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AMIT KELKAR (Shell Catalysts & Technologies)
Nitrogen slip is a key variable in balancing the performance of the pretreat and cracking catalysts for cycle optimization. The optimum nitrogen slip depends on the specific unit objectives and constraints. In general, the pretreat temperature is adjusted to maintain sufficiently low slip to the cracking catalyst to achieve the target conversion while also balancing the activity to fully utilize both catalyst systems. Lower nitrogen slip results in improved product quality and more volume swell but can lead to faster utilization of the pretreat system which may reach EOR before the cracking catalyst.

Hydrocrackers with separate pretreat and cracking reactors are often equipped with an inter-reactor sample to track N-slip. In units without adequate sampling, top cracking bed dT is used as a real time indicator of nitrogen poisoning. Decreasing top bed dT suggests increasing N slip which in turn requires higher cracking severity to maintain conversion. If the slip is high enough, N inhibition can progress down the lower beds as evident in loss of dT. Nitrogen inhibition is reversible and cracking activity recovers once pretreat WABT is raised. Cracking bed temperatures should be monitored carefully as N-slip is lowered to avoid excessive cracking as activity recovers. It is important to note that N-slip is one of many variables that impacts bed dT and should only be used as a qualitative indicator.
Kinetic modeling and pilot plant testing are useful tools to understand impact of N slip on cracking activity and selectivity. 
Customized catalyst system design is critical in optimizing pretreat and cracking activity for maximum performance. Pretreat limited units are designed with a robust cracking catalyst system that can withstand higher N slip in the later part of the cycle without loss of conversion. A H2 constrained unit might need to be designed with a higher N slip and the appropriate cracking system to meet performance objectives. In some instances, the HCPT WABT is rapidly raised at start of the cycle to the optimum temperature for maximum aromatic saturation and maintained there. This is often the case with highly aromatic feeds like LCO to maximize volume swell.
In addition to WABT, temperature profile is an important handle in balancing pretreat and cracking severity. For pretreat, optimum catalyst utilization is achieved by operating in an equal bed outlet mode so that each bed deactivates at a similar rate. At times, pretreat must be operated in an ascending profile to generate enough heat input for the cracking reactor or to optimize metals uptake. On the other hand, cracking beds should be operated at equal bed dT to ensure similar deactivation. The top bed is exposed to the highest nitrogen inhibition and maintaining equal bed dT means operating in a descending profile. This can lead to higher temperature in the lower beds and possible runaway in case of loss of quench. Hence the recommendation is to target lower dT in the top bed and equal dT’s for the lower beds. 
Understanding the nature of the molecules being converted and specific unit constraints is key to selecting the appropriate catalyst system and managing the temperature profile for overall unit optimization.

ROBERT STEINBERG (Motiva Enterprises)

The pretreat temperature is ideally set to hit a desired nitrogen slip to the cracking catalyst. However, in some circumstances the nitrogen slip cannot be measured. This could be because there are pretreat and cracking beds in the same reactor with no way to sample between beds. Or, even when there is a separate pretreat reactor there may not be facilities to obtain a sample of the effluent before it is mixed with recycle oil to the cracking reactor or effluent from the cracking reactor. In such circumstances the pretreat temperature needs to be set based on cracking catalyst performance.

A typical nitrogen slip from the pretreat reactor would be in the range of 20-60 ppmw but can sometimes be higher or lower. The target depends on how active the cracking catalyst is and the need to balance pretreat and cracking catalyst life. More nitrogen slip allows a reduction in pretreat temperature and extends pretreat catalyst life. Less nitrogen slip increases the activity of the cracking catalyst – this can be used to operate at a lower temperature and extend cracking catalyst life or increase conversion.

The simplest way to operate the pretreat and cracking beds is to monitor their respective weighted average bed temperatures (WABT) and exotherms. If the exotherm in the lead cracking bed decreases at the same inlet temperature, nitrogen slip may have increased and the pretreat WABT should be increased. This principal applies to a single stage unit where there is simply pretreat followed by cracking catalyst and also to a two stage recycle unit – in both cases a higher pretreat WABT will reduce nitrogen slip to the cracking catalyst and increase conversion at the same cracking bed temperature.

When measuring nitrogen slip, it is best to look at the nitrogen content of the unconverted oil (UCO) after naphtha and diesel have been removed. Catalyst vendors will recommend a target nitrogen slip that should work well for their catalysts in a particular unit but this may need to be adjusted as operating conditions change and the catalyst ages.

