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A recent study for a grassroots 30,000-b/sd hydrotreater that can produce ultra-low sulfur diesel (ULSD) helped assess the critical process Refining design issues for meeting future ULSD requirements. The feed, a 70:30 mixture mixture of virgin distillate and FCC FCC light cycle cycle oil, was the basis for a previous Mustang Engineers & Constructors LP study for revamping an existing facility to meet the ULSD specification.1 This first of two articles presents a comprehensive overview of design con-
Ultra-low sulfur diesel In mid 2006, the US Environmental Protection Agency will begin enforcing rules that mandate a maximum sulfur content of 15 ppm (wt) for on-road diesel fuel. Refiners are considering several strategies to comply with this rule: • Revamping existing facilities commissioned in the early 1990s that now meet the low-sulfur diesel specification of 500 ppm (wt) sulfur. • Constructing grassroots facilities to meet the new specification. • Executing various combinations of new and revamped facilities.
The problem Middle distillates contain many sulfur species, including mercaptans, sulfides, thiophenes, and aromatic sulfur compounds. Sterically hindered dibenzothiophenes are a group of aromatic sulfur compounds compound s that are most difficult to remove; they constitute a large fraction that remains in diesel after hydrotreating to the current specification specification 500 ppm (wt). This is particularly true for diesel fuels that contain significant quantities of cracked stocks, like FCC light cycle oil, that contain a large concentrati concentration on of aromatic sulfur compounds. Effective removal of these species requires tailored catalysts and process conditions. Refiners must also consider other factors, such as feed nitrogen content and aromatics equilibrium. These issues have been addressed in detail in the literature.2-7
siderations for grassroots ULSD hydrotreaters.The drotreaters. The chemistry of middle distillate sulfur sulfur species (see box), reaction process variables, and major factors influencing reactor design are covered. Engineering aspects of hydrotreater design are are discussed, includ including ing a suitable operating pressure level that satisfies reaction conditions and the practical limits of piping mechanical design. Alternative process configurations for ULSD hydrotreaters weree studied, wer studied, along with key design parameters and metallurgical considerations for major equipment. Part Pa rt 2, next wee week, k, will presen presentt process simulations for different process configurations at reactor inlet pressures of 800 and 1,100 psi. We estimated capital costs and life cycle costs for three cases.
ULSD HYDROTREATING—1
Study outlines optimum ULSD hydrotreater design
Design considerations The ULSD hydrotreater designer must account for these factors during the front-end design process: • Feed characte characteristics ristics and variabi variability lity.. • Other produ product-qua ct-quality lity requi requirerements, especi especiall allyy cetane index. • Ca Catal talyst yst sel select ection ion.. • Optimi Optimizati zation on of of reactor reactor proc process ess variables. • Equipm Equipment ent design requir requirements ements.. • Rel Relia iabil bility ity. • Futur Futuree off-road off-road diesel sulfu sulfurr stanstandards. • Minim Minimizing izing prod product uct contam contaminat ination. ion. • Handl Handling ing of off-sp off-specif ecificati ication on diesel product.
Les Harwell Sam Thakkar Stan Polcar R.E. Pa Palmer lmer Mustang Engineers & Constructors LP Houston Pankaj Pa nkaj H. Desai Akzo Nobel Catalysts LLC Houston
Process flow Fig.. 1 shows a simplified process Fig flow diagram for a diesel hydrotreater. Stripper bottoms heats fresh feed from the surge surge drum, which then mixmixes with recycle recycle hydrogen. hydrogen. Reactor effluent further heats the combined feed, which is then heated in the charge heater. Reactor inter-bed quench may be required depending on the volume of Based on a presentation to the 2003 NPRA An nual Meeting, Mar Mar.. 23-25, 2003, San Antonio. Antonio.
Reprinted with revisions to format, from the July 28, & August 4, 2003 editions of OIL & GAS JOURNAL Copyright 2003 by PennW PennWell ell Corporation
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cracked stocks—FCC light cycle oil and light coker gas oil—in the feed. Fig. 1 shows recycle gas used as quench. Reactor effluent heats the combined feed and flows to a hot, highpressure separator (HHPS). Vapor from the HHPS heats the recycle gas and stripper charge before cooling in the product condenser and entering the cold, high-pressure separator (CHPS). Wash water, injected upstream of the product’s condenser, helps remove ammonium hydrosulfide. HHPS liquid combines with heated CHPS liquid and flows to the product stripper. Vapor from the CHPS contacts with amine in a scrubber for H2S removal and flows to the recycle compressor
suction drum. Makeup hydrogen is compressed and combines with recycle gas in the suction drum. The compressor suction drum can purge some recycle gas to improve recycle gas hydrogen purity. The quantity depends on the reactor’s hydrogen-partial-pressure requirements and makeup hydrogen purity. In the product stripper, superheated steam feeds the tower’s bottom and helps remove H2S. Stripper overhead vapors condense and flow to the stripper accumulator. Accumulator vapor and liquid, known as wild naphtha, are processed in offsite facilities. Stripped ULSD product supplies heat to the feed stream and then flows to dry-
ing facilities, which can be a coalescersalt dryer or vacuum drying system.
Feed, product characteristics Sulfur, nitrogen, and aromatics have the greatest impact on the ULSD facility design. Feed nitrogen has a significant impact on the operating pressure. The refiner must remove nitrogen to essentially the same level as sulfur to meet the ultra-low sulfur target. This means that the catalyst and hydrogen partial pressure must be consistent with a high nitrogen-removal operation. Light coker gas oil and FCC light cycle oil normally contain most of the feed nitrogen. Feed aromatics content
ULSD HYDROTREATER
Fig. 1
Makeup H2 LCO
Feed heater Makeup H2 compression
Straight run diesel
Reactor Recycle compressor Amine treating
Quench
Feed
ULSD product
Gas to fuel
Hot high-pressure separator
Sour water Sour naphtha
Cold high-pressure separator
Steam Stripper
Wash water
Sour water
governs chemical hydrogen consumption at the low space velocities and high hydrogen partial pressures required for ULSD production. Product cetane or gravity generally determines the amount of cracked stocks that a refiner can include in the feed. There is a small increase in the gravity and cetane index during the hydrotreating reaction. If a significant improvement in cetane (3-5 numbers) or gravity is required, a multistage design using aromatics saturation catalysts in the second stage may be a more economical option. The magnitude of improvement in gravity and cetane determines this design choice. Design product sulfur is a key target, not only from a process design standpoint, but also for offsite storage and transfer considerations. Most new and revamp facility designs are for a product sulfur content of 8-10 ppm (wt). Refiners should pilot test the feed to confirm reaction process conditions. Testing for variations in feed characteristics, especially the FCC light cycle oil and coker light gas oil’s back-end distillation, is important because product separation in these facilities is notoriously poor. This can result in a temporary spike in the most difficult-to-treat sulfur compounds in the hydrotreater feed and requires a higher reactor temperature. This increases the catalyst deactivation rate and directionally reduces cycle length.
