A recent study for a grassroots 30,000-b/sd hydrotreater that can pro- duce ultra-low sulfur diesel (ULSD) helped as- sess the critical process design issues for meeting future ULSD requirements. The feed, a 70:30 mixture of virgin distillate and FCC light cycle oil, was the basis for a previous Mustang Engi- neers & Constructors LP study for re- vamping an existing facility to meet the ULSD specification. 1 This first of two articles presents a comprehensive overview of design con- siderations for grassroots ULSD hy- drotreaters. The chemistry of middle distillate sulfur species (see box), reac- tion process variables, and major fac- tors influencing reactor design are cov- ered. Engineering aspects of hydrotreater design are discussed, including a suit- able operating pressure level that sat- isfies reaction conditions and the practical limits of piping mechanical design. Alternative process configurations for ULSD hydrotreaters were studied, along with key design pa- rameters and metallur- gical considerations for major equipment. Part 2, next week, will present process simulations for different process configurations at reactor inlet pressures of 800 and 1,100 psi.We es- timated capital costs and life cycle costs for three cases. Design considerations The ULSD hydrotreater designer must account for these factors during the front-end design process: • Feed characteristics and variability. • Other product-quality require- ments, especially cetane index. • Catalyst selection. • Optimization of reactor process variables. • Equipment design requirements. • Reliability. • Future off-road diesel sulfur stan- dards. • Minimizing product contamination. • Handling of off-specification diesel product. Process flow Fig. 1 shows a simplified process flow diagram for a diesel hydrotreater. Stripper bottoms heats fresh feed from the surge drum, which then mix- es with recycle hydrogen. Reactor efflu- ent further heats the combined feed, which is then heated in the charge heater. Reactor inter-bed quench may be re- quired depending on the volume of Study outlines optimum ULSD hydrotreater design P ROCESSING Les Harwell Sam Thakkar Stan Polcar R.E. Palmer Mustang Engineers & Constructors LP Houston Pankaj H. Desai Akzo Nobel Catalysts LLC Houston Refining Based on a presentation to the 2003 NPRA An- nual Meeting, Mar. 23-25, 2003, San Antonio. 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 sev- eral strategies to comply with this rule: • Revamping existing facilities commis- sioned 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, thio- phenes, and aromatic sulfur compounds. Sterically hindered dibenzothio- phenes are a group of aromatic sulfur compounds that are most difficult to re- move; they constitute a large fraction that remains in diesel after hydrotreat- ing to the current 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 concentration of aromatic sulfur compounds. Effective removal of these species re- quires tailored catalysts and process conditions. Refiners must also consider other factors, such as feed nitrogen con- tent and aromatics equilibrium. These issues have been addressed in detail in the literature. 2-7 ULSD HYDROTREATING—1 Reprinted with revisions to format, from the July 28, & August 4, 2003 editions of OIL & GAS JOURNAL Copyright 2003 by PennWell Corporation
Transcript
A recent study for agrassroots 30,000-b/sdhydrotreater that can pro-duce ultra-low sulfurdiesel (ULSD) helped as-sess the critical processdesign issues for meeting
future ULSD requirements.The feed, a 70:30 mixture of virgin
distillate and FCC light cycle oil, wasthe basis for a previous Mustang Engi-neers & Constructors LP study for re-vamping an existing facility to meet theULSD specification.1
This first of two articles presents acomprehensive overview of design con-
siderations for grassroots ULSD hy-drotreaters. The chemistry of middledistillate sulfur species (see box), reac-tion process variables, and major fac-tors influencing reactor design are cov-ered.
Engineering aspects of hydrotreaterdesign are discussed, including a suit-able operating pressure level that sat-isfies reaction conditions and thepractical limits of piping mechanicaldesign.
Alternative processconfigurations forULSD hydrotreaterswere studied, alongwith key design pa-rameters and metallur-gical considerationsfor major equipment.
Part 2, next week, will presentprocess simulations for differentprocess configurations at reactor inletpressures of 800 and 1,100 psi. We es-timated capital costs and life cycle costsfor three cases.
dards.• Minimizing product contamination.• Handling of off-specification
diesel product.
Process flowFig. 1 shows a simplified process
flow diagram for a diesel hydrotreater.Stripper bottoms heats fresh feed
from the surge drum, which then mix-es with recycle hydrogen. Reactor efflu-ent further heats the combined feed,which is then heated in the chargeheater.
Reactor inter-bed quench may be re-quired depending on the volume of
Study outlines optimum ULSD hydrotreater design
P R O C E S S I N G
Les HarwellSam ThakkarStan PolcarR.E. PalmerMustang Engineers & Constructors LPHouston
Pankaj H. DesaiAkzo Nobel Catalysts LLCHouston
Refining
Based on a presentation to the 2003 NPRA An-nual Meeting, Mar. 23-25, 2003, San Antonio.
Ultra-low sulfur diesel
In mid 2006, the US EnvironmentalProtection Agency will begin enforcingrules that mandate a maximum sulfurcontent of 15 ppm (wt) for on-roaddiesel fuel. Refiners are considering sev-eral strategies to comply with this rule:• Revamping existing facilities commis-sioned in the early 1990s that now meetthe low-sulfur diesel specification of 500ppm (wt) sulfur.• Constructing grassroots facilities tomeet the new specification.• Executing various combinations ofnew and revamped facilities.
The problem
Middle distillates contain many sulfurspecies, including mercaptans, sulfides, thio-phenes, and aromatic sulfur compounds.
Sterically hindered dibenzothio-phenes are a group of aromatic sulfurcompounds that are most difficult to re-move; they constitute a large fractionthat remains in diesel after hydrotreat-ing to the current specification 500 ppm(wt). This is particularly true for dieselfuels that contain significant quantitiesof cracked stocks, like FCC light cycleoil, that contain a large concentration ofaromatic sulfur compounds.