ROBERT STEINBERG (Motiva Enterprises)
The pretreat temperature is ideally set to hit a desired nitrogen slip to the cracking catalyst. However, in some circumstances the nitrogen slip cannot be measured. This could be because there are pretreat and cracking beds in the same reactor with no way to sample between beds. Or, even when there is a separate pretreat reactor there may not be facilities to obtain a sample of the effluent before it is mixed with recycle oil to the cracking reactor or effluent from the cracking reactor. In such circumstances the pretreat temperature needs to be set based on cracking catalyst performance.

A typical nitrogen slip from the pretreat reactor would be in the range of 20-60 ppmw but can sometimes be higher or lower. The target depends on how active the cracking catalyst is and the need to balance pretreat and cracking catalyst life. More nitrogen slip allows a reduction in pretreat temperature and extends pretreat catalyst life. Less nitrogen slip increases the activity of the cracking catalyst – this can be used to operate at a lower temperature and extend cracking catalyst life or increase conversion.

The simplest way to operate the pretreat and cracking beds is to monitor their respective weighted average bed temperatures (WABT) and exotherms. If the exotherm in the lead cracking bed decreases at the same inlet temperature, nitrogen slip may have increased and the pretreat WABT should be increased. This principal applies to a single stage unit where there is simply pretreat followed by cracking catalyst and also to a two stage recycle unit – in both cases a higher pretreat WABT will reduce nitrogen slip to the cracking catalyst and increase conversion at the same cracking bed temperature.

When measuring nitrogen slip, it is best to look at the nitrogen content of the unconverted oil (UCO) after naphtha and diesel have been removed. Catalyst vendors will recommend a target nitrogen slip that should work well for their catalysts in a particular unit but this may need to be adjusted as operating conditions change and the catalyst ages.

SYED SHAH (Honeywell UOP)
In a hydrocracking unit, the most effective way to optimize the nitrogen slip to the cracking catalyst is by taking a pretreat effluent sample and testing it for nitrogen using ASTM D4629 or ASTM D5762. Knowing the nitrogen slip from the pretreating catalyst allows determining the relative performance and stability between the pretreat and cracking catalyst systems to make optimum use of each. UOP Unicracking design includes a sample point on pretreat reactor effluent that is specifically designed to take a sample of this hot stream. Based on nitrogen slip result, pretreat temperatures are adjusted to maintain the nitrogen slip close to the target required for optimum performance of the cracking catalyst. It is important to balance the deactivation rates of treating and cracking catalyst. Over-converting of nitrogen can result in accelerated deactivation of the pretreat catalyst. Under-converting of nitrogen can result in higher temperatures for the hydrocracking catalyst along with reduced yield selectivity and product quality.  Depending on specific unit objectives such as product quality or maximum hydrogenation, optimum target nitrogen slip may be anywhere from <1 ppmw to >100 ppmw.

Hydrocracking units without an intermediate sample point between the pretreating and hydrocracking catalysts are becoming more common due to units with both catalyst types in a single reactor or due to reluctance based on safety concerns with taking the inter-reactor sample. Without the intermediate sample, judgment must be made based on the information available, although this is difficult to do accurately. First priority is that both catalysts must be operated within the safe operating limits of the unit including heater duty, quench availability, and bed temperature rise limits. If these conditions are satisfied, then estimating techniques for monitoring catalyst performance can be considered.

The total temperature rise of the pretreating catalyst is the primary indication of hydrotreating severity. With a given feed composition, the relative temperature rise between pretreating and hydrocracking can be evaluated. A decreasing pretreat temperature rise along with increasing hydrocracking temperature rise may imply a shift of the hydrotreating reactions to the hydrocracking catalyst. In this case, a higher pretreating temperature may return balance to the catalyst severities. In the past, a decreased temperature rise in the first hydrocracking bed was interpreted as an increased nitrogen slip, particularly with noble metal catalysts. However, today’s base metal hydrocracking catalysts are more tolerant of increased nitrogen slip and can often have a significant temperature rise from hydrotreating reactions alone. Therefore the hydrocracking temperature rise may not always be reliable for this purpose.  The catalyst supplier can assist by providing operating targets and operating response curves.  

RAHUL SINGH (Haldor Topsoe, Inc)
A hydrocracker unit consists of a hydrocracker pretreat section and a hydrocracking section. This can be housed in one reactor or two or more separate reactors depending on the unit configuration.  Hydrocrackers process a variety of feeds (gas oils, cycle oils, coker gas oils, deasphalted oils, coker naphthas, etc.) with a wide range of properties (S, N, SG, Aromatics, SIMDIST) under variety of process conditions (T, P, LHSV, H2/oil).  Feed is introduced in the pretreat (P/T) section, where predominately saturation reactions (hydrodenitrogenation [HDN], hydrodesulfurization [HDS] and hydrodearomatization [HDA]) occur on the catalyst surface.  The effluent of the pretreater is then processed over the cracking catalyst where cracking and saturations reactions occur to produce the desired products, i.e., naphtha, jet, diesel, with desired yields and specifications.  
A hydrocracking catalyst is comprised of metals (Ni, W, Pt, Pd) supported on zeolites.  Cracking reactions occur on the zeolite while saturation reactions occur on the metal sites.  Nitrogen in the pretreat effluent, i.e., nitrogen slip, has an important role in optimizing the operation of a hydrocracker.  Various ways to determine the optimum nitrogen slip to the cracking reactor are discussed below.

