends to be nearly free of metals. Metals areremoved as metal sulfides. Unlike sulfur and nitrogen, which are converted and “escape”as H2S and NH3, the vanadium and nickel removed are absorbed on the catalyst. Thesemetals are known to plug the catalyst pores, and this pore blockage results in catalystdeactivation.
The conversion of Rams carbon is economically important if LC-Fining vacuum bot-toms are fed to a downstream coking unit. A lower-Rams-carbon-content residual productto the coking unit means less coke make and thus a higher yield of liquid fractions that cansubsequently be converted to transportation fuels. Another option to limit the coke makein the downstream coking unit is to maximize the pitch conversion at the LC-Fining resid-ual hydrocracker.
Residual Conversion Limits
There are many factors that affect the sediment formation rate and consequently the reac-tor operability and residual conversion limits, including
● Residual asphaltene content● CCR reactivity● Thermal severity (conversion)● Catalyst type and activity● Hydrogen partial pressure● Type and quantity of diluents● Residual resin content● Reactor temperature
Of these, the first six have the greatest influence. Many pilot-plant tests showed that sedi-ment formation is directly proportional to the asphaltene content of the feed and inverselyproportional to the CCR reactivity.
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Coke Precursor/Organic Sediment Formation
The nature and origin of the coke precursors are often not precisely known. However, amechanism of sediment (i.e., coke precursor) formation in processes that involve thermalcracking in addition to hydrocracking and hydrogenation, such as the LC-Fining process,has been generally postulated as described by the following reaction chemistry.
Thermal cracking—formation of free radicals
R – R → R* � R*
Free radicals react to form olefins or asphaltenes:
R – CH2 – CH2* → R – CH � CH2 � H*
R* � R� → R – R�
Termination or recapping of free radicals by hydrogenation
R* � H* → R – H
At elevated temperatures, thermal cracking reactions generate free-radical speciesdue to the rupture of carbon-carbon bonds. The free radicals may react with hydrogenin the presence of the catalyst to form stable products. This reaction predominates inthe LC-Fining process where high hydrogen partial pressures are always maintained. Ifproper conditions are not maintained, the free radicals may also combine with otherfree radicals to form higher-molecular-weight free radicals. This chain reaction cancontinue until very high-molecular-weight, insoluble species (coke precursors/sedi-ments) are produced. As the temperature is increased to obtain higher conversions, therate of generation of free radicals, and consequently coke precursors, can increase, cre-ating phase separation and potential instability in the reactor if it is allowed to exceedthe solubility limit.
Means of Controlling Coke Precursors/Organic Sediments
Control of coke precursors (organic sediments) can be accomplished in three ways: (1)Their formation is minimized or eliminated by using extremely high hydrogen partial pres-sures or very active catalyst; (2) the coke precursors are maintained in solution by addingaromatic diluents; and/or (3) the coke precursors are removed from the system.
The catalyst used in the LC-Fining process has an excellent ability to control the for-mation of these coke precursors, and aromatic diluents have been used successfully.
The continuous, physical removal of coke precursors (via filtration, centrifugation, etc.)from the reactor loop can be accomplished by bottoms recycle and removal of the cokeprecursors from the recycle stream, an approach pilot-tested and patented by Lummus forhigh conversion LC-Fining.
Catalyst Deactivation
The rate at which a catalyst deactivates during residual oil hydrocracking is a complexfunction of many parameters that can be categorized into three distinct classes. The firstconsists of the physical and chemical properties of the residual feedstock to be processed.The second is concerned with the nature of the catalyst itself. The third is the effect of the
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operating conditions used to obtain the desired levels of conversion and desulfurization(temperature, space velocity, hydrogen partial pressure, etc.).
The most significant causes of catalyst deactivation are metals and carbon laydown.Concurrent with the desulfurization of residuals is a demetallization reaction. Residualcracking products have nickel and vanadium contents markedly lower than those in thefeed to the unit. The metals accumulate on the catalyst, causing deactivation. It has beenproposed that the organometallics simply block the outer physical surface of the catalyst.
Carbon laydown on catalyst is influenced by feedstock characteristics and conversionseverity. Carbon accumulation is high in all operating scenarios, ranging from slightlyunder 10 wt % for low-temperature hydrodesulfurization of atmospheric residuum to over40 wt % for high conversion of vacuum bottoms.
Residual hydrocracking is apparently diffusion-controlled. It has been found that 1/32
in catalyst performs better than 1/16 in catalyst.