hydrogen partial pressure, and reactor temperature during the process design phase for a given cycle length and treating severity. Catalyst deactivation rate falls with increasing hydrogen partial pressure; therefore, the space velocity can increase for a constant cycle length. This is, however, at the expense of higher hydrogen consumption. Another important variable is the ratio of total hydrogen supplied to chemical hydrogen consumption. This value should be 5-6 for feeds with significant quantities of cracked stocks. This ratio is normally less important during the process design, but it can control the design hydrogen-circulation rate, especially when a refiner exclusively uses high-purity makeup hydrogen. When setting the hydrogen partial pressure and total operating pressure, one should be aware that the alloy piping flanges limit the reactor pressure, especially when the feed contains significant quantities of cracked stocks. This includes piping between the feed exchangers and feed heater, between the heater and reactor, and between the reactor and combined feed exchangers. Piping is normally a 300-series stainless steel; for 600-psi flanges this corresponds to a maximum operating pressure of around 800 psi at the reactor inlet. For 900-psi flanges, the reactor operating pressure can increase to about 1,100 psi.These pressures account for relieving conditions. These criteria should not necessarily govern a refiner’s selection of operating pressure; however, pressures barely exceeding these limits and requiring the Reaction process variables next-higher-rated ANSI flange class are Key reaction process variables include: costly. • Space velocity. Actual pipe wall thickness is calcu• Hydrogen partial pressure. lated from design conditions; the limit• Makeup hydrogen purity. ing value is used for the piping specifi• Ratio of total hydrogen to chemi- cation in the reactor loop’s alloy seccal hydrogen consumption. tions. • Cycle length. Actual pipe materials can also be • Reactor temperature. 316, 321, and 347 grades of stainless steel. Type 347 has a higher allowable Refiners can use a nickel-molybdenum catalyst for feeds that have high aro- stress value than the other grades but is matics or nitrogen; an appropriate selec- normally more expensive. Availability tion of graded catalysts in the bed’s top and delivery time for the various stainwill mitigate reactor pressure buildup. less components are also a major conRefiners must optimize reactor space sideration. velocity, hydrogen treat gas quantity, The existing equipment and reactor
loop’s piping limit the hydrogen partial pressure in revamp designs. A higher treat gas rate can increase the hydrogen partial pressure; this is usually limited due to an associated increase in the reactor loop’s pressure drop and the corresponding maximum operating pressure of system components. Lower-purity makeup hydrogen requires higher hydrogen circulation rates to maintain a target hydrogen partial pressure; it may even require a purge stream from the cold separator. If makeup hydrogen purity is too low, there is no combination of recycle rate and purge that will achieve the target reactor outlet partial pressure. For a revamp design, increasing makeup hydrogen purity is the most effective way to increase the hydrogen partial pressure. Catalyst cycle lengths of 24-36 months are typical for new designs. This is because, at some point, factors other than catalyst activity (such as reactor pressure drop) will limit the cycle. For a fixed space velocity, cycle length increases with a higher hydrogen partial pressure. Maximum reactor outlet temperature at end-of-cycle catalyst conditions is usually 725-750° F. to avoid aromatics saturation equilibrium constraints and to maintain product color. The amount of cracked stocks in the feed and the crude source also influence this temperature. Hydrotreating catalyst performance correlations for reactor temperature is usually based on the weighted average bed temperature. WABT is the reactor inlet temperature plus two thirds of the reactor temperature rise. Quenching limits the temperature rise to 40-50° F. in each bed. For a 50° F. temperature rise and a 725° F. maximum reactor outlet temperature, the end-of-run WABT is (7252 50) + ⁄ 3(50) = 708° F. The start-of-run WABT must be high enough to remove the required amount of sulfur and nitrogen.The catalyst deactivation rate at the design space velocity and hydrogen partial pressure determines cycle length. During the cycle, WABT will increase 30-50° F. with lower deactivation rates at higher hydrogen partial pressures.
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Reactor design A superficial mass velocity set at 2,000-5,000 lb/hr/sq ft determines the initial reactor diameter. A value of 3,500 allows for a reasonable turndown rate and some upside allowance. Mass flow includes all hydrocarbons and hydrogen at the reactor inlet. The refiner should consider a twotrain design if the required reactor diameter is larger than could be shop fabricated and shipped to the plant site—usually 12-14 ft for overland shipment. For diesel hydrotreaters, this is normally 30,000-40,000 b/sd for each reactor train. Refiners with adequate water access and dock facilities can consider larger
reactors and larger single-train capacities. An appropriate correlation8 estimates the pressure drop for two-phase flow through the reactor. The pressure drop calculation should also include interbed quench. Catalyst physical properties, including void fraction and equivalent particle diameter, have a significant impact on reactor pressure drop. To ensure good distribution in the catalyst beds, most ULSD reactors are dense loaded instead of sock loaded. The clean pressure drop should be 0.5-1.0 psi/ft of catalyst bed. For a required catalyst volume, bed heights are either based on the heat of
reaction and maximum temperature rise or set at 30-40 ft.The designer then estimates the clean pressure drop for each bed. The designer determines the maximum overall pressure drop when designing the bed support and compares it to the catalyst crush strength, which is the sum of: • Fouled bed pressure drop at 175% of calculated clean drop. • Dead weight of catalyst and support material. • Coke deposits at 30% of catalyst dead weight. • Liquid holdup. • 15-psi allowance for depressurization.