Effective removal of these species re-quires tailored catalysts and processconditions. Refiners must also considerother factors, such as feed nitrogen con-tent and aromatics equilibrium. Theseissues have been addressed in detail inthe literature.2-7
ULSD HYDROTREATING—1
Reprinted with revisions to format, from the July 28, & August 4, 2003 editions of OIL & GAS JOURNALCopyright 2003 by PennWell Corporation
cracked stocks—FCC light cycle oil andlight coker gas oil—in the feed. Fig. 1shows recycle gas used as quench.
Reactor effluent heats the com-bined feed and flows to a hot, high-pressure separator (HHPS).Vapor fromthe HHPS heats the recycle gas andstripper charge before cooling in theproduct condenser and entering thecold, high-pressure separator (CHPS).
Wash water, injected upstream of theproduct’s condenser, helps remove am-monium hydrosulfide. HHPS liquidcombines with heated CHPS liquid andflows to the product stripper.
Vapor from the CHPS contacts withamine in a scrubber for H2S removaland flows to the recycle compressor
suction drum. Makeup hydrogen iscompressed and combines with recyclegas in the suction drum.The compres-sor suction drum can purge some recy-cle gas to improve recycle gas hydrogenpurity.The quantity depends on the re-actor’s hydrogen-partial-pressure re-quirements and makeup hydrogen pu-rity.
In the product stripper, superheatedsteam feeds the tower’s bottom andhelps remove H2S. Stripper overheadvapors condense and flow to the strip-per accumulator. Accumulator vaporand liquid, known as wild naphtha, areprocessed in offsite facilities.
Stripped ULSD product supplies heatto the feed stream and then flows to dry-
ing facilities, which can be a coalescer-salt dryer or vacuum drying system.
Feed, product characteristicsSulfur, nitrogen, and aromatics have
the greatest impact on the ULSD facilitydesign. Feed nitrogen has a significantimpact on the operating pressure.Therefiner must remove nitrogen to essen-tially the same level as sulfur to meetthe ultra-low sulfur target.
This means that the catalyst and hy-drogen partial pressure must be consis-tent with a high nitrogen-removal op-eration.
Light coker gas oil and FCC light cy-cle oil normally contain most of thefeed nitrogen. Feed aromatics content
P R O C E S S I N G
Coldhigh-pressure
separatorHothigh-pressure
separator
Recyclecompressor
Reactor
Feed heater
Makeup H2compression
Aminetreating
LCO
Straight run diesel
Feed
Quench
Sour water
Sour water
Sour naphtha
Gas to fuel
Steam
Stripper
ULSD product
Makeup H2
Wash water
ULSD HYDROTREATER Fig. 1
governs chemical hydrogen consump-tion at the low space velocities andhigh hydrogen partial pressures re-quired for ULSD production.
Product cetane or gravity generallydetermines the amount of crackedstocks that a refiner can include in thefeed.
There is a small increase in the grav-ity and cetane index during the hy-drotreating reaction.
If a significant improvement incetane (3-5 numbers) or gravity is re-quired, a multistage design using aro-matics saturation catalysts in the secondstage may be a more economical op-tion.The magnitude of improvement ingravity and cetane determines this de-sign choice.
Design product sulfur is a key target,not only from a process design stand-point, but also for offsite storage andtransfer considerations.
Most new and revamp facility designsare for a product sulfur content of 8-10ppm (wt).
Refiners should pilot test the feed toconfirm reaction process conditions.Testing for variations in feed character-istics, especially the FCC light cycle oiland coker light gas oil’s back-end distil-lation, is important because productseparation in these facilities is notori-ously poor.
This can result in a temporary spikein the most difficult-to-treat sulfurcompounds in the hydrotreater feedand requires a higher reactor tempera-ture.This increases the catalyst deactiva-tion rate and directionally reduces cyclelength.
Reaction process variablesKey reaction process variables include:• Space velocity.• Hydrogen partial pressure.• Makeup hydrogen purity.• Ratio of total hydrogen to chemi-
cal hydrogen consumption.• Cycle length.• Reactor temperature.Refiners can use a nickel-molybde-
num catalyst for feeds that have high aro-matics or nitrogen; an appropriate selec-tion of graded catalysts in the bed’s topwill mitigate reactor pressure buildup.
Refiners must optimize reactor spacevelocity, hydrogen treat gas quantity,
hydrogen partial pressure, and reactortemperature during the process designphase for a given cycle length andtreating severity.
Catalyst deactivation rate falls withincreasing hydrogen partial pressure;therefore, the space velocity can in-crease for a constant cycle length.Thisis, however, at the expense of higherhydrogen consumption.
Another important variable is the ra-tio of total hydrogen supplied to chem-ical hydrogen consumption.This valueshould be 5-6 for feeds with significantquantities of cracked stocks.This ratiois normally less important during theprocess design, but it can control thedesign hydrogen-circulation rate, espe-cially when a refiner exclusively useshigh-purity makeup hydrogen.
When setting the hydrogen partialpressure and total operating pressure,one should be aware that the alloy pip-ing flanges limit the reactor pressure,especially when the feed contains sig-nificant quantities of cracked stocks.This includes piping between the feedexchangers and feed heater, betweenthe heater and reactor, and between thereactor and combined feed exchangers.
Piping is normally a 300-seriesstainless steel; for 600-psi flanges thiscorresponds to a maximum operatingpressure of around 800 psi at the reac-tor inlet. For 900-psi flanges, the reac-tor operating pressure can increase toabout 1,100 psi. These pressures ac-count for relieving conditions.
These criteria should not necessarilygovern a refiner’s selection of operatingpressure; however, pressures barely ex-ceeding these limits and requiring thenext-higher-rated ANSI flange class arecostly.
Actual pipe wall thickness is calcu-lated from design conditions; the limit-ing value is used for the piping specifi-cation in the reactor loop’s alloy sec-tions.
Actual pipe materials can also be316, 321, and 347 grades of stainlesssteel. Type 347 has a higher allowablestress value than the other grades but isnormally more expensive. Availabilityand delivery time for the various stain-less components are also a major con-sideration.
The existing equipment and reactor
loop’s piping limit the hydrogen partialpressure in revamp designs. A highertreat gas rate can increase the hydrogenpartial pressure; this is usually limiteddue to an associated increase in the re-actor loop’s pressure drop and the cor-responding maximum operating pres-sure of system components.