(a)    Cycle Length - Refiners have a target of cycle length of the hydrocracker.  A majority of hydrocrackers are pretreat activity limited. This means that the EOR WABT is reached earlier for the pretreat catalyst compared to the hydrocracking catalyst.  EOR WABTs are known and depend on design specifications or process limitations.  To meet the cycle length, SOR WABT of the P/T can be determined from the expected deactivation rates and EOR WABT.  The SOR WABT of the P/T determines the nitrogen slip based on the reaction conditions.  For example, a higher nitrogen slip will allow the HDC P/T to be operated at a lower temperature and hence, a longer cycle length can be expected.  A high zeolitic catalyst can be placed in the hydrocracker to counter the effect of high nitrogen slip on the activity of the HDC catalyst.

(b)    Volume Swell - Nitrogen slip is determined by the HDN activity of the pretreat catalyst.  For the same wt % conversion in a hydrocracker, a high HDN, HDA, HDS activity of a certain P/T catalyst would generate a P/T effluent with a high API, which will be further processed in a hydrocracker producing a higher volume swell.  If there is hydrogen available on the site, it is always good to use the hydrogen to upgrade the feed to valuable products, and increased swell. The desired volume swell can be translated into API upgrade, which can be interpreted into HDN activity and ultimately a function of WABT for a given feed.  A disadvantage of operating in a higher nitrogen slip mode is putting the burden of this work on the cracking catalyst, which causes it to lose its selectivity to desired products.  This selectivity deteriorates from SOR to EOR when operating in high nitrogen slip mode.

(c)    Hydrogen consumption - Saturation reactions, i.e., HDS, HDN and HDA occur on the pretreat catalyst which require hydrogen. The available hydrogen for chemical consumption to the unit is determined from make-up gas rates and hydrogen purity.  Nitrogen slip depends on the mono, di and tri+ aromatic saturation activity of the P/T catalyst, which combine with HDS and HDN activity to control hydrogen consumption.  Hence, available hydrogen consumption from M/U can be correlated to nitrogen slip from the P/T.

(d)    Conversion - Every hydrocracker unit has a specific conversion target.  Nitrogen slip is a factor in determining the required WABT for the hydrocracker to achieve a specific conversion.  Cycle length requirement, together with conversion, in Hydrocracker can be used to determine the optimum nitrogen slip.  At higher nitrogen slip, the hydrocracking catalyst must run hotter and thus will shift the selectivity from liquid towards gas make.  It is therefore necessary to use P/T catalysts which produces lower nitrogen slip and still meet cycle length requirement.

(e)    Product quality - Hydrocrackers add hydrogen, to lower quality oils, which produces quality products (Naphtha, Jet and Diesel).  Most of the product specifications (Sulfur, Nitrogen, API, Cetane Index, Smoke point) are a function of saturation which is strongly affected by nitrogen slip from the P/T.  A low nitrogen slip from the P/T would produce inter-stage effluent with high API upgrade. This automatically translates into a better product quality, after it gets further upgraded on a hydrocracking catalyst.  Therefore, the product property objectives would set the HDN activity of the P/T.  Nitrogen in the P/T effluent has a strong affinity to adsorb on the HDC catalyst and hinder the cracking and saturation activity, which cannot be compensated by increasing the temperature as the product selectivity would change towards higher gas makes.

(f)    Process constraints - Last, the M/U availability, quench limitations and treat gas play a role in determining the nitrogen slip from the P/T.  Hydrogen availability has to be above a minimum number for a good deactivation rate of the P/T catalyst.  We know from earlier discussions that nitrogen slip is related to hydrogen consumption in the P/T.  Therefore, hydrogen consumption and availability can be used to determine the required nitrogen slip.  Saturation reactions occur in the pre-treater, which is exothermic.  A low nitrogen slip would result from a high degree of saturation activity which would result in large exotherms.  Operations can dictate the accepted temperature profile, i.e., equal bed outlet or ascending temperature profile, or what is the difference in temperature between HDC P/T last bed outlet temperature and HDC inlet temperature.  The available quenches and reactor heater limitations would be a factor in determining the target N slip.  