P RESSURE PROFILE
Fig. 2
Makeup H2 1,200
LCO
1,198 1,100
Feed heater
1,102
Straight run diesel
1,252
Makeup H2 compression
957 1,054 ∆P = 233
1,152 1,290
Reactor
1,190
Recycle compressor
1,025
1,274 1,120 1,174
Amine treating
Quench
1,123
1,060
1,022
963
Feed
1,154
ULSD product
1,002 1,100
1,254 1,186 1,096 998
1,086
1,073
1,065
988
975
968
Gas to fuel
1,176 1,276
1,000
Sour water Sour naphtha
Pressure, psi Normal
Relieving
978
965 1,062
Hot high-pressure separator
Cold high-pressure separator
Steam Stripper
Wash water
Sour water
Pressure drops for the feed distributor and redistributors, quench internals, and collector are added to the bed values to provide the overall fouled-reactor pressure drop. For ULSD, the feed distributor and redistributor design is crucial for obtaining the target product sulfur level. Several designs provide internal reactor quench to limit the temperature rise. The normal method is to use hydrogen-rich recycle gas. Quench gas also causes liquid vaporization in the reactor, which provides an extra heat sink for heat removal. Another less-common approach is to recycle CHPS liquid to the reactor. The quench limits the reactor bed temperature rise to 40-50° F. 1 4Reactor metallurgy is typically 1 ⁄ 1 1 chrome, ⁄ 2-molybdenum with ⁄ 8-in. 347 stainless overlay or cladding.The metallurgy choice often depends on the vessel thickness, which affects the cost. Cladding is usually more economical for vessels with a wall thickness up to about 4 in. An overlay is normally used with thicker walls.
1 units normally have 1 ⁄ 4-chrome shells 1 and 400-series tubes. Hot shells are 1 ⁄ 4chrome with 300-series stainless cladding and 300-series stainless tubes.
a HHPS to debottleneck an existing hydrotreater that is hydraulically limited in the reactor loop. This revamp would require, however, a modest increase in inlet flow and power for the recycle Charge heater compressor. The feed heater’s normal duty Other options include adding hot should not be less than 20% of the and cold low-pressure separators.These combined feed duty to the reactor’s in- designs marginally unload the upper let temperature at end-of-cycle condipart of the stripper and improve LPG tions; this allows for turndown and recovery. 1 heat balance control. A design margin HHPS separator base material is 1 ⁄ 41 of 10% of the combined feed exchang- chrome ⁄ 2-molybdenum, normally with er duty accommodates accelerated exa clad or overlay lining, which is 300 changer fouling. or 400-series stainless depending on The number of heater passes should the actual predicted corrosion rate. match the number of parallel feed-efThe CHPS is carbon steel resistant to fluent exchanger trains.The refiner can hydrogen-induced cracking. A 300-seflow control feed and recycle hydrogen ries stainless steel mesh pad helps coafor each train upstream of the exchang- lesce the wash water. The boot is also ers. This prevents maldistribution of the hydrogen-induced-cracking resistant 1 4-in. corrosion allowance. two-phase stream entering the heater steel with a ⁄ passes. The limit for one heater pass is Reactor products condenser 15,000-18,000 b/sd of fresh feed at A final condenser provides cooling typical diesel hydrotreater conditions. before liquid and vapor separation reThis limit is based on a maximum 8-in. gardless of the reactor product separadiameter heater tube and an overall tor configuration. An air cooler is typipressure drop of 40-50 psig. cally used, sometimes supplemented Feed exchangers Heater tube metallurgy is typically with a water-trim cooler. Feed exchangers are generally speci- 347 stainless steel. Piping and equipment metallurgy in fied to optimize heat economy; howevthis area requires special attention due er, enough duty should be allowed so to the presence of ammonia and H 2S in Reactor-product separation that the charge heater has a reasonable Several alternative arrangements can the reactor effluent. When this stream amount of turndown and overall heat- handle reactor effluent after the comcools, these two compounds combine balance control. bined feed exchangers. to form ammonium hydrosulfide, Feed exchange should not be more In general, refiners must decide which condenses as a solid on inlet than 80-85% of the total duty needed whether to use a HHPS in addition to a piping and in cooler tubes. to heat the feed to the reactor tempera- CHPS, or a CHPS only. The HHPS usualRecirculating wash water, injected ture at end-of-cycle conditions. ly operates around 500-550° F. and, before the cooler, introduces a corroThe reactor feed-effluent exchanger’s therefore, the vapors can heat recycle sive aqueous solution in the stream to mechanical design requires minimal gas and cold separator liquid. prevent plugging and under-deposit leakage. Many refiners prefer a pullA HHPS improves overall heat econ- corrosion. through tube bundle with a floating omy and results in smaller product Tube metallurgy selection is based head because it is easier to clean; how- stripper and auxiliary equipment. This on ammonia and H2S in the effluent. ever, the floating head cover is a poten- arrangement also improves oil-water The designer should ensure a target tial leak source. A U-tube bundle avoids separation in the CHPS. maximum value for Kp of 0.15 before this potential problem and is less exFor heavy gas oil hydrotreating, a choosing alloy metallurgy, where:9 pensive. hot separator is necessary for adequate Kp = mole % H2S x mole % NH3 Tubes should be seal welded to the oil-water separation.This is less of a The designer should also: tube sheet to prevent leaks from the concern for diesel hydrotreating unless • Determine makeup water quantity rolled tube joints. there are significant quantities of to limit ammonium hydrosulfide conMaterial selection for feed-effluent cracked materials, which decrease API centration in the condensed water exchangers is based on predicted corro- gravity. phase to 4 wt %. One must assume sion rates. Feed-effluent exchangers Disadvantages of this design are a 5- equal molar amounts of ammonia and have several shells in each train with 10% lower recycle gas purity that reH2S in the CHPS water and that essenvarying metallurgies. Cold shells are all quires more horsepower in the recycle tially all ammonia dissolves in the wacarbon steel. Intermediate-temperature hydrogen compressor. Refiners can add ter.