LC-FINING REACTOR
Following is a schematic of an ebullated bed LC-Fining reactor (Fig. 8.7.1). Fresh feed andhydrogen enter the reactor at the bottom and pass up through a catalyst bed wherehydrodesulfurization and other cracking and hydrogenation reactions occur. A portion ofthe product at the top of the reactor is recycled by means of an internally mounted recyclepump. This provides the flow necessary to keep the catalyst bed in a state of motion some-what expanded over its settled level (i.e., ebullated). The catalyst level is monitored andcontrolled by radioactive density detectors, where the source is contained inside the reac-tor and the detectors are mounted outside. Temperature is monitored by internal couplesand skin couples. The performance of the ebullated bed is continuously monitored andcontrolled with the density detectors and temperature measurements that verify proper dis-tribution of gas and liquid throughout the catalyst bed. Temperature deviations outside thenormal expected ranges that might suggest maldistribution will cause the distributed con-trol system (DCS) to activate alarms or initiate automatic shutdown on the heaters, hydro-gen feed, and/or reactor section, as required.
Catalyst is added and withdrawn while the reactor is in operation. The reactors can bestaged in series, where the product from the first reactor passes to a second reactor and, ifnecessary, to a third reactor. After the final reactor, the product goes to a high-pressure/high-temperature separator.
LC-FINING PROCESS FLOW SCHEMATICS
Process Description
Following is a simplified process flow diagram of an LC-Fining unit with a close-coupled,integrated, fixed-bed hydrotreater/hydrocracker (Fig. 8.7.2).
Oil feed and hydrogen are heated separately, combined, and then passed into the hydro-cracking reactor in an upflow fashion through an ebullated bed of catalyst. Under the effectsof time, temperature, and hydrogen pressure, and aided by the catalysts, the feed oil iscracked and hydrogenated to produce lighter, higher-quality products. A portion of the liquidproduct from the large pan at the top of the reactor is recycled through the central downcomerby means of a pump mounted in the bottom head of the reactor. This flow gives the neededvelocity for bed expansion and aids in maintaining near-isothermal reactor temperature.
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The hydrocracking reactions are exothermic, resulting in a temperature rise from inletto outlet depending upon the reaction operating severity. However, because of the mixingeffect of the internal recycle liquid, the bed operates essentially isothermally. Catalyst isadded and withdrawn batchwise to maintain an equilibrium catalyst activity without theneed for unit shutdown.
Reactor products flow to the high-pressure/high-temperature separator. Vapor effluentfrom the separator is let down in pressure before heat exchange, removal of condensates,and purification. Handling the recycle gas at low pressure offers considerable savings ininvestment over the conventional high-pressure recycle gas purification system.
After stripping, the recycle liquid is pumped through the coke precursor removal step(a physical means of separation such as centrifugation) where very small quantities ofinsoluble heavy hydrocarbons or carbonaceous solids are removed. The clean liquid recy-cle then passes to the suction drum of the feed pump. Net product from the top of the recy-cle stripper goes to fractionation; net heavy oil product is directed from the stripperbottoms pump discharge to vacuum fractionation.
The LC-Fining reactor effluent vapor, along with distillate recovered from the heavy oilstripper overhead, any virgin atmospheric gas oil recovered in the prefractionator upstreamof the LC-Fining unit, and vacuum gas oil recovered in the LC-Fining vacuum fractiona-tor, is all charged to a “wide-cut,” close-coupled, integrated, fixed-bed hydrotreater/hydro-cracker located immediately downstream from the last ebullated bed LC-Fining reactor.The inlet temperature to the first bed is controlled by adjusting the amount of heat extract-ed from the LC-Fining reactor vapor stream, and the temperature of the distillate liquid iscontrolled by a combination of hydrogen and liquid quench. The effluent from thehydrotreating reactors is separated into a vapor and heavy distillate liquid stream, with the
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Catalyst Addition Line
Density DetectorRadiation Source Well
Density Detectors
Normal Bed Level
CatalystWithdrawal Line
Recycle Pump
SkinTC's
Thermowell NozzleEffluent
Feed
FIGURE 8.7.1 LC-Fining reactor.
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liquid stream routed to the hydrotreated distillate fractionator. The vapor stream is amine-treated, purified through a pressure swing absorption (PSA) or membrane unit, and recom-pressed and recirculated as treat gas to the LC-Fining reactors.