(g)    Operating conditions - The hydroprocessing reactions are all a function of hydrogen-to-oil, temperature, pressure and feed properties (S, N, SG, SIMDIST).  These factors, together with all of the above, are used to determine the nitrogen slip from the P/T.  The goal is to have a high HDN activity to produce low nitrogen slip which leads to a scenario of superior profits to the refinery by producing high volume swell, excellent product properties, heavy feed processing ability, extended cycle length, operational stability and improved selectivity of the products from SOR to EOR.  Haldor Topsoe offers industry leading superior HyBRIM™ and Hyswell™ hydrocracker P/T catalyst providing refiners with the choice of optimizing their hydrocracker performance, ease of operation and excellent profits over the cycle.  
 

SUNHEIL ABDO and MICHAEL PEDERSEN (Honeywell UOP)
Tolerance for water is dependent on catalyst type and state.  Prior to and during sulfiding and start of run cycle conditioning, hydrotreating catalysts can be quite vulnerable to water.  Moisture will promote mobility of catalyst metals resulting in poorer metals dispersion and lower catalyst effectiveness.  This can be particularly significant for many of the Type II formulations that feature loose association of metals with the catalyst support.  In general, hydrotreating catalysts are very tolerant of even high levels of moisture after proper sulfiding and catalyst conditioning.
Hydrocracking catalysts are bifunctional, featuring both metal hydrogenation and solid acid cracking characteristics.  Each function is susceptible to high levels of moisture.  As with hydrotreating catalysts, metals mobility and agglomeration facilitated by higher levels of moisture are a concern.  For noble metal hydrocracking catalysts, moisture has been limited to as low as 1.5 psi (0.1 kg/cm2) partial pressure.  For base metals hydrocracking catalysts with moderate to high zeolite content, startup and operation at up to 20 psi (1.4 kg/cm2) water partial pressure has been successful.  There are a variety of catalyst formulations so it is advisable to follow the procedures and limits specified by the catalyst provider.  Incidentally, moisture, even at low concentration, is a key contributor to the initial steep increase in temperature requirement for zeolitic catalysts at start of run.  
Particularly during unit startup and catalyst activation there are several potential sources of water.  These may include moisture adsorbed on catalyst from the environment, water precursors in catalyst formulation, water generated from catalyst activation and vaporized process wash water as well as water associated with other process chemicals that may be introduced.  So water is present in hydroprocessing catalyst systems.  The key to success is to control the rate of release of water to a level acceptable for the system.  
For conventional hydroprocessing operations it is generally recommended to avoid free (liquid) water in feed.  But biofuels hydrotreating and renewables processing catalysts must remain robust in high moisture environments out of necessity since water is a prominent reaction product.  And some resid hydrotreating catalysts function better upon introduction of water with the feed.
 

BRANDON MILLER (Shell Catalysts & Technologies)
Operating temperature and bed temperature rise during the first month of operation can be critical with regards to safety and stable unit performance. In general, the limits of temperature and temperature rise are meant to keep the unit within the mechanical & metallurgical design of the equipment, to maximize the catalyst performance, to minimize the catalyst deactivation for the entirety of the cycle, and to meet the desired product quality. In addition, several factors influence the current and potential temperature and temperature rise of any given unit including feed rate/type, treat gas availability, and catalyst type.

One of the main risks early in the cycle with high temperature or high temperature rise is localized hydrogen deficiency leading to increased deactivation and reduced performance. Freshly-sulfided catalyst has not yet accumulated a significant amount of carbon deposits on its surface, and is therefore in a highly active state, sometimes called hyperactive or ultra-active. With the catalyst in this state, coke precursors in the feed speedily react to produce a molecule with an extremely reactive free radical site. Ideally, this site would react with hydrogen; but because of the accelerated reaction rate at the surface of the catalyst in this highly active state, there can be localized hydrogen deficiency. Without hydrogen readily available to react with the free radical site, the molecule will polymerize or condense with another molecule or it may just deposit on the catalyst surface as coke. When coke deposits in this hurried, sloppy way, it often blocks the entrance of the catalyst pores or active sites, resulting in premature permanent deactivation. 

To help reduce this risk, it is generally advised that cracked stocks be avoided during the first 3 days of operation on fresh catalyst, and then slowly introduced following the 3-day period. Cracked and heavy feedstocks are rich with coke precursors (PNAs, asphaltenes, etc.), so controlling these feeds early on minimizes the quantity of coke precursors and allows the initial layer of coke to deposit on the catalyst surface in a more controlled manner. Once the initial layer of coke has laid down, it helps stabilize the active sites and prevent agglomeration of metals. This maximizes long term catalyst activity.