Lower-purity makeup hydrogen re-quires higher hydrogen circulationrates to maintain a target hydrogen par-tial pressure; it may even require apurge stream from the cold separator. Ifmakeup hydrogen purity is too low,there is no combination of recycle rateand purge that will achieve the targetreactor outlet partial pressure.
For a revamp design, increasingmakeup hydrogen purity is the mosteffective way to increase the hydrogenpartial pressure.
Catalyst cycle lengths of 24-36months are typical for new designs.This is because, at some point, factorsother than catalyst activity (such as re-actor pressure drop) will limit the cy-cle. For a fixed space velocity, cyclelength increases with a higher hydro-gen partial pressure.
Maximum reactor outlet temperatureat end-of-cycle catalyst conditions isusually 725-750° F. to avoid aromaticssaturation equilibrium constraints andto maintain product color.The amountof cracked stocks in the feed and thecrude source also influence this tem-perature.
Hydrotreating catalyst performancecorrelations for reactor temperature isusually based on the weighted averagebed temperature. WABT is the reactorinlet temperature plus two thirds of thereactor temperature rise. Quenchinglimits the temperature rise to 40-50° F.in each bed.
For a 50° F. temperature rise and a725° F. maximum reactor outlet tem-perature, the end-of-run WABT is (725-50) + 2⁄3(50) = 708° F.
The start-of-run WABT must be highenough to remove the required amountof sulfur and nitrogen.The catalyst de-activation rate at the design space ve-locity and hydrogen partial pressuredetermines cycle length. During the cy-cle, WABT will increase 30-50° F. withlower deactivation rates at higher hy-drogen partial pressures.
Reactor designA superficial mass velocity set at
2,000-5,000 lb/hr/sq ft determinesthe initial reactor diameter. A value of3,500 allows for a reasonable turn-down rate and some upside allowance.Mass flow includes all hydrocarbonsand hydrogen at the reactor inlet.
The refiner should consider a two-train design if the required reactor di-ameter is larger than could be shopfabricated and shipped to the plantsite—usually 12-14 ft for overlandshipment. For diesel hydrotreaters, thisis normally 30,000-40,000 b/sd foreach reactor train.
Refiners with adequate water accessand dock facilities can consider larger
reactors and larger single-train capaci-ties.
An appropriate correlation8 estimatesthe pressure drop for two-phase flowthrough the reactor.The pressure dropcalculation should also include inter-bed quench.
Catalyst physical properties, includ-ing void fraction and equivalent parti-cle diameter, have a significant impacton reactor pressure drop.To ensuregood distribution in the catalyst beds,most ULSD reactors are dense loadedinstead of sock loaded.The clean pres-sure drop should be 0.5-1.0 psi/ft ofcatalyst bed.
For a required catalyst volume, bedheights are either based on the heat of
reaction and maximum temperaturerise or set at 30-40 ft. The designerthen estimates the clean pressure dropfor each bed.
The designer determines the maxi-mum overall pressure drop when de-signing the bed support and comparesit to the catalyst crush strength, whichis the sum of:
• Fouled bed pressure drop at 175%of calculated clean drop.
• Dead weight of catalyst and sup-port material.
• Coke deposits at 30% of catalystdead weight.
• Liquid holdup.• 15-psi allowance for depressuriza-
tion.
P R O C E S S I N G
Coldhigh-pressure
separatorHothigh-pressure
separator
Recyclecompressor
Reactor
Feed heater
Makeup H2compression
Aminetreating
LCO
Straight run diesel
Feed
Quench
Sour water
Sour water
Sour naphtha
Gas to fuel
Steam
Stripper
ULSD product
Makeup H2
Wash waterPressure, psi
Normal Relieving
1,000
1,176 1,276
965 1,062
978
988
1,0861,1861,096
998
1,154
1,254
1,274
1,1741,120
1,022
1,025
1,123 1,060
963
1,290
∆P = 233
1,190
1,252
1,152
1,200
1,102
1,073
975
1,065
968
1,002 1,100
957 1,054
1,198 1,100
PRESSURE PROFILE Fig. 2
Pressure drops for the feed distribu-tor and redistributors, quench inter-nals, and collector are added to the bedvalues to provide the overall fouled-re-actor pressure drop. For ULSD, the feeddistributor and redistributor design iscrucial for obtaining the target productsulfur level.
Several designs provide internal reac-tor quench to limit the temperaturerise.The normal method is to use hy-drogen-rich recycle gas. Quench gas al-so causes liquid vaporization in the re-actor, which provides an extra heat sinkfor heat removal.
Another less-common approach is torecycle CHPS liquid to the reactor.Thequench limits the reactor bed tempera-ture rise to 40-50° F.
Reactor metallurgy is typically 11⁄4-chrome, 1⁄2-molybdenum with 1⁄8-in. 347stainless overlay or cladding.The metal-lurgy choice often depends on the ves-sel thickness, which affects the cost.Cladding is usually more economicalfor vessels with a wall thickness up toabout 4 in. An overlay is normally usedwith thicker walls.
Feed exchangersFeed exchangers are generally speci-
fied to optimize heat economy; howev-er, enough duty should be allowed sothat the charge heater has a reasonableamount of turndown and overall heat-balance control.
Feed exchange should not be morethan 80-85% of the total duty neededto heat the feed to the reactor tempera-ture at end-of-cycle conditions.
The reactor feed-effluent exchanger’smechanical design requires minimalleakage. Many refiners prefer a pull-through tube bundle with a floatinghead because it is easier to clean; how-ever, the floating head cover is a poten-tial leak source. A U-tube bundle avoidsthis potential problem and is less ex-pensive.
Tubes should be seal welded to thetube sheet to prevent leaks from therolled tube joints.
Material selection for feed-effluentexchangers is based on predicted corro-sion rates. Feed-effluent exchangershave several shells in each train withvarying metallurgies. Cold shells are allcarbon steel. Intermediate-temperature
units normally have 11⁄4-chrome shellsand 400-series tubes. Hot shells are 11⁄4-chrome with 300-series stainlesscladding and 300-series stainless tubes.