The high-conversion (�80 percent) LC-Fining process differs from the basic processin that bottoms recycle is practiced. The recycle liquid is let down in pressure and passesto the recycle stripper where it is fractionated to the proper boiling range for return to thereactor. In this way, the concentration of bottoms in the reactor, and therefore the reactionproducts slate, can be controlled.
LC-FINING TECHNOLOGY ADVANTAGES
Several advances in the LC-Fining residual hydroconversion technology have significant-ly reduced the capital investment and have further extended the conversion limits and pro-cessing limitations. These include
One of the key areas of process optimization resulted in the H2 purification system. In ear-ly designs (Amoco), a lean oil system was used to purify the recycle gas, and the maxi-mum purity achievable was 82 percent. In 1984, Lummus developed and patented alow-pressure H2 purification system, which has been utilized in all commercial operatingunits since. With low-pressure H2 purification, the gas exiting the last reactor is immedi-ately let down and cooled at low pressure and then purified in a PSA unit. This permitshigh hydrogen treat gas purities, generally exceeding 97 vol %. As a result, the treat gascirculation rates were reduced by 50 to 60 percent and the reaction system design pressureby 10 percent, while still satisfying the hydrogen partial pressure requirements. Thischange, in conjunction with replacing 12 high-pressure equipment services (includinghigh-pressure exchangers and drums; high-pressure lean oil and amine absorbers; and leanoil, amine, and wash water pumps) with low-pressure equipment, significantly reduced theLC-Fining unit investment cost. The only drawback with this low-pressure H2 purificationscheme (i.e., using a PSA system) is that the recycle gas had to be recompressed from lowpressure back to reactor operating pressure, requiring an increase of 25 to 30 percent in theoverall power consumption.
In 1998, the use of membranes was evaluated for purification of the recycle gas, andsimilar treat gas purities were achieved with membranes as with a PSA system.Membranes allow the same reduced reaction system pressure and lower treat gas circula-tion rates as with a PSA unit, but with the added benefit that the purified recycle gas isavailable at higher pressures. Consequently, the recycle gas can be recompressed in a sin-
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gle stage of compression versus the two or three stages of compression required with aPSA system, resulting in a 20 percent reduction in the overall unit power consumption. Inaddition, based on current high-pressure equipment pricing, the unit investment is slightlyless with membranes than with a PSA unit.
Low Treat Gas Rates
By using high-purity recycle gas, it is possible to achieve the desired hydrogen partial pres-sure with much lower hydrogen treat gas rates. The low gas rates have two primary bene-fits: They greatly reduce unit investment associated with the heating, cooling, purification,and recompression of the recycle hydrogen; and they significantly reduce the gas superfi-cial velocity and therefore the gas holdup in the reactor. This provides for greater liquidresidence time per unit reactor volume, thereby reducing the reactor volume required toachieve the desired thermal conversion and catalytic contaminants removal. Conversely,when designing to a maximum allowable superficial gas velocity within a specified reac-tor diameter constraint, it is possible to substantially increase the LC-Fining unit capacity.
Reduced gas rates enhance overall reactor operation since internal liquid recirculationis increased as a result of the reduced gas superficial velocity and holdup. This results inbetter back-mixing of liquid and catalyst within the reactor, thereby minimizing incidencesof catalyst bed slumping and channeling and flow maldistribution. The use of low treat gasrates is seen in all commercially operating LC-Fining units, which operate with a totalhydrogen/chemical hydrogen ratio of 2.5:3.
Integrated Fixed-Bed Hydrotreater/Hydrocracker
Several recent designs incorporated a close-coupled, integrated, fixed-bed hydrotreater/hydrocracker immediately downstream of the LC-Fining reactors. In this design, the vaporstream from the LC-Fining reactors, the distillate recovered from the heavy oil stripperoverhead, and the straight-run atmospheric and vacuum gas oils are fed to a wide-cut,fixed-bed hydrotreater/hydrocracker operating at essentially the same pressure level as theLC-Fining reactors.
Excess hydrogen contained in the LC-Fining reactor effluent vapor is used tohydrotreat the distillate fractions. Additional hydrogen, equivalent only to the chemicalhydrogen consumed in the fixed-bed reactor, is introduced as quench to the second andthird catalyst beds. If necessary, the remaining portion of the reaction heat is dissipated byinjecting liquid quench recycled from the separator downstream of the hydrotreater/hydro-cracker.