Treat gas balance is also critical in ensuring there are no pockets of hydrogen deficiency in the reactor, thereby minimizing coke formation in reactor beds. Since temperature rise is a result of hydrogen consumption reactions, the bottom of high heat release beds often experiences increased coking tendency due to the reduced hydrogen partial pressure. This can be especially true in top beds where the feed is in its least saturated state. During the initial month of operation many units are especially vulnerable because the heat release profile has just changed from near-spent catalyst operating at end of run conditions, to fresh catalyst at start of run conditions. For instance, the top beds, which had been poisoned and deactivated over the course of the previous cycle, may now generate significantly more heat release. So, the treat/quench gas strategy will need to be adjusted, sometimes drastically, to ensure desirable hydrogen to oil ratios throughout the bed.

Another thing to keep in mind is that coming out of a prolonged downtime the feed mix may be different than usual. Sometimes during an outage, inventories of high value feeds buildup in tankage and upon starting the unit there is an urgency to run-off the excess feeds (cracked stocks, heavy cuts, etc.) at higher-than-normal rates/ratios. In some cases, the combined effect of running more cracked stocks during the most active month of the catalyst’s life can be limiting. For example, a unit that is not usually limited by hydrogen consumption may need to reduce operating temperature during the initial weeks of operation to keep the hydrogen consumption within the availability limits. 
Often when we talk about temperature, or temperature rise, we refer to an average number. However, it is important to realize that many of the limitations faced will be determined by the peak temperature of a reactor/bed, not the average. Most hydroprocessing beds have some amount of radial flow maldistribution. The amount of maldistribution varies wildly and can be caused by imperfections in the reactor internals or installation thereof, catalyst loading, reactor configuration, etc. This maldistribution leads to differences in the radial temperature profile that must be considered when setting unit limits. For example, peak temperature can be critical when considering metallurgical, product aromatic, or mercaptan recombination limitations. So, depending on the amount of maldistribution experienced/expected, the strictness of the limits may need to be adjusted.
Some of the limits we encounter are dynamic and can depend on the current operating conditions of the unit. Here are two examples in addition to a feed change discussed previously:

•    Due to maldistribution, more temperature rise in a bed leads to larger variation in peak temperatures, leading to a changing limit. This may lead to slightly more generous temperature rise limits at start of run compared to the previous limit at end of run.
•    Thermal cracking increases with the absolute temperature. Early in the cycle, temperatures may be low and there is minimal thermal cracking taking place, but at the higher end-of-run temperatures there will be much more thermal cracking. So, the maximum acceptable temperature rise limit may vary with the current absolute temperature. For example, in some units, a high temperature rise could be acceptable at the low start of run temperatures and unacceptable at the higher end of run temperatures when thermal cracking is more prevalent.

Operating temperature and bed temperature rise limits are often set by the maximum allowable temperature of the equipment, under the unit’s operating conditions. These limits are impacted by the feed coming into the unit, the unit operating conditions, the catalyst type, and the operator’s ability to keep the unit within the limits under abnormal or excursion events (e.g. quench availability, quench valve operation, etc.). The overall temperature or temperature rise in any given bed may be limited, especially at start of run, by a combination of these factors; the intent being to safeguard the unit against a temperature excursion, or runaway that cannot be easily controlled. These safeguarding limits and guidelines should be reviewed on a case-by-case basis with the appropriate safety experts, unit designer, and catalyst vendor.

Once all the safety & design limits are accounted for, temperature and temperature profile can be set based on meeting the target outcomes of the unit; whether that be equal bed outlet temperatures at maximum aromatic saturation, or an ascending profile tailored to give fixed product quality. During the first month of operation we typically advise units to slowly and incrementally move towards their ideal operating conditions to avoid inadvertently causing a localized disruption that could lead to premature deactivation, and to avoid overreacting to the changing conditions without giving the unit time to equilibrate and respond.
 

ROBERT STEINBERG (Motiva Enterprises)
There are many things that can contribute to fouling of feed/effluent exchangers. Fouling can occur on either the feed or product side of the exchangers.
 

Possibilities sources of fouling on the feed side include:

•    Dissolved O2. Oxygen can get into feeds if they come from a tank that is not N2 blanketed, this is especially likely if feeds have been imported from another site via a barge. Oxygen can also be present if a feed come from a vacuum tower with an air leak. Corrosion inhibitors or oxygen scavengers injected into the feed as far upstream as possible may help. The best method to remove oxygen is to add an O2 Stripper on the stream that contains oxyg

en.

•    Caustic. Small amounts of caustic that was not water washed can lead to severe fouling.

•    Particulates, scale, corrosion particles. FeS scale is often found in refinery streams. If the source is known, corrosion inhibitors may be able to reduce the amount of scale. Good feed filters may be able to remove some of the scale but FeS particulates can be small enough to pass through most feed filters.