Charge heaterThe feed heater’s normal duty
should not be less than 20% of thecombined feed duty to the reactor’s in-let temperature at end-of-cycle condi-tions; this allows for turndown andheat balance control. A design marginof 10% of the combined feed exchang-er duty accommodates accelerated ex-changer fouling.
The number of heater passes shouldmatch the number of parallel feed-ef-fluent exchanger trains.The refiner canflow control feed and recycle hydrogenfor each train upstream of the exchang-ers. This prevents maldistribution of thetwo-phase stream entering the heaterpasses.
The limit for one heater pass is15,000-18,000 b/sd of fresh feed attypical diesel hydrotreater conditions.This limit is based on a maximum 8-in.diameter heater tube and an overallpressure drop of 40-50 psig.
Heater tube metallurgy is typically347 stainless steel.
Reactor-product separationSeveral alternative arrangements can
handle reactor effluent after the com-bined feed exchangers.
In general, refiners must decidewhether to use a HHPS in addition to aCHPS, or a CHPS only.The HHPS usual-ly operates around 500-550° F. and,therefore, the vapors can heat recyclegas and cold separator liquid.
A HHPS improves overall heat econ-omy and results in smaller productstripper and auxiliary equipment.Thisarrangement also improves oil-waterseparation in the CHPS.
For heavy gas oil hydrotreating, ahot separator is necessary for adequateoil-water separation.This is less of aconcern for diesel hydrotreating unlessthere are significant quantities ofcracked materials, which decrease APIgravity.
Disadvantages of this design are a 5-10% lower recycle gas purity that re-quires more horsepower in the recyclehydrogen compressor. Refiners can add
a HHPS to debottleneck an existing hy-drotreater that is hydraulically limitedin the reactor loop.This revamp wouldrequire, however, a modest increase ininlet flow and power for the recyclecompressor.
Other options include adding hotand cold low-pressure separators. Thesedesigns marginally unload the upperpart of the stripper and improve LPGrecovery.
HHPS separator base material is 11⁄4-chrome 1⁄2-molybdenum, normally witha clad or overlay lining, which is 300or 400-series stainless depending onthe actual predicted corrosion rate.
The CHPS is carbon steel resistant tohydrogen-induced cracking. A 300-se-ries stainless steel mesh pad helps coa-lesce the wash water.The boot is alsohydrogen-induced-cracking resistantsteel with a 1⁄4-in. corrosion allowance.
Reactor products condenserA final condenser provides cooling
before liquid and vapor separation re-gardless of the reactor product separa-tor configuration. An air cooler is typi-cally used, sometimes supplementedwith a water-trim cooler.
Piping and equipment metallurgy inthis area requires special attention dueto the presence of ammonia and H2S inthe reactor effluent. When this streamcools, these two compounds combineto form ammonium hydrosulfide,which condenses as a solid on inletpiping and in cooler tubes.
Recirculating wash water, injectedbefore the cooler, introduces a corro-sive aqueous solution in the stream toprevent plugging and under-depositcorrosion.
Tube metallurgy selection is basedon ammonia and H2S in the effluent.The designer should ensure a targetmaximum value for Kp of 0.15 beforechoosing alloy metallurgy, where:9
Kp = mole % H2S x mole % NH3The designer should also:• Determine makeup water quantity
to limit ammonium hydrosulfide con-centration in the condensed waterphase to 4 wt %. One must assumeequal molar amounts of ammonia andH2S in the CHPS water and that essen-tially all ammonia dissolves in the wa-ter.
• Determine the quantity of circu-lating water to ensure that 25% of thetotal water injected in the liquid phaseis upstream of the condenser.
• Limit the mixed-phase velocitiesin the piping and tubes to 20 fps forcarbon steel and 30 fps for alloys.Theminimum velocity should be 10 fps.
• Provide symmetrical piping intoand out of the condenser.
• Ensure that the wash water sourceis a stripper dedicated to hydrotreatingspent wash water if stripped sour wateris used.
• Ensure that wash water is injectedinto the main effluent line upstream ofthe condenser, or independently intoeach air cooler nozzle through a distri-bution device like a spray nozzle or re-striction orifice. Injecting into pipingprovides more time for mixing and va-porizing the water and is less expen-sive. Injection into individual air con-denser nozzles ensures good distribu-tion in each air cooler bay.
Typical alloys, if required, are In-coloy 800 or 825, and duplex stainlesssteel 2205.
CompressionInlet flow and head requirements
determine the recycle compressor de-sign. Adequate suction flow is requiredfor a centrifugal compressor, and thehead should be low enough to limit thenumber of stages so that only a singlebody machine is needed; this is usuallya maximum of 10 stages.
In the past, most diesel hydrotreatershad an inadequate volume of recyclegas for a centrifugal compressor. New,larger ULSD hydrotreaters that circulate3,000-5,000 cu ft of gas/bbl of feedare candidates for centrifugal compres-sors.
A reciprocating compressor used forrecycle normally also provides makeupservice in a multiple-throw machine.For reciprocating compressors greaterthan 500 hp, two 60% capacity ma-chines are typically used. For smallercompressors, two 100% units are used.
Makeup compressor discharge feedsto the recycle compressor suction ordischarge.
Makeup gas flowing to the suctionincreases recycle gas purity and vol-ume, and decreases the molecular
weight.This is a potential problem for acentrifugal recycle compressor becausethe lower molecular weight requiresmore head and potentially more stages.
This is not as serious for a recipro-cating machine; however, higher hydro-gen purity increases the discharge tem-perature, which is limited to 275° F. perAPI-618.
Makeup gas normally feeds to therecycle compressordischarge unless itneeds an addition-al stage of makeupcompression.Thishas the advantageof maintainingsome hydrogenflow through thereactor loop incase of a recyclecompressor emer-gency outage.
Reactor-loop hydraulicsA hydraulic profile for the reactor
loop provides the recycle compressorhead requirements and establishes thedesign pressure of piping and equipment.