LC-Fining reactor effluent vapor is first contacted with VGO in a wash tower in orderto remove any potential residual entrainment and entrainment of catalyst fines into thefixed-bed reactor. To maintain the desired HDS/HDN fixed-bed catalyst activity over the run length, a certain percentage of demetallization catalyst is included on the top of thefirst bed to remove metals and CCR contained in the hydrotreated feed fraction. Residualcarryover is mitigated by providing additional surge upstream and by having spare (i.e.,redundant) upstream separator and wash tower level control systems.
By incorporating the fixed-bed hydrotreater within the LC-Fining reaction system, theHP system service count is reduced from approximately 14 pieces (for a stand-alonehydrotreater) to only 6. In addition, since excess hydrogen in the LC-Fining reactor efflu-ent vapor is used to hydrotreat the straight-run and LC-Fining distillate fractions, the needfor additional recycle gas compression is eliminated. As a result, the investment is signif-icantly reduced compared to that for a stand-alone hydrotreater/hydrocracker.
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Interreactor Separator/Stripper
Recently proposed process design configurations incorporate the use of an interreactorseparator/stripper, which permits higher liquid capacities to be achieved for a given reac-tor cross section. Gas superficial velocities through the downstream reactors are reducedby separating the vapor between reaction stages and routing the vapor to the final reactoreffluent separator. This design provides for parallel flow of gas to each reaction stage whilemaintaining the benefits of series flow liquid operations. In a conventionally designed unit,the effluent vapor from the upstream reactor is combined with additional treat/quench gas,and all the vapor is directed to the downstream reactors. It is possible to process muchhigher feed throughputs to the LC-Fining unit for a given reactor cross-sectional area.
The reduction in stripped feed to the second-stage reactor and the associated increase inresidual concentration enable reactor volume to be reduced for a given capacity and conver-sion, while the higher final-stage hydrogen partial pressure can be utilized to reduce eitherthe catalyst addition rate or the reactor operating/design pressure. More important, the reduc-tion in the paraffinic naphtha and light distillate fractions in the more highly converted third-stage liquid reduces the sediment formation for a given residual conversion. The residualconversion limits can be further extended due to this change in liquid composition.
Liquid Recycle Pan
A modified, two-stage, liquid recycle pan design, which increased conversion in the reac-tor and minimized upsets associated with the recycle pump bed expansion system, was putin operation at Amoco in 1986 and at Syncrude in 1988. The design operated with a morequiescent zone in and above the pan than the original version, minimizing the entrainmentof gas bubbles into the recycle fluid, which in turn minimized gas holdup in the reactor.With this modification at Syncrude’s LC-Fining unit, the conversion increased approxi-mately 4 percent for the same reactor operating conditions.
Subsequently, Amoco and Syncrude, with Lummus’s participation, developed a newpan design aimed at permitting operation at still higher capacities and treat gas rates.Following the installation of this new pan at Syncrude in 1996, Syncrude has been able tocharge 58,000 BPSD of 650°F� A-tar (unit originally designed for 40,000 BPSD), at sim-ilar reactor gas inlet superficial velocities, while achieving a 975°F� conversion of 57 to58 vol %. Thus, the installation of this new pan increased conversion an additional 2 to 3vol %.
Bottom Head Feed Distributor
Based on cold flow modeling work done by Amoco in 1985/1986, a bottom head feed dis-tributor was added to the LC-Fining reactors. It provides for better mixing of the feed oil,gas, and recycle oil while providing for better distribution of oil and gas to the cap and ris-er assembly.
Reactor Distributor Design
The primary reactor grid distributor is a bubble cap type with slotted risers for distributionof vapor and liquid to each cap. Each riser contains a seat and ball to prevent back-migra-tion of catalyst below the grid into the reactor bottom head, should the recycle pump stop.The seat contains a small V notch to permit oil to be drained from the reactor.
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Amoco performed substantial cold flow modeling of the distributor grid that led to theinstallation of catalyst slides. This was found to reduce instances of localized catalyst set-tling near the wall, maintaining a cleaner reactor environment and increasing the runlength between turnarounds.
Alternate grid modifications are being considered, such as using larger risers aroundthe reactor perimeter and/or varying the riser slot heights to either increase the flow of bothgas and liquid, or, preferentially, increase the flow of gas, at the reactor wall.