•    Dirty feed. Cracked feeds, especially coker gasoils, tend to be dirty and have small coke particles. Good filtering is essential. If not done at the upstream unit it needs to be done on the hydroprocessing unit. It is often a good practice to filter both places in case one of the filters is bypassed.

•    Salt in Feeds. If crude oil is not properly desalted there can be salts left in heavy feeds. Salts from other sources can also be present at times. A water wash or a desalter can remove salts.

•    High temperatures. High skin temperatures tend to increase fouling. High temperatures may be unavoidable when exchanged against reactor effluent, especially in the hotter shells. An exchanger design that increases velocity and promotes turbulence on the feed side will increase the heat transfer coefficient and reduce skin temperatures. Injecting hydrogen upstream of the exchanger will help.
•    Low velocities. Lower velocities in the exchanger reduce pressure drop but lead to higher skin temperatures, make it easier for particulates to stick to tube surfaces and increase fouling. Injecting hydrogen upstream of the exchanger will help. Recycling hydrotreated product when the unit is turned down will maintain higher velocities in the exchangers.
•    Cracked feeds. Cracked feeds have olefins and sometimes di-olefins which can polymerize and  are more prone to fouling. Cracked feeds can be a particularly severe problem if dissolved oxygen is present. A selective hydrogenation unit or reactor can be used to saturate di-olefins at a relatively low temperature upstream of the main reactor before the feed gets hot enough for severe fouling to occur.
•    Asphaltene precipitation. This is normally only an issue with resid units. Mixing different feeds, especially a lighter more paraffinic feed with resid, can create incompatible mixtures and cause asphaltene precipitation.
 

Reactor effluent is normally cleaner than reactor feed. Olefins get saturated and dissolved oxygen gets converted to water in the reactor. The reactor effluent will always have hydrogen which tends to keep velocities high. However, there are some possible sources of fouling on the reactor effluent side:
•    Salt precipitation. H2S, NH3 and HCl are normally present. These form ammonium chloride (NH4Cl) and ammonium bisulfide (NH4HS) salts when temperatures fall below the salt formation point. The salt point is dependent on the operating pressure and concentration of H2S, NH3 and HCl. Salt point curves can be found in API Recommended Practice RP-932B Design, Materials, Fabrication, Operation, and Inspection Guidelines for Corrosion Control in Hydroprocessing Reactor Effluent Air Cooler (REAC) Systems. Typical precipitation temperatures are in the 300-400°F range for NH4Cl and around 100°F for NH4HS. In addition to fouling, these salts can be extremely corrosive if water is present. Dry salts are not corrosive but an intermittent water wash may be needed to remove them once fouling occurs.
•    Polynuclear aromatics. This is normally only an issue with hydrocrackers, especially the 2nd stage of a two-stage recycle unit. If conversion is too high the PNA concentration can get high enough that they become insoluble in the oil. The lighter cracked products can cause PNA’s to precipitate in exchangers as the effluent cools and more of the naphtha range material condenses.
 

JOE RYDBERG (CITGO)
In our recent experience, fouling on the “feed side” of the feed/effluent exchangers in Naphtha units  is due to corrosion products (Fe) entering with the feed, and processing recycled Naphtha’s particularly from LPG Caustic Disulfide separators.   The recycled naphtha’s can have higher levels of Sodium and Salts (likely amine degradation products that build up in the caustic). 

Other causes can be contamination of cracked stocks into the virgin stocks system.  Exposure to oxygen will cause gum formation.  Crude supply sources have unknown diluents.  Refineries are now collecting more material from various refinery sources and rerunning as slops, for example introduction of flare gas recovery liquids, reprocessed as slop oil; re-processing/chemical cleaning liquids pumped to slop system.  
Use of chemical additives (organic dispersant, antipolymerant, oxygen scavengers) can be used and are used within CITGO to mitigate fouling.  Proper tracking of heat exchanger fouling is important and can aid in scheduling cleanings (requiring unit shutdowns) outside of turnarounds, during catalyst change-outs, etc.  When dealing with especially challenging feeds and / or extending cycle length goals, installation of spare feed/effluent heat exchangers could be value added approach   
Effluent side fouling typically is caused by inadequate water wash, presence of NH4Cl in addition to FeS. HCl can also react to create additional FeS in the presence of H2S.  

ERIC LIN (Norton Engineering Consultants, Inc.)
In a hydrocracker with liquid recycle (could be single stage recycle or two-stage recycle), there exists the possibility of HPNA (Heavy Polynuclear Aromatics) buildup at the bottom of the fractionator.  Although the overall conversion will decrease, the best solution is to have a dedicated bleed stream out of the unit (FCC is a typical destination) to prevent this buildup.  High asphaltenes in the feed are usually a sure sign of HPNA production.