Reactor inlet pressure is set based onconditions that the catalyst supplier es-tablishes, and the pressure limits ofstandard piping flanges.
Two case studies are presented inPart 2 for reactor inlet pressures of 800and 1,100 psig.
Table 1 shows a typical allowance forreactor-loop equipment and pipingpressure drop for the 1,100-psig case.
Fig. 2 shows a point-to-point pres-sure profile for normal operations andfor the relieving case in which theCHPS is 10% greater than the operat-ing pressure.
Table 2 shows maximum pressuresat relieving conditions and typical de-sign temperatures for critical pipingservices in the reactor section com-pared to a maximum allowable pressurefor 900-psi, type-321 stainless flanges.
Table 2 shows that the assumedpressure profile is consistent with 900-psi flanges in the reactor loop’s alloypiping.The relieving pressure profilesets the design pressure for equipmentin the reactor loop.
Early in the design process, the de-signer develops a sized equipment listthat allows for plot plan studies.Thedesigner checks preliminary reactor-
loop pipe sketches and rechecks al-lowances for piping pressure drops be-fore specifying the recycle compressor.
The designer checks reactor-loophydraulics a final time after receivingequipment vendor information and af-ter completing piping isometrics forthe reactor loop.
We established design pressures inthe reactor loop assuming the CHPS re-lief valve set pressure is 10% higher thanthe normal operating pressure.The re-finer can set this margin at 5% with apilot-operated relief valve, which willlower the reactor loop’s design pressures.
A 10% margin in a grassroots facilityallows for future capacity increaseswhile staying within equipment designpressure limits. API RecommendedPractice 521, Appendix G recommendsa design pressure of 105% of the set-tling-out pressure, which would makethe design pressure about 115% of thenormal operating pressure.
While the designer is establishingequipment design conditions, insuffi-cient information exists to calculate thesettling-out pressure accurately. When
P R O C E S S I N G
Equipment item Pressure drop, psi
Combined feed exchangers, both sides 40Feed heater 50Reactor, fouled 75HHPS vapor-recycle hydrogen, both sides 20HHPS vapor-CHPS liquid 10Effluent air condenser 10Amine system, knock out drums 5Piping 25––––––––––––––––––––––––––––––––––––– –––Total reactor loop 235
Exchanger to heater 1,252 700 1,260Heater to reactor 1,200 775 1,242Reactor to exchanger 1,123 825 1,233
PIPING CONDITIONS Table 2
the heat and material balances andprocess flow diagram are completed,however, one can use preliminaryequipment sizing information, assump-tions about pipe quantities, and the re-actor loop’s pressure profile to estimatethe settling-out pressure.
When actual equipment and pipingdesign information is available, one canre-estimate the settling-out pressure todetermine the margin available com-pared to the CHPS relief valve set point.If the settling-out pressure exceeds theset pressure, a stepwise calculation ordynamic simulation can estimate therelief rate.
Product stripperThe product stripper removes H2S
and light hydrocarbons from the ULSDproduct. Light material removal mustbe high enough to meet a flash-pointspecification of 140° F.This ensures re-moval of nearly all H2S.
The stripper requires about 30 trays.Overhead products include sour gas, un-stabilized naphtha, and sour water.Thedesign also includes equipment for inter-mittently water washing the condenser.
Steam or vapors from a fired reboilerserve as the stripping medium. Due totemperature constraints of about 700°F. maximum, the stripper operatingpressure is limited to 40-50 psig for afixed reboiler.
Unless the refinery has a low-pres-sure gas recovery system, a small com-pressor and spare are needed to handlethe offgas.
A steam stripper can operate at pres-sures greater than 100 psig, which al-lows routing into an existing refinerysour fuel gas system.
Diesel product from a steam strippermust have all water removed.Vacuumdrying or coalescing followed by saltdrying can accomplish the water re-moval. If this type of system already ex-ists, the steam-stripping option has alower capital cost vs. a fired reboiler. Asteam-stripped tower is also a littlesmaller than a reboiled tower.
If the product drying system is notin place, the refiner should consider thefired reboiler option. It will also createless sour water to process.
Product stripper material is killed car-bon steel with 410 stainless steel trays
and caps.The overhead condenser andreceiver are hydrogen-induced-crackingresistant steel.The overhead receiverboot has a 1⁄4-in. corrosion allowance.
Feed filtrationAlthough distillate stocks are rela-
tively clean, feed filtration is importantto mitigate exchanger and reactor plug-ging. A cartridge filter with 25-�m re-tention is typical for this application.Cracked feeds should feed the hy-drotreater hot from the upstream facili-ties or from gas blanketed storage. ✦
References1. Palmer, R.E., Ripperger, G.L.,
Migliavacca, J.M., “Revamp your hy-drotreater to manufacture ultra low sul-fur diesel fuel,” presented to the 2001NPRA Annual Meeting, Mar. 18-20,New Orleans, paper AM-01-22.
2. Leliveld, B., Mayo, S., Miyauchi,Y.,and Plantenga, F., “Elegant solutions forultra low sulfur diesel,” presented to the2001 NPRA Annual Meeting, Mar. 18-20, New Orleans, paper AM-01-09.
3. Brevoord, E., Gudde, N., Hoekstre,G., Mayo, S., and Plantenga, F., “ULSDreal life: Commercial performance ofStars and Nebula technology,” present-ed to the 2002 NPRA Annual Meeting,Mar. 17-19, 2002, San Antonio, paperAM-02-38.
4. Landue, M.V., Catalysis Today,Vol.36 (1997), pp. 393-429.
5. Piehl, R.L., “Refiners tame efflu-ent air-cooler corrosion,” OGJ, Aug. 18,1975, pp. 119-20.
6.Torrisi, S., “Proven best practices forULSD production,” presented to the2002 NPRA Annual Meeting, Mar. 17-19,2002, San Antonio, paper AM-02-35.
7. Schmidt, M., “Premium perform-ance hydrotreating with Axens HR 400series hydrotreating catalysts,” present-ed to the 2002 NPRA Annual Meeting,Mar. 17-19, 2002, San Antonio, paperAM-02-57.