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• Determine the quantity of circulating water to ensure that 25% of the total water injected in the liquid phase is upstream of the condenser. • Limit the mixed-phase velocities in the piping and tubes to 20 fps for carbon steel and 30 fps for alloys. The minimum velocity should be 10 fps. • Provide symmetrical piping into and out of the condenser. • Ensure that the wash water source is a stripper dedicated to hydrotreating spent wash water if stripped sour water is used. • Ensure that wash water is injected into the main effluent line upstream of the condenser, or independently into each air cooler nozzle through a distribution device like a spray nozzle or restriction orifice. Injecting into piping provides more time for mixing and vaporizing the water and is less expensive. Injection into individual air condenser nozzles ensures good distribution in each air cooler bay. Typical alloys, if required, are Incoloy 800 or 825, and duplex stainless steel 2205.
weight.This is a potential problem for a Table 2 shows that the assumed centrifugal recycle compressor because pressure profile is consistent with 900the lower molecular weight requires psi flanges in the reactor loop’s alloy more head and potentially more stages. piping. The relieving pressure profile This is not as serious for a reciprosets the design pressure for equipment cating machine; however, higher hydro- in the reactor loop. gen purity increases the discharge temEarly in the design process, the deperature, which is limited to 275° F. per signer develops a sized equipment list API-618. that allows for plot plan studies. The Makeup gas normally feeds to the designer checks preliminary reactorrecycle compressor discharge unless it P RESSURE ALLOWANCES Table 1 needs an additionEquipment item Pressure drop, psi al stage of makeup Combined feed exchangers, both sides 40 compression.This Feed heater 50 has the advantage Reactor, fouled 75 HHPS vapor-recycle hydrogen, both sides 20 of maintaining HHPS vapor-CHPS liquid 10 some hydrogen Effluent air condenser 10 Amine system, knock out drums 5 flow through the Piping 25 ––––––––––––––––––––––––––––––––––––– ––– reactor loop in Total reactor loop 235 case of a recycle compressor emergency outage. loop pipe sketches and rechecks allowances for piping pressure drops before specifying the recycle compressor. Reactor-loop hydraulics A hydraulic profile for the reactor The designer checks reactor-loop hydraulics a final time after receiving loop provides the recycle compressor head requirements and establishes the equipment vendor information and afdesign pressure of piping and equipment. ter completing piping isometrics for Compression Reactor inlet pressure is set based on the reactor loop. Inlet flow and head requirements conditions that the catalyst supplier esWe established design pressures in determine the recycle compressor detablishes, and the pressure limits of the reactor loop assuming the CHPS resign. Adequate suction flow is required standard piping flanges. lief valve set pressure is 10% higher than for a centrifugal compressor, and the Two case studies are presented in the normal operating pressure.The rehead should be low enough to limit the Part 2 for reactor inlet pressures of 800 finer can set this margin at 5% with a number of stages so that only a single and 1,100 psig. pilot-operated relief valve, which will body machine is needed; this is usually Table 1 shows a typical allowance for lower the reactor loop’s design pressures. a maximum of 10 stages. reactor-loop equipment and piping A 10% margin in a grassroots facility allows for future capacity increases In the past, most diesel hydrotreaters pressure drop for the 1,100-psig case. had an inadequate volume of recycle Fig. 2 shows a point-to-point pres- while staying within equipment design gas for a centrifugal compressor. New, sure profile for normal operations and pressure limits. API Recommended larger ULSD hydrotreaters that circulate for the relieving case in which the Practice 521, Appendix G recommends 3,000-5,000 cu ft of gas/bbl of feed CHPS is 10% greater than the operata design pressure of 105% of the setare candidates for centrifugal compres- ing pressure. tling-out pressure, which would make sors. Table 2 shows maximum pressures the design pressure about 115% of the A reciprocating compressor used for at relieving conditions and typical denormal operating pressure. recycle normally also provides makeup sign temperatures for critical piping While the designer is establishing service in a multiple-throw machine. services in the reactor section comequipment design conditions, insuffiFor reciprocating compressors greater pared to a maximum allowable pressure cient information exists to calculate the than 500 hp, two 60% capacity mafor 900-psi, type-321 stainless flanges. settling-out pressure accurately. When chines are typically used. For smaller compressors, two 100% units are used. P IPING CONDITIONS Table 2 Makeup compressor discharge feeds to the recycle compressor suction or Pipe section Design pressure, Design temperature, Maximum allowable psi °F. pressure, psi discharge. Exchanger to heater 1,252 700 1,260 Makeup gas flowing to the suction Heater to reactor 1,200 775 1,242 increases recycle gas purity and volReactor to exchanger 1,123 825 1,233 ume, and decreases the molecular
the heat and material balances and process flow diagram are completed, however, one can use preliminary equipment sizing information, assumptions about pipe quantities, and the reactor loop’s pressure profile to estimate the settling-out pressure. When actual equipment and piping design information is available, one can re-estimate the settling-out pressure to determine the margin available compared to the CHPS relief valve set point. If the settling-out pressure exceeds the set pressure, a stepwise calculation or dynamic simulation can estimate the relief rate.