In a residue hydrocracker, the existence of sediment can cause similar fouling in these exchangers.  Sediment can typically be mitigated with the use of slurry oil as a cutter (much easier to acquire for units that also have an FCC nearby).  

SAM LORDO (Consultant)
Fouling in the circuit ahead of the furnace and furnace can be caused by inorganic solids, or polymerization of feed components (organic fouling).  Mitigating fouling from inorganic solids, such as iron sulfide and other corrosion byproducts, sand and silt (from imported feedstocks) is primarily done using feed filters. The pore size is best at 1-5 micron.  The filter can be cartridge style, sand filters.  Some filter arrangement would have backwash capability.
Fouling downstream of the reactor may include ammonium chloride (NH4Cl). Typically, a well-designed Waterwash is used.  The use of salt dispersants are also applicable where Waterwash is feasible
Organic fouling could be from:
•    Stream that contain olefinic/diolefinic components which when exposed to elevated temperatures ass found in the hydroprocessing units
•    O2 contamination of feed or feed component streams
Mitigation of this source of fouling can be done using an appropriate chemical additive, such as dispersant and/or antiploymerant.
 

MICHAEL PEDERSEN (Honeywell UOP)
Most hydroprocessing catalysts require a conditioning period at start of run to allow the active sites to stabilize.  One aspect of this process is the common industry practice to avoid processing cracked feedstocks during the first few days of operation.  Prior to conditioning, fresh catalysts have a high tendency to generate excessive coke when operated with reactive feedstocks or at normal unit operating severity.  A short period of mild operating conditions can pay big dividends in overall catalyst cycle performance while high severity operation at start of run can have substantial negative impact on apparent catalyst activity and cycle length.  In general, catalysts that are claimed not to require conditioning have been artificially inhibited prior to delivery.  
Hydroprocessing catalysts encompass a wide variety of formulations, so a general set of conditioning guidelines is not applicable.  For a specific catalyst system, instructions from the supplier should be followed.
 

SIMERJEET SINGH and RAJESH SIVADASAN (Honeywell UOP)
Fouling of feed/effluent exchangers in hydrotreating units is a common problem leading to throughput losses, increased energy consumption, unit downtime and maintenance expenses for exchanger cleaning. Fouling happens due to changes in feedstock quality, exchanger temperature, fluid velocity, degree of vaporization and exchanger configuration leading to formation of hard carbon deposits (coking), deposition of undesirable polymers (polymerization) and corrosion products. 
For Coker Naphtha Hydrotreater:
• Feed quality issues:
Coker naphtha (CN), by the nature of thermal cracking reactions, contains free radicals, which react with diolefins and olefins to form oligomers and polymers. By itself CN presents a fouling problem in a NHT, however when combined with stored SRN there exists the potential for significant fouling. Storage of CN prior to processing can have disastrous results, as the combination of diolefins, free radicals, and oxygen (peroxides) can lead to rapid fouling on the feed side of the combined feed exchanger (CFE), the NHT charge heater, and the NHT reactor. This fouling can be serious enough to cause premature pressure drop increase along with loss of heat transfer due to fouling in a matter of days if not hours. The downtime associated with addressing this fouling costs the refiner time and money.
The highly reactive diolefins in CN are the four carbon and five carbon species, at the front end of the boiling range. Longer chain diolefins tend to be reactive, but less reactive than the short chain diolefins. Simply increasing the initial boiling point of CN (reducing the quantity of light diolefins) may reduce the tendency of CN to cause fouling. When cracked stocks with significant diolefin concentrations are present, it is UOP’s practice to include a diolefin saturation reactor as a first, low temperature reaction stage in a two-stage reactor system. In this reactor, most of the diolefins are saturated. This reactor is located in between CFE shells and its position is selected such that the inlet temperature is in the range of 320-370°F.
•    Design considerations:  
o    Feed tank blanketing  
o    Design of feed tanks (Fixed/ floating roof) 
o    Hydrogen Injection to preheat exchangers 
o    “Over-Sized” exchangers for clean duty 
o    Exchanger velocities  
o    Dry Point location 

For VGO HDT:
•    Feed quality issues:
o    Fouling is also experienced in units that run straight run feed only, so it is not just a phenomenon that requires cracked olefinic feeds.
o    Fouling from asphaltene precipitation.
•    Design considerations:
o    Same design considerations as coker naphtha HDT except the dry point location.
o    Thermal cracking of feed VGO in feed effluent exchanger can be of main issue if separate feed heating is being used as design feature over combined feed heating. 
•    Fouling Mitigation Strategies:
Many methods exist for managing fouling. The costs of these methods vary, as does their effectiveness. In order to choose the most effective method for managing fouling, an understanding of the source of foulant precursors should be established. Analytical methods are available that can be used to characterize a feed for gums, asphaltene or stability in the presence of oxygen. While these methods may or may not provide a complete solution to exactly where the fouling problem comes from, they may help to characterize the different feeds at a given site and help narrow down the probable root cause. 