8. Larkins, R.P., White, R.R., and Jef-fery, D.W., “Two Phase Concurrent Flowin Packed Beds,” AICHE Journal, June1961, pp. 231-39.
9. Knudsen, K.G, Cooper, B.H., andTopsoe, H., “Catalyst and Process Tech-nologies for Ultra Low Sulfur Diesel,”Applied Catalysis A: General,Vol. 189(1999), pp. 205-15.
The authorsLeslie J. Harwell is a seniorconsulting process engineer atMustang Engineers & Con-structors LP, Houston. He isprimarily involved in ultra-lowsulfur projects and midstreamgas processing. Harwell has alsoworked for Fish EngineeringCorp. and Litwin Engineering & Construction Inc.He holds a BS (1967) in chemical engineeringfrom the University of Texas,Austin. Harwell is aregistered professional engineer in Texas and amember of the Gas Processors Suppliers Associa-tion.
Shrikant (Sam) Thakkar is asenior specialist for Chevron-Texaco Corp., Bellaire,Tex. Hepreviously worked for MustangEngineers & Contractors LP(when this article was written).He holds a BS (1981) inchemical engineering from
Bombay University.
Stan Polcar is a principalprocess engineer for MustangEngineers & Constructors LP.He has also worked for LitwinEngineering & ConstructionInc. and The Pritchard Corp.Polcar holds a BSc (1968) inchemical engineering from CaseWestern Reserve University, Cleveland. He is a reg-istered chemical engineer in Texas.
R.E. Palmer is the manager ofdownstream process engineeringfor Mustang Engineers & Con-structors LP. He is responsiblefor process design and market-ing support for all refining,petrochemical, and chemicalprojects. He has led numerous
studies, technology evaluations, and projects relat-ing to clean fuels production. Palmer previouslyspent 23 years with Litwin Engineers & Con-structors Inc. and 5 years with Conoco Inc. Hehas a BS in chemical engineering from the Uni-versity of Missouri, Rolla.
Pankaj H. Desai is the newbusiness development managerfor Akzo Nobel Catalysts LLC,Houston. He joined Akzo Nobelin 1980 and has held variouspositions in hydroprocessingand FCC catalyst research anddevelopment. Desai holds aBTech degree (1974) in chemical engineeringfrom Indian Institute of Technology, Kanpur, and aPhD in chemical engineering from the Universityof Houston.
A recent study for agrassroots 30,000-b/sdhydrotreater that can pro-duce ultra-low sulfurdiesel (ULSD) helped as-sess the critical processdesign issues for meetingfuture ULSD require-ments.
Part 1 (OGJ, July 28, 2003, p. 50) ofthis two-part series discussed alterna-
tive process configurations for ULSDhydrotreaters with a discussion of keydesign parameters and metallurgicalconsiderations for major pieces ofequipment. Engineering aspects of hy-drotreater design were covered, includ-
ing a suitable operating pressure levelthat satisfies reaction conditions andthe practical limits of piping mechani-cal design.
Part 2 covers process simulations fora number of process configurations forreactor inlet pressures from 800 to1,100 psi. We estimated capital costsand life-cycle costs for three of themost promising cases.
The more important conclusions ofthis study are:
• A grassroots hydrotreater for ULSDcan have a reactor operating pressure aslow as 800 psi if high-purity makeuphydrogen is used.This approach lacksthe robustness of a 1,100-psi facilityespecially in its ability to handle upsets.This disadvantage, which is difficult toquantify, has estimated capital and alife-cycle cost advantages of 10% and12%, respectively.
• A good feedstock characterizationincluding off-design variations is essen-tial for proper selection of catalyst, re-action conditions, and process configu-ration. We recommend pilot testing.
• Capital and operating costs are
Study identifies optimum operatingconditions for ULSD hydrotreaters
P R O C E S S I N G
Les HarwellSam ThakkarStan PolcarR.E. PalmerMustang Engineers & Constructors LPHouston
Pankaj H. DesaiAkzo NobelCatalysts LLC
Houston
Cold high-pressureseparator
Recyclecompressor
Reactor
Feed heaterMakeup H2
compression
Aminetreating
LCO
Straight-rundiesel
Feed
Quench
Sour water
Sour water
Sour naphtha
Gas to fuel
Steam
ULSD product
Makeup H2
Wash water
ULSD HYDROTREATER, CHPS ONLY Fig. 1
Refining
ULSD HYDROTREATING—Conclusion
Based on a presentation to the 2003 NPRA An-nual Meeting, Mar. 23-25, 2003, San Antonio.
lower for a two-separatorsystem—hot high-pres-sure separator (HHPS)and cold high-pressureseparator (CHPS)—com-pared to a single CHPSdesign.
• Minimal chemicalhydrogen consumption,in conjunction with aprocess design that elimi-nates reactor loop purgeand minimizes solutionlosses, can significantlyreduce the project’s life-cycle cost.
• Operating cost is themajor component in unitlife-cycle cost.
• At different pressures(700-1,100 psig) andfeed and process condi-tions, the increase inchemical hydrogen con-sumption is a modest15%. For ULSD designs,hydrogen consumption ismainly a function of feedcharacteristics, particularlyaromatics content, andunit operating pressure.
• Makeup hydrogenpurity has a major impacton the recycle gas com-pressor. For a makeup hy-drogen purity of 85 mole%, the centrifugal recyclecompressor requires twoor three stages. For high-purity hydrogen makeup,eight or nine stages are required. It isimpractical to design for this range ofpurities.
• Recycle gas purification via recon-tact with liquid from a cold low-pres-sure separator (CLPS) is not justifiedbased on energy savings. For an exist-ing unit, however, higher-purity recyclegas can extend cycle length or helpmeet the treating specification.
Case studiesTable 1 shows the feed characteris-
tics used for this study.The feed is ablend of two-thirds straight-run dieseland one-third FCC light cycle oil(LCO); the combined nitrogen contentis 242 ppm (wt).
Tables 2 and 3 show various reactorparameters and product yields andqualities using Akzo Nobel CatalystsLLC’s Stars KF-848 high-activity, nickel-moly catalyst for different reactor inletpressures and hydrogen partial pres-sures at the reactor outlet. The catalystis dense loaded into the reactor.