Product stripper The product stripper removes H2S and light hydrocarbons from the ULSD product. Light material removal must be high enough to meet a flash-point specification of 140° F. This ensures removal of nearly all H 2S. The stripper requires about 30 trays. Overhead products include sour gas, unstabilized naphtha, and sour water.The design also includes equipment for intermittently water washing the condenser. Steam or vapors from a fired reboiler serve as the stripping medium. Due to temperature constraints of about 700° F. maximum, the stripper operating pressure is limited to 40-50 psig for a fixed reboiler. Unless the refinery has a low-pressure gas recovery system, a small compressor and spare are needed to handle the offgas. A steam stripper can operate at pressures greater than 100 psig, which allows routing into an existing refinery sour fuel gas system. Diesel product from a steam stripper must have all water removed.Vacuum drying or coalescing followed by salt drying can accomplish the water removal. If this type of system already exists, the steam-stripping option has a lower capital cost vs. a fired reboiler. A steam-stripped tower is also a little smaller than a reboiled tower. If the product drying system is not in place, the refiner should consider the fired reboiler option. It will also create less sour water to process. Product stripper material is killed carbon steel with 410 stainless steel trays
The authors Leslie J. Harwell is a senior consulting process engineer at Mustang Engineers & Constructors LP, Houston. He is primarily involved in ultra-low sulfur projects and midstream Feed filtration gas processing. Harwell has also Although distillate stocks are relatively clean, feed filtration is important worked for Fish Engineering Corp. and Litwin Engineering & Construction Inc. to mitigate exchanger and reactor plug- He holds a BS (1967) in chemical engineering ging. A cartridge filter with 25-m re- from the University of Texas,Austin. Harwell is a tention is typical for this application. registered professional engineer in Texas and a Cracked feeds should feed the hy member of the Gas Processors Suppliers Associadrotreater hot from the upstream facili- tion. ties or from gas blanketed storage. ✦ Shrikant (Sam) Thakkar is a senior specialist for ChevronReferences Texaco Corp., Bellaire,Tex. He 1. Palmer, R.E., Ripperger, G.L., previously worked for Mustang Migliavacca, J.M., “Revamp your hyEngineers & Contractors LP drotreater to manufacture ultra low sul(when this article was written). fur diesel fuel,” presented to the 2001 He holds a BS (1981) in chemical engineering from NPRA Annual Meeting, Mar. 18-20, Bombay University. New Orleans, paper AM-01-22. 2. Leliveld, B., Mayo, S., Miyauchi,Y., and Plantenga, F., “Elegant solutions for Stan Polcar is a principal process engineer for Mustang ultra low sulfur diesel,” presented to the Engineers & Constructors LP. 2001 NPRA Annual Meeting, Mar. 18He has also worked for Litwin 20, New Orleans, paper AM-01-09. Engineering & Construction 3. Brevoord, E., Gudde, N., Hoekstre, Inc. and The Pritchard Corp. Polcar holds a BSc (1968) in G., Mayo, S., and Plantenga, F., “ULSD chemical engineering from Case real life: Commercial performance of Stars and Nebula technology,” present- Western Reserve University, Cleveland. He is a reged to the 2002 NPRA Annual Meeting, istered chemical engineer in Texas. Mar. 17-19, 2002, San Antonio, paper R.E. Palmer is the manager of AM-02-38. downstream process engineering 4. Landue, M.V., Catalysis Today, Vol. for Mustang Engineers & Con36 (1997), pp. 393-429. structors LP. He is responsible 5. Piehl, R.L., “Refiners tame efflufor process design and marketing support for all refining, ent air-cooler corrosion,” OGJ, Aug. 18, petrochemical, and chemical 1975, pp. 119-20. projects. He has led numerous 6.Torrisi, S., “Proven best practices for studies, technology evaluations, and projects relatULSD production,” presented to the ing to clean fuels production. Palmer previously 2002 NPRA Annual Meeting, Mar. 17-19, spent 23 years with Litwin Engineers & Con2002, San Antonio, paper AM-02-35. structors Inc. and 5 years with Conoco Inc. He 7. Schmidt, M., “Premium perform- has a BS in chemical engineering from the Uniance hydrotreating with Axens HR 400 versity of Missouri, Rolla. series hydrotreating catalysts,” presented to the 2002 NPRA Annual Meeting, Pankaj H. Desai is the new business development manager Mar. 17-19, 2002, San Antonio, paper for Akzo Nobel Catalysts LLC, AM-02-57. Houston. He joined Akzo Nobel 8. Larkins, R.P., White, R.R., and Jef- in 1980 and has held various fery, D.W., “Two Phase Concurrent Flow positions in hydroprocessing in Packed Beds,” AICHE Journal, June and FCC catalyst research and development. Desai holds a 1961, pp. 231-39. BTech degree (1974) in chemical engineering 9. Knudsen, K.G, Cooper, B.H., and Topsoe, H., “Catalyst and Process Tech- from Indian Institute of Technology, Kanpur, and a PhD in chemical engineering from the University nologies for Ultra Low Sulfur Diesel,” of Houston. Applied Catalysis A: General, Vol. 189 (1999), pp. 205-15. and caps.The overhead condenser and receiver are hydrogen-induced-cracking resistant steel.The overhead receiver 1 boot has a ⁄ 4-in. corrosion allowance.
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ULSD HYDROTREATING— Conclusion
A recent study for a grassroots 30,000-b/sd hydrotreater that can produce ultra-low sulfur diesel (ULSD) helped assess the critical process Refining design issues for meeting future ULSD requirements. Part 1 (OGJ, July 28, 2003, p. 50) of this two-part series discussed alterna-
Study identifies optimum operating conditions for ULSD hydrotreaters Les Harwell Sam Thakkar Stan Polcar R.E. Palmer Mustang Engineers & Constructors LP Houston Pankaj H. Desai Akzo Nobel Catalysts LLC Houston
tive process configurations for ULSD hydrotreaters with a discussion of key design parameters and metallurgical considerations for major pieces of equipment. Engineering aspects of hydrotreater design were covered, includBased on a presentation to the 2003 NPRA An nual Meeting, Mar. 23-25, 2003, San Antonio.
ing a suitable operating pressure level that satisfies reaction conditions and the practical limits of piping mechanical design. Part 2 covers process simulations for a number of process configurations for reactor inlet pressures from 800 to 1,100 psi. We estimated capital costs and life-cycle costs for three of the most promising cases. The more important conclusions of this study are: • A grassroots hydrotreater for ULSD can have a reactor operating pressure as low as 800 psi if high-purity makeup hydrogen is used.This approach lacks the robustness of a 1,100-psi facility especially in its ability to handle upsets. This disadvantage, which is difficult to quantify, has estimated capital and a life-cycle cost advantages of 10% and 12%, respectively. • A good feedstock characterization including off-design variations is essential for proper selection of catalyst, reaction conditions, and process configuration. We recommend pilot testing. • Capital and operating costs are