•    Avoid oxygen contamination of feed.
Direct feeding – Supply feedstock to hydrotreater from upstream unit without using intermediate tankage.
Benefits:
o    Eliminates residence time in intermediate tankage, thus minimizing formation of other free radicals. 
o    By far the cheapest solution and reduces working capital. 
Risks:
o    Lacks flexibility to accommodate swings in feedstock rate and unit outages. 

Tank blanketing – If tanks must be used, they should be blanketed. Nitrogen is the best blanketing gas owing to its reliably low O2 content and ease of venting to atmosphere. Gas blanketed internal floating roof tanks are most effective in minimizing oil contact with O2 and evaporation losses to blanket gas.
Benefits:
o    Commercially just as effective as direct feed and overcomes all the limitations.
o    O2 cannot react if not in system, therefore should reduce foulant generation.
Risks:
o    Cannot impact O2 brought in with import through other feeds
o    Choice of correct seal for floating roof and its periodic checking and maintenance. 

•    Remove Oxygen from Contaminated Feeds.
Oxygen Stripper – Strips out free O2, including import O2 and removes the potential for further formation of peroxides. Common scheme is for ambient temperature hydrogen stripping of the feed to fuel gas system. 
Benefits:
o    Only feed streams exposed to O2 need to be stripped.
o    Maximizing direct feed to the unit in combination with stripping the small O2 contaminant stream is generally more economical than stripping the complete feed stream.
Risks:
o    Residence time, particularly in imports, may result in some polymer reaction occurring. 
o    Expensive option in terms of equipment, and is not so effective if the peroxides/ polymer has already been formed upstream of the stripper.

Injection of anti-oxidant chemical – Antioxidant chemicals have been used with a degree of success in some locations.
Benefits:
o    Act as chain stoppers that react preferentially with O2 and peroxides, making them unavailable to take part in free radical polymerization reactions.
Risks:
o    Although chemical treatment can help, it is not always successful and it tends to be most effective when the antioxidant is dosed into the upstream unit rundown ahead of the storage tank.

•    Remove foulant/prevent laydown.
Hydrogen treat gas injection – Inject hydrogen treat gas upstream instead of downstream of preheat exchangers.
Benefits:
o    Hydrogen gas increases turbulence and can also help to reduce polymer formation reactions.
o    For VGO HDT hydrogen injection especially for units with separate heating of VGO will prevent thermal cracking of VGO.
o    Avoid dry point in exchanger areas where the feedstock is completely evaporated towards dryness as severe fouling may happen. Polymer and gum tends to build up on the shell-side behind baffles, because of relatively stagnant zone. Evaporation of feed leaves less liquid solvent for the gums and gums get deposited.  Most severe at the liquid dry point.

•    Modify exchanger design – Modify exchanger internals, maintain high velocities in exchangers, appropriately oversize exchangers to lower high tube wall temperatures below the critical temperature required for coking or polymerization.
Parallel exchanger – Flexibility for bypassing and cleaning.
Benefits:
o    Clean all exchangers on-the-run, extra exchangers mean no loss of throughput to clean.
Risks:
o    No reduction in rate of fouling.
o    Additional design features required (such as PRV’s) to safely by-pass/isolate exchangers.

•    Anti-foulant chemical injection. 
Benefits:
o    A reduction in the rate of fouling.
Risks:
o    Fouling mechanisms will still occur, probably downstream.

•    Prevent corrosion
Corrosion resistant tube metallurgy – select appropriate tube metallurgy to prevent formation of corrosion products that aid the process of foulant formation such as naphthenic acids or high TAN feeds.
Benefit:
o    Easy to implement for new unit and revamp of existing unit.
Risks:
o    May not be the best solution as metallurgy upgrade is expensive and components other than tubes can still provide corrosion products to aid fouling.
 

IHSAN RAAD (Shell Catalysts & Technologies)
 There are several types of fouling in Hydrotreating feed/effluent exchanger units, the three most common types in the industry are:
1.    Inorganic particulates.
2.    Organic deposits.
3.    Ammonium salts.
Each type of fouling has its own characteristics and deposition locations. Knowledge of the type of fouling and the underlying deposition mechanism is essential to tackle the fouling problem. This can either be done by eliminating the root-cause, or by selecting a fit-for-purpose and cost-effective abatement approach.
 

1.    Inorganic particulates: Inorganic fouling is mainly caused as a result of iron sulfide, sodium or coke fines that can either be carried from upstream units or generated in-situ in the preheat