Table 2 shows the different cyclelengths at each pressure level to obtaina 10-ppm (wt) sulfur product given aconstant liquid hourly space velocity(LHSV). Cycle length varies from 20-22months at a 700-psig reactor pressureto 33-35 months at 1,100 psig.
The higher hydrogen partial pres-sure corresponding to each reactor inletpressure decreases catalyst deactivation
and, therefore, increases the pre-dicted cycle length for a fixedquantity of catalyst. Table 3shows the results with a con-stant cycle length of 23-25months and an LHSV thatchanges according to reactor in-let pressure. As the reactor inletpressure increases from 700 to1,100 psi, LHSV increases 35%,which means that less catalyst isneeded.
For both cases, the makeuphydrogen is 85 mole %. Chemi-cal hydrogen consumption onlyincreases 14-15% for the twoextremes in reactor inlet pres-sure.This increase, however, canhave a significant impact on theproject’s life-cycle cost, especial-ly if the refinery is hydrogenlimited.
For ULSD applications, thequantity of cracked stocks in thefeed strongly influences chemi-cal hydrogen consumption. Hy-drogen solution loss (nonchem-ical consumption) increaseswith pressure.
Tables 2 and 3 LHSVs were1.0-1.75 hr–1, which is rathermodest with a feed that containsa substantial LCO fraction.Thisis due to a somewhat low feed-nitrogen content, a processingobjective of lower product sul-fur only and not cetane im-provement, and the applicationof new generation, high-activitycatalysts.
If the feed contains more nitrogen,which is common with coker light gasoil, the volume of catalyst required tomeet the target sulfur specification in-creases significantly. Also, changing feedsulfur quantities and sulfur speciesmight move the design to a more con-servative LHSV.
We further analyzed reactor yieldsand process conditions for a LHSV of1.25 hr–1 and reactor inlet pressures of800 and 1,100 psi. We also consideredother design options, including:
• A two-separator system with anHHPS and CHPS.
• A single-separator system with aCHPS only.
• An HHPS-CHPS system that en-
Straight-run Property gas oil Light cycle oil Blend
End of run Feed 100.00 100.00 100.00C4 and lighter 0.17 0.17 0.16Naphtha 3.54 3.46 3.41Distillate 95.78 95.97 96.11
ULSD REACTION CONDITIONS, CONSTANT LHSV Table 2
Reactor inlet pressure, psig 700 900 1,100
Reactor LHSV, hr–1 Base Base x 1.20 Base x 1.35Start of run WABT, °F. Base Base + 5° F. Base + 7° F.Chemical hydrogenconsumption, cu ft/bbl Base Base x 1.08 Base x 1.15
P R O C E S S I N Griches recycle hydrogenusing a CLPS and recon-tacts the liquid with re-actor effluent that feedsthe products condenser.
• High-purity hy-drogen makeup for thetwo-separator case.
• High-purity hy-drogen makeup for thesingle separator case.
Table 4 shows thelow-purity makeup hy-drogen composition. We assumed thathigh-purity makeup gas was 99.9 mole% hydrogen.
The low-purity makeup is typical ofa blend of hydrogen from a catalytic re-former and a high-purity source, suchas membrane purification or purchasedhydrogen.
Table 5 shows the results of simula-tions for the 13 cases corresponding to
the various options. Foreach case, the LHSV andproduct sulfur contentwere constant. We setthe recycle hydrogenrate to achieve a ratio oftotal hydrogen to thereactor to chemical hy-drogen consumption of5.5 for the 1,100-psicase, and 6.0 for the800-psi case.
Based on the calcu-lated reactor outlet hydrogen partialpressure, Akzo Nobel estimated the cy-cle length, chemical hydrogen con-sumption, and weighted average bedtemperature to calculate the final reac-tor yields and other process conditions.
Case descriptionsWe used the simulation results to
prepare heat and material balances,
equipment sizes, and design character-istics for three cases:
• Case 1 base case: 1,100-psi reactorinlet with a HHPS, CLPS, and 85-mole% makeup hydrogen purity.
• Case 2: 1,100-psi separator with aCHPS only and 85-mole % makeup hy-drogen purity.
• Case 13: 800-psi reactor inlet witha HHPS, CLPS, and 99.9-mole % make-up hydrogen purity.
We specified a single, three-bed re-actor with inter-bed hydrogen quenchfor all the cases. We used a centrifugalcompressor for recycle hydrogen andtwo-stage, reciprocating compressorwith a spare for makeup.The chargeheater has two passes and steam gener-ation in the convection section.Theproduct stripper uses steam for strip-ping rather than a fired reboiler.
We prepared capital cost estimatesfor each of the three cases and estimat-
ed maintenance costs and the con-sumption of utilities, catalysts, andchemicals. We discounted the operatingcosts over 20 years and added them tothe investment to calculate the overalllife-cycle cost for each case.The dis-count was based on an assumed corpo-rate hurdle rate of 15%.
Cases 1-5 are based on 1,100-psi re-actor inlet pressure and low-purity hy-drogen makeup. Case 1 has an HHPSfor the base option. Case 2 has a singleCHPS (Fig. 1) that results in a higher-purity recycle gas and a corresponding-ly lower horsepower for the recycle gascompressor.The HHPS design is moreenergy efficient, which results in a re-duction of 36.4 MMbtu/hr of fired du-ty in the charge heater.
Estimated capital invest-ments for Cases 1 and 2 are$62.9 million and $71.1 mil-lion, respectively.These costsare for inside battery limitsonly and include the initialload of catalyst, chemicals, andspare parts.
Construction labor ratesand productivity are based ona US Gulf Coast location. Weprepared these estimates byobtaining equipment pricingand then applying appropriatefactors to get a total installedcost. One should use these fig-ures carefully because actualfacility location, owners’ costs,and design preferences cansubstantially affect final cost.