Vacuum Bottoms Recycle Operation without Coke Precursor Removal
Vacuum bottoms recycle (VBR) operations have been extensively pilot-tested, and VBRwas found beneficial to
● Increase residual conversion● Minimize hydrogen consumption● Maximize yield of vacuum gas oil● Minimize light ends (C1-C3, gas) make● Maximize unit capacity for a given level of conversion and reactor volume
The most significant advantage of VBR operation is the increase in residual conversion atthe same operating severity. With VBR, the residual concentration increases within thereactors, thereby increasing the conversion rate.
VBR does have some drawbacks. The operating company may prefer the higher yieldof lighter distillates realized without recycle. Recycle is a form of back-mixing and canresult in higher impurities in the product. Recycle of vacuum bottoms at higher tempera-tures is ineffective for control of sediment, as the sediment formation can rise rapidly, cre-ating potential difficulties for maintaining proper catalyst bed ebullition.
COMMERCIAL OPERATIONS
Residue hydroconversion units that utilize the ebullated bed LC-Fining technology aresummarized in Table 8.7.1. Shell Canada is scheduled to start up in 2003 and Petro-Canadaaround 2005.
PROCESS FLEXIBILITY
The LC-Fining unit has great inherent flexibility to meet variations in feedquality/throughput, product quality, and reaction operating severities (temperature, spacevelocity, conversion, etc.). This flexibility is a direct result of the ebullated catalyst bedreactor system. In an ebullated bed unit, if the metals or sulfur content of the feed increas-es, the product quality is maintained by increasing catalyst consumption. Conversely, thecatalyst consumption is reduced if the feed quality improves.
BP has utilized this flexibility to process heavy sour vacuum bottoms from a blend ofdifferent crudes, including Maya and Bachequero. At the same time, BP has increased con-version of vacuum bottoms to distillate to 75 to 80 percent typically, at full feed rate, andup to 92 percent at reduced feed rates. The original design was 60 percent at full feed rate.
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The world’s first ebullated bed residual upgrader operated by Cities Service OilCompany utilized this flexibility to process atmospheric bottoms, FCC heavy cycle oil,propane deasphalter bottoms, and vacuum bottoms.
The Syncrude unit was originally designed to process 500°F� Athabasca bitumen con-taining 55 wt % 975°F� vacuum residual. More recently, they installed a vacuum pre-fractionation system and are now processing a blend of atmospheric and vacuum bottomscontaining 75 wt % 975°F� vacuum residual.
Sufficient operating flexibility is also normally provided in the design to enable the unitto operate in the future with VBR, which provides for future options to increase either con-version or unit throughput.
Hydrogen consumption 120–340 N m3/m3 (700–2000 SCF B)
Desulfurization 60–95 percent
Demetallization 70–98 percent
CCR reduction 40–75 percent
WIDE RANGE OF FEEDSTOCKS
A wide range of heavy oils have been processed in LC-Fining units. For example, the BPunit handles many of the poorest-quality vacuum residual in the world, including Mexican,Venezuelan, and Middle Eastern. Feed typically is under 5° API and has more than 4 wt %sulfur and more than 400 ppm metals. Table 8.7.2 shows the major crudes processed byBP at Texas City to produce LC-Fining feedstock from 1984 to 1992.
Table 8.7.3 shows the BP unit operating results.
YIELDS AND PRODUCT QUALITY
LC-Fining unit product yields for processing Arabian heavy vacuum bottoms to conver-sion levels of 40, 65, and 80 percent are listed in Table 8.7.4. All these conversions can beachieved in the same LC-Fining unit, illustrating the great flexibility of the process.
The yield structure and product properties are estimated from generalized correlationsthat were derived from extensive pilot-plant and commercial data. Typical product proper-ties for a 65 vol % conversion case are shown in Table 8.7.5.
8.80 HYDROTREATING
CHEVRON LUMMUS GLOBAL EBULLATED BED BOTTOM-OF-THE-BARRELHYDROCONVERSION (LC-FINING) PROCESS
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EBULLATED BED BOTTOM-OF-THE-BARREL HYDROCONVERSION 8.81
TABLE 8.7.2 Major Crudes Processed by BP at Texas City to Produce LC-Fining Feedstock, 1984 to 1992
Maya Menemota Djeno Basrah LightHondo Gulf of Suez Mix Arab Medium Basrah HeavyHeavy Bachequero Isthmus Arab Heavy Qua IboeLloydminster Alberta Light Coban MereyAlaskan North Slope Pilon Cold Lake LeonaWest Texas C Laguana Peace River KuwaitKhafji BCF-17 Yemen KirkukJobo Tia Juana Pesado Olmeca
Feedstock: Blend of Maya (over 40 percent), Venezuelan, Middle Eastern, domestic includingANS, and other vacuum bottoms.