ULSD HYDROTREATER, CHPS ONLY
Fig. 1
Makeup H2 LCO
Feed heater Makeup H2 compression
Straight-run diesel Reactor Recycle compressor Quench
Amine treating
Feed ULSD product
Gas to fuel
Cold highpressure separator
Sour water Sour naphtha Steam
Wash water
Sour water
lower for a two-separator F EED PROPERTIES and, therefore, increases the preTable 1 system—hot high-presdicted cycle length for a fixed Straight-run sure separator (HHPS) quantity of catalyst. Table 3 Property gas oil Light cycle oil Blend shows the results with a conand cold high-pressure Flow rate, 1,000 b/sd 20 10 30 Gravity, °API 34.5 20.0 29.4 separator (CHPS)—comstant cycle length of 23-25 Sulfur, wt % 0.8 2.0 1.2 pared to a single CHPS months and an LHSV that N, ppm (wt) 100 500 242 Total aromatics, vol % 20.0 65.0 35.0 design. changes according to reactor inMonoaromatics, vol % 12.8 57.2 27.6 • Minimal chemical let pressure. As the reactor inlet Polyaromatics, vol % 7.2 7.8 7.4 Bromine number 1 10 4.2 hydrogen consumption, pressure increases from 700 to D86 90 vol % distillation temperature, °F. 640 640 640 in conjunction with a 1,100 psi, LHSV increases 35%, Cetane index 49.9 30.3 42.6 process design that elimiwhich means that less catalyst is nates reactor loop purge needed. and minimizes solution For both cases, the makeup ULSD REACTION CONDITIONS, CONSTANT LHSV Table 2 losses, can significantly hydrogen is 85 mole %. ChemiReactor inlet pressure, psig 700 900 1,100 reduce the project’s lifecal hydrogen consumption only Start of run WABT, °F. Base Base – 7° F. Base – 13° F. increases 14-15% for the two cycle cost. Chemical hydrogen • Operating cost is the extremes in reactor inlet presconsumption, cu ft/bbl Base Base x 1.075 Base x 1.14 Cycle length, months 20-22 29-31 38-40 major component in unit sure. This increase, however, can Product sulfur, ppm (wt) 10 10 10 Product nitrogen, ppm (wt) <10 <10 <10 life-cycle cost. have a significant impact on the Yields, wt % • At different pressures project’s life-cycle cost, especialStart of run Feed 100.00 100.00 100.00 C and lighter 0.07 0.07 0.07 4 (700-1,100 psig) and ly if the refinery is hydrogen Naphtha 3.32 3.27 3.24 feed and process condilimited. Distillate 96.04 96.15 96.22 End of run Feed 100.00 100.00 100.00 tions, the increase in For ULSD applications, the C4 and lighter 0.17 0.17 0.16 Naphtha 3.54 3.46 3.41 chemical hydrogen conquantity of cracked stocks in the Distillate 95.78 95.97 96.11 sumption is a modest feed strongly influences chemi15%. For ULSD designs, cal hydrogen consumption. Hyhydrogen consumption is ULSD REACTION CONDITIONS, CONSTANT CYCLE LENGTH Table 3 drogen solution loss (nonchemical consumption) increases mainly a function of feed Reactor inlet pressure, psig 700 900 1,100 characteristics, particularly with pressure. –1 Reactor LHSV, hr Base Base x 1.20 Base x 1.35 aromatics content, and Tables 2 and 3 LHSVs were Start of run WABT, °F. Base Base + 5° F. Base + 7° F. unit operating pressure. 1.0-1.75 hr–1, which is rather Chemical hydrogen consumption, cu ft/bbl Base Base x 1.08 Base x 1.15 • Makeup hydrogen modest with a feed that contains Cycle length, months 23-25 23-25 23-25 Product sulfur, ppm (wt) 10 10 10 purity has a major impact a substantial LCO fraction.This Product nitrogen, ppm (wt) <10 <10 <10 on the recycle gas comis due to a somewhat low feedYields, wt % Start of run Feed 100.00 100.00 100.00 pressor. For a makeup hynitrogen content, a processing C4 and lighter 0.07 0.07 0.07 drogen purity of 85 mole objective of lower product sulNaphtha 3.30 3.29 3.28 Distillate 96.05 96.12 96.18 %, the centrifugal recycle fur only and not cetane imEnd of run Feed 100.00 100.00 100.00 C and lighter 0.17 0.16 0.17 compressor requires two provement, and the application 4 Naphtha 3.53 3.46 3.42 or three stages. For highof new generation, high-activity Distillate 95.8 96.0 96.1 purity hydrogen makeup, catalysts. eight or nine stages are required. It is Tables 2 and 3 show various reactor If the feed contains more nitrogen, impractical to design for this range of parameters and product yields and which is common with coker light gas purities. qualities using Akzo Nobel Catalysts oil, the volume of catalyst required to • Recycle gas purification via recon- LLC’s Stars KF-848 high-activity, nickel- meet the target sulfur specification intact with liquid from a cold low-presmoly catalyst for different reactor inlet creases significantly. Also, changing feed sure separator (CLPS) is not justified pressures and hydrogen partial pressulfur quantities and sulfur species based on energy savings. For an existsures at the reactor outlet. The catalyst might move the design to a more coning unit, however, higher-purity recycle is dense loaded into the reactor. servative LHSV. gas can extend cycle length or help Table 2 shows the different cycle We further analyzed reactor yields meet the treating specification. lengths at each pressure level to obtain and process conditions for a LHSV of a 10-ppm (wt) sulfur product given a 1.25 hr–1 and reactor inlet pressures of Case studies constant liquid hourly space velocity 800 and 1,100 psi. We also considered Table 1 shows the feed characteris(LHSV). Cycle length varies from 20-22 other design options, including: tics used for this study. The feed is a months at a 700-psig reactor pressure • A two-separator system with an blend of two-thirds straight-run diesel to 33-35 months at 1,100 psig. HHPS and CHPS. and one-third FCC light cycle oil The higher hydrogen partial pres• A single-separator system with a (LCO); the combined nitrogen content sure corresponding to each reactor inlet CHPS only. is 242 ppm (wt). pressure decreases catalyst deactivation • An HHPS-CHPS system that en-