Intuitively, Case 1 with twoseparators and an additionalHHPS vapor-recycle hydrogenheat exchanger should requiremore capital than the singleCHPS case (Case 2). Because
the ULSD productis steam stripped,however, Case 2requires significantadditional heat ex-change to raise theCHPS liquid tem-perature for effec-tive H2S removal.This practicallydoubles the chargeheater’s size,which increasesthe size of the fi-nal effluent aircooler.
Table 6 summa-rizes the results ofthis analysis. Case 1 has an investmentand operating cost advantage vs. Case 2,at least for the design feed and otherfactors we chose for this study.
Cases 3-5 include a CLPS and liquidrecontact with HHPS vapor flowing tothe effluent condenser (Fig. 2). Eachcase shows the impact of more enrich-ment liquid on the recycle compressorvolume and power requirements.
A comparison of Case 1 and Case 5shows a maximum advantage of 177hp for the enrichment option. Includ-ing recycle liquid pump power adjust-
ments, horsepower savings allow an in-vestment of around $110,000 based ona 3-year simple payout.The equipmentrequired for this option may be moreexpensive than the savings.Thisarrangement might make sense for arevamp in which the recycle compres-sor is the bottleneck.
Cases 6 and 7 use high-purity hy-drogen for the CHPS, and HHPS-CHPSarrangements at a 1,100-psi reactoroutlet pressure.
The reactor outlet hydrogen partialpressure increases significantly with a
First year costs, million $Utilities 2.3 3.0 1.8Maintenance1 1.0 1.0 0.8Catalyst, chemicals2 14.7 14.2 13.0
Capital costs, million $3 62.9 71.1 56.7Discounted operating costs4 244.4 256.4 214.2Life-cycle costs 314.9 336.5 278.1
1Turnaround costs every 4 years are included in the life-cycle cost but not annualmaintenance costs. 2Catalyst and chemicals include hydrogen costs. 3Capital cost in-cludes initial supplies of catalysts, chemicals, and spare parts. 4Discounted value ofoperating costs based on 15% interest rate and 20-year project life.
LIFE-CYCLE COSTS Table 7
Cold low-pressureseparator
Cold high-pressureseparator
Hot high-pressureseparator
Recyclecompressor
Reactor
Feed heaterMakeup H2
compression
Aminetreating
LCO
Straight-rundiesel
Feed
Quench
Sour water
Gas
Sour water
Sour naphtha
Gas to fuel
Steam
ULSD product
Makeup H2
Wash water
ULSD HYDROTREATER WITH HHPS, CHPS, CLPS Fig. 2
P R O C E S S I N G
corresponding improvement in cyclelength to 47-49 months from 33-35months for Case 7 vs. Case 1.
Cases 6 and 7 have a compressorhorsepower almost 30% lower thanCase 1; however, due to a substantialdecrease in molecular weight, the com-pressor head more than doubles, whichnecessitates an 8-stage centrifugal com-pressor vs. a 2-stage machine for Case 1.
Cases 8-13 have a reactor inlet pres-sure of 800 psi. Cases 8-11 use low-purity makeup hydrogen.
The hydrogen partial pressure is 448psia for Case 8, which is far less thanAkzo Nobel’s recommended minimumof 600 psia for a grassroots ULSD hy-drotreater.
For Cases 9 and 10, we increased thereactor loop purge to 5 and 10 MMcfd,respectively, to try to increase the hy-drogen partial pressure.The partialpressure in Case 10 improved to 508psia at the high purge rate, still signifi-cantly below the minimum value.
Case 11 shows that, for a HHPS, thehydrogen partial pressure of 408 psia is
even lower than the CHPS-only cases.Cases 12 and 13 use an 800-psi re-
actor outlet pressure with high-purityhydrogen makeup, and single and two-separator options.The hydrogen partialpressure is just greater than the mini-mum. Estimated US Gulf Coast capitalcost for Case 13 is $56.7 million.
Life-cycle costWe prepared a life-cycle cost analysis
for Cases 1, 2, and 13.For this analysis, life-cycle costs are
the sum of the capital investment—in-cluding initial catalyst, chemicals, andspare parts—and the present value ofoperating costs during the project’s life,excluding operating labor, taxes, andinsurance.
We used a discount rate based on anassumed corporate hurdle rate of 15%and a 20-year project life for this calcu-lation.Table 7 shows the cost results in-cluding estimates for maintenance costsand utility, catalyst, and chemical con-sumption.
The Case 2 design, due to a higher
charge duty, produces moremedium-pressure steam thanis required for the productstriper.
Overall, the 800-psi casewith two separators has thelowest life-cycle costs, whichwe expected; however, thiscase barely meets the mini-mum hydrogen partial pres-sure target even with high-purity makeup hydrogen.
Many refiners installedgrassroots hydrotreaters, with
reactor inlet operating pressures ofaround 800 psi, to satisfy the 1993low-sulfur diesel regulations of 500-ppm (wt) sulfur. Owners will revampmost of these facilities to produceULSD even though the hydrogen partialpressures are significantly less than theminimum recommended for grassrootsULSD hydrotreaters. This will requireproportionately more catalyst and per-haps shorter cycle lengths.
For the 1,100-psi cases, the two-sep-arator arrangement has the lowest life-cycle cost.
Fig. 3 shows that operating costs,even when discounted over the pro-ject’s life, are the major contributor tototal life-cycle cost for the three cases.This stresses the importance of an opti-mum process design and minimizedutility and maintenance costs. Choosingthe right reactor operating conditionsalso can reduce chemical hydrogenconsumption.
Fig. 4 compares the operating costcomponents. Chemical hydrogen con-sumption has a larger impact than oth-er operating costs.
A process design that minimizes oreliminates the need for a purge streamfrom the recycle gas results in lessmakeup hydrogen needed and a lowerlife-cycle cost. One should also consid-er the economics of hydrotreating in-crementally cracked stocks for ULSD vs.other uses such as heating oil, cutterstock, or fuel.
The hydrogen cost in Fig. 4 is$2.40/Mcf. We assumed that the refin-ery is hydrogen limited and any re-former hydrogen the refiner uses backsout an equivalent quantity of purchasedmaterial. ✦