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CATALYSTS
A series of catalysts are available for use in LC-Fining units. The first-generation catalystsin commercial use had adequate HDM/HDS activity with acceptable sediment levels. Thesewere less expensive than more recently developed, enhanced contaminant removal/sedimentcontrol catalysts. New-generation catalysts are needed to produce low-sulfur fuel oils (fromvacuum bottoms) of 1 wt % sulfur or less with minimum sediment levels (�0.15 wt %) forpipeline stability. The other requirement of a good LC-Fining catalyst is to maintainimproved reactor operability/stability at high-temperature/high-residual conversions.
The residual hydroprocessing catalysts are small (1/32 to 1/8 in), extruded, cylindricalpellets made from an aluminum base. The pellets are impregnated with active metals (Co,Ni, Mo, W, and other proprietary materials) that have good hydrogenation, demetallation,desulfurization, and sediment control activity. Catalyst manufacturing processes are tai-lored to manipulate physical and mechanical properties such as size (length and diameter),attrition resistance, crush strength, pore size distribution, pore volume, and effective sur-face area. Catalytic performance is affected by the complicated nature of the “active site”and dispersion and distribution of activators and promoters.
Pore size control and distribution are key factors in the behavior and formulation ofresidual conversion catalysts. The pore sizes need to be sufficiently large to allow the dif-fusion of the large residual/asphaltene molecules that require upgrading. Unfortunately asthe pore diameter increases, the surface area and the hydrogenation activity decrease. Thediffusion of large molecules is reduced further because of pore mouth plugging due to car-bon laydown and metal sulfide buildup from vanadium and nickel atoms that are removedfrom the residual feed. Metal sulfides are formed from the oxidative state of the catalyst inthe LC-Fining reactor environment (presulfiding reactions with sulfur in heavy oils, etc.).
Catalysts are also optimized for specific functions—such as metals removal, sulfurremoval, carbon residue reduction, and high conversion—while maintaining a clean prod-uct low in organic sediments. The catalyst system developed by BP for its LC-Fining unitat Texas City utilizes a proprietary demetallization catalyst in the first reactor and a high-activity nickel/Mo desulfurization catalyst in the second and third reactors.
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One of the key features of the LC-Fining process is the use of countercurrent catalystaddition to optimize catalyst usage. Fresh catalyst is added to the third reactor, then reusedby withdrawing it and adding it to the second reactor. The catalyst can then be used a thirdtime by withdrawing it from the second reactor and adding that material to the first reac-tor. Catalyst cascading results in higher overall kinetics rate constants and therefore betteroverall catalyst utilization based on the concentration of metals in the spent catalyst dis-charged from the first stage. This mode of addition and withdrawal has the added benefitof exposing the most highly converted residual to the most active catalyst. This reduces thesediment formation in the last reactor and thus allows reactor operability and conversionlimits to be extended.
INVESTMENT COSTS
Compared to other residual hydrogenation processes, the LC-Fining process has severalintrinsic advantages:
● Very high conversion levels● Low investment cost● Lower operating costs● Lower hydrogen losses● More efficient hydrogen and heat recovery● Lower maintenance
Much of the cost of a hydrogenation unit is connected to the gas recycle rate; high gasrecycle rate results in high compressor, piping, furnace, heat exchanger, and separatorcosts. The LC-Fining process is the lowest-cost commercially proven residual hydrogena-tion process due to the low total hydrogen rate and the proprietary low-pressure recoverysystem. The low-pressure recovery system saves 8 to 10 percent of the capital investmentof an LC-Fining unit, which translates to $10 million to $30 million U.S. or more, depend-ing on plant size. Gas losses are also maintained at a low level by using low hydrogen cir-culation rates.
When the integrated hydrotreater/hydrocracker is incorporated into the LC-Fining unit,additional savings in investment of as much as 40 percent of the cost of separatehydrotreating facilities are possible.
Depending on feedstock properties, operating severities, product requirements, andprocessing objectives, the typical ISBL investment cost of an LC-Fining unit ranges from$2000 to $5000 U.S. per BPSD.
EBULLATED BED BOTTOM-OF-THE-BARREL HYDROCONVERSION 8.83
CHEVRON LUMMUS GLOBAL EBULLATED BED BOTTOM-OF-THE-BARRELHYDROCONVERSION (LC-FINING) PROCESS