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riches recycle hydrogen the various options. For using a CLPS and recon- M AKEUP HYDROGEN Table 4 each case, the LHSV and tacts the liquid with re- COMPOSITION product sulfur content were constant. We set actor effluent that feeds Component Fraction, vol % the products condenser. the recycle hydrogen Hydrogen 85.00 Nitrogen 1.00 • High-purity hyrate to achieve a ratio of Methane 3.21 drogen makeup for the total hydrogen to the Ethane 3.10 Propane 3.80 two-separator case. reactor to chemical hyi-Butane 1.27 • High-purity hydrogen consumption of n-Butane 1.10 i-Pentane 0.64 drogen makeup for the 5.5 for the 1,100-psi n-Pentane 0.32 C6+ 0.56 single separator case. case, and 6.0 for the Total 100.00 Table 4 shows the 800-psi case. low-purity makeup hyBased on the calcudrogen composition. We assumed that lated reactor outlet hydrogen partial high-purity makeup gas was 99.9 mole pressure, Akzo Nobel estimated the cycle length, chemical hydrogen con% hydrogen. The low-purity makeup is typical of sumption, and weighted average bed a blend of hydrogen from a catalytic re- temperature to calculate the final reacformer and a high-purity source, such tor yields and other process conditions. as membrane purification or purchased hydrogen. Case descriptions Table 5 shows the results of simulaWe used the simulation results to tions for the 13 cases corresponding to prepare heat and material balances,
equipment sizes, and design characteristics for three cases: • Case 1 base case: 1,100-psi reactor inlet with a HHPS, CLPS, and 85-mole % makeup hydrogen purity. • Case 2: 1,100-psi separator with a CHPS only and 85-mole % makeup hydrogen purity. • Case 13: 800-psi reactor inlet with a HHPS, CLPS, and 99.9-mole % makeup hydrogen purity. We specified a single, three-bed reactor with inter-bed hydrogen quench for all the cases. We used a centrifugal compressor for recycle hydrogen and two-stage, reciprocating compressor with a spare for makeup.The charge heater has two passes and steam generation in the convection section. The product stripper uses steam for stripping rather than a fired reboiler. We prepared capital cost estimates for each of the three cases and estimat-
SIMULATION RESULTS Case
Table 5
1
2
Reactor inlet pressure, psig 1,100 1,100 Makeup H2 purity, mole % 85 85 Number of separators 2 1 Cold high-pressure separator Temperature, °F. 110 110 Pressure, psig 965 993 H2 partial pressure at reactor outlet, psia 612 674 Cycle length, months 33-35 36-38 Chemical H2 consumption, cu ft/bbl 400-500 400-500 Makeup H2 flow, MMcfd 19.4 18.8 H2 at reactor inlet: Chemical H2 consumption, ratio 5.5 5.5 Total H2 at reactor inlet, cu ft/bbl 2,673 2,673 H2 purity at reactor inlet, mole % 70.2 76.5 Purge, MMcfd 1 1 Enrichment liquid, b/sd 0 0 Recycle gas Flow rate, MMcfd 114.5 104.4 H2 purity, mole % 67.0 75.0 Molecular weight 8.8 6.6 Recycle gas compressor Normal bhp 1,902 1,741 Number of stages 2 3 Stripper Offgas flow, MMcfd 4.0 3.5 Overhead liquid flow, b/sd 1,117 1,140 Overhead liquid cut point, °F. 350 350 ULSD product Flow, b/sd 29,841 29,856 Gravity, °API 33.0 33.1 Cetane index 45.5 45.5 Sulfur, ppm (wt) 10 10 Nitrogen, ppm (wt) <10 <10
3
4
5
6
7
8
9
10
11
1,100
1,100
1,100
1,100
1,100
85
85
85
99.9
3
3
3
110 965
110 965
644
12
13
800
800
800
800
800
800
99.9
85
85
85
85
99.9
99.9
1
2
1
1
1
2
1
2
110 965
110 993
110 965
110 693
110 693
110 693
110 665
110 693
110 665
661
687
898
893
448
492
508
408
612
611
34-36
36-38
37-39
48-50
47-49
18-20
22-24
22-24
<20
33-35
33-35
400-500
400-500
400-500
400-500
400-500
400-500
400-500
400-500
400-500
19.8
19.8
19.8
15.3
15.3
17.0
20.9
25.9
17.3
13.8
14.5
5.5
5.5
5.5
5.5
5.5
6.0
6.0
6.0
6.0
6.0
6.0
2,673
2,673
2,668
2,668
2,668
2,622
2,622
2,622
2,622
2,622
2,622
73.5 1
75.2 1
77.9 1
98.5 0
98.1 0
72.2 1
79.9 5
82.7 10
65.2 1
97.9 0
98.0 0
3,126
6,117
11,756
0
0
0
0
0
0
0
0
109.0 71.0 7.9
106.6 73.0 7.4
103.0 76.0 6.7
81.3 98.1 2.3
81.6 97.6 2.5
109.5 69.4 7.5
98.0 79.0 6.3
95.0 82.0 5.9
122.1 61.0 9.7
80.4 97.4 2.4
80.5 97.3 2.6
1,769 3
1,502 2
1,725 3
1,368 9
1,382 8
2,503 4
2,311 5
2,238 5
2,760 3
1,405 9
1,787 8
3.62
3.4
3.1
1.3
2.1
2.9
2.7
2.7
3.0
0.9
1.4
1,080
1,046
1,000
942
971
1,131
1,182
1,226
1,101
969
995
350
350
350
350
350
350
350
350
350
350
350
29,821 33.1 45.5 10 <10
29,814 33.1 45.5 10 <10
29,810 33.1 45.5 10 <10
29,919 33.1 45.5 10 <10
29,859 33.1 45.5 10 <10
29,834 32.5 44.7 10 <10
29,829 32.5 44.7 10 <10
29,824 32.5 44.7 10 <10
29,837 32.5 44.7 10 <10
29,885 32.6 44.8 10 <10
29,882 32.6 44.8 10 <10
400-500 400-500
the ULSD product LIFE-CYCLE COSTS Table 7 is steam stripped, Case 1 2 13 Case 1 2 however, Case 2 Reactor pressure, psi 1,100 1,100 800 Reactor inlet pressure, requires significant Arrangement 2 separators 1 separator 2 separators psig 1,100 1,100 additional heat exHHPS Yes No Utility consumption CLPS Yes Yes change to raise the Power, kw 3,500 3,418 2,666 Recycle gas flow rate, Medium pressure steam, lb/hr 2,770 -9,000 4,400 cu ft/bbl 3,803 3,493 CHPS liquid temBoiler feedwater, lb/hr 8,000 20,890 6,160 Recycle gas flow rate, Fuel, MMbtu/hr 33.8 70.2 25.3 perature for effecMMcfd 114.5 104.4 Cooling water, gpm 1,354 1,500 1,000 Recycle gas horsepower 1, 902 1, 741 tive H2S removal. Total power consumption, First year costs, million $ This practically kw-hr 3,500 3,418 Utilities 2.3 3.0 1.8 Charge heater fired duty, doubles the charge Maintenance 1 1.0 1.0 0.8 MMbtu/hr 33.8 70.2 2 Catalyst, chemicals 14.7 14.2 13.0 Investment, million $ 62.9 71.1 heater’s size, Capital costs, million $3 62.9 71.1 56.7 which increases Discounted operating costs 4 244.4 256.4 214.2 ed maintenance costs and the conthe size of the fiLife-cycle costs 314.9 336.5 278.1 sumption of utilities, catalysts, and nal effluent air 1 Turnaround costs every 4 years are included in the life-cycle cost but not annual 2 3 maintenance costs. Catalyst and chemicals include hydrogen costs. Capital cost inchemicals. We discounted the operating cooler. 4 cludes initial supplies of catalysts, chemicals, and spare parts. Discounted value of costs over 20 years and added them to Table 6 summaoperating costs based on 15% interest rate and 20-year project life. the investment to calculate the overall rizes the results of life-cycle cost for each case.The disthis analysis. Case 1 has an investment ments, horsepower savings allow an incount was based on an assumed corpo- and operating cost advantage vs. Case 2, vestment of around $110,000 based on rate hurdle rate of 15%. at least for the design feed and other a 3-year simple payout. The equipment Cases 1-5 are based on 1,100-psi re- factors we chose for this study. required for this option may be more actor inlet pressure and low-purity hyCases 3-5 include a CLPS and liquid expensive than the savings. This drogen makeup. Case 1 has an HHPS recontact with HHPS vapor flowing to arrangement might make sense for a for the base option. Case 2 has a single the effluent condenser (Fig. 2). Each revamp in which the recycle compresCHPS (Fig. 1) that results in a highercase shows the impact of more enrich- sor is the bottleneck. purity recycle gas and a corresponding- ment liquid on the recycle compressor Cases 6 and 7 use high-purity hyly lower horsepower for the recycle gas volume and power requirements. drogen for the CHPS, and HHPS-CHPS compressor. The HHPS design is more A comparison of Case 1 and Case 5 arrangements at a 1,100-psi reactor energy efficient, which results in a reshows a maximum advantage of 177 outlet pressure. duction of 36.4 MMbtu/hr of fired du- hp for the enrichment option. IncludThe reactor outlet hydrogen partial ty in the charge heater. ing recycle liquid pump power adjust- pressure increases significantly with a Estimated capital investments for Cases 1 and 2 are Fig. 2 ULSD HYDROTREATER WITH HHPS, CHPS, CLPS $62.9 million and $71.1 milMakeup H2 lion, respectively. These costs LCO are for inside battery limits Feed heater Makeup H2 only and include the initial compression Straight-run load of catalyst, chemicals, and diesel spare parts. Reactor Construction labor rates Recycle and productivity are based on compressor Amine Quench a US Gulf Coast location. We treating prepared these estimates by obtaining equipment pricing Feed and then applying appropriate ULSD product factors to get a total installed cost. One should use these figures carefully because actual Gas Cold lowfacility location, owners’ costs, pressure Gas to fuel separator and design preferences can substantially affect final cost. Cold highHot highpressure Intuitively, Case 1 with two pressure separator separators and an additional separator Sour water HHPS vapor-recycle hydrogen heat exchanger should require Sour naphtha Steam more capital than the single Sour water Wash water CHPS case (Case 2). Because
HHPS ADVANTAGES
Table 6
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charge duty, produces more Fig. 3 ULSD 20-YEAR COSTS medium-pressure steam than 400 is required for the product $ 350 striper. n o i 300 Overall, the 800-psi case l l i ISBL capital cost with two separators has the m250 , t Operating costs, s lowest life-cycle costs, which o 200 discounted to c we expected; however, this e 150 net present value l c case barely meets the mini y c Total life-cycle cost - 100 e mum hydrogen partial pres f (operating and capital) i L 50 sure target even with high0 purity makeup hydrogen. Case 1 Case 2 Case 13 Many refiners installed grassroots hydrotreaters, with reactor inlet operating pressures of Fig. 4 ULSD 20-YEAR OPERATING COSTS around 800 psi, to satisfy the 1993 160 low-sulfur diesel regulations of 500Case 1 140 ppm (wt) sulfur. Owners will revamp Case 2 $ most of these facilities to produce n 120 Case 13 o ULSD even though the hydrogen partial i l l i 100 pressures are significantly less than the m , t minimum recommended for grassroots s o 80 c ULSD hydrotreaters. This will require e l c 60 proportionately more catalyst and per y c haps shorter cycle lengths. e f i 40 L For the 1,100-psi cases, the two-separator arrangement has the lowest life20 cycle cost. 0 Fig. 3 shows that operating costs, Utilities Maintenance Catalyst Chemicals even when discounted over the proincluding H2 ject’s life, are the major contributor to corresponding improvement in cycle even lower than the CHPS-only cases. total life-cycle cost for the three cases. length to 47-49 months from 33-35 Cases 12 and 13 use an 800-psi reThis stresses the importance of an optimonths for Case 7 vs. Case 1. actor outlet pressure with high-purity mum process design and minimized Cases 6 and 7 have a compressor hydrogen makeup, and single and two- utility and maintenance costs. Choosing horsepower almost 30% lower than separator options.The hydrogen partial the right reactor operating conditions Case 1; however, due to a substantial pressure is just greater than the minialso can reduce chemical hydrogen decrease in molecular weight, the com- mum. Estimated US Gulf Coast capital consumption. Fig. 4 compares the operating cost pressor head more than doubles, which cost for Case 13 is $56.7 million. necessitates an 8-stage centrifugal comcomponents. Chemical hydrogen conpressor vs. a 2-stage machine for Case 1. Life-cycle cost sumption has a larger impact than othCases 8-13 have a reactor inlet presWe prepared a life-cycle cost analysis er operating costs. sure of 800 psi. Cases 8-11 use lowfor Cases 1, 2, and 13. A process design that minimizes or purity makeup hydrogen. For this analysis, life-cycle costs are eliminates the need for a purge stream The hydrogen partial pressure is 448 the sum of the capital investment—in- from the recycle gas results in less psia for Case 8, which is far less than cluding initial catalyst, chemicals, and makeup hydrogen needed and a lower Akzo Nobel’s recommended minimum spare parts—and the present value of life-cycle cost. One should also considof 600 psia for a grassroots ULSD hyoperating costs during the project’s life, er the economics of hydrotreating indrotreater. excluding operating labor, taxes, and crementally cracked stocks for ULSD vs. For Cases 9 and 10, we increased the insurance. other uses such as heating oil, cutter reactor loop purge to 5 and 10 MMcfd, We used a discount rate based on an stock, or fuel. respectively, to try to increase the hyassumed corporate hurdle rate of 15% The hydrogen cost in Fig. 4 is drogen partial pressure. The partial and a 20-year project life for this calcu- $2.40/Mcf. We assumed that the refinpressure in Case 10 improved to 508 lation. Table 7 shows the cost results in- ery is hydrogen limited and any repsia at the high purge rate, still signifi- cluding estimates for maintenance costs former hydrogen the refiner uses backs cantly below the minimum value. and utility, catalyst, and chemical conout an equivalent quantity of purchased Case 11 shows that, for a HHPS, the sumption. material. ✦ hydrogen partial pressure of 408 psia is The Case 2 design, due to a higher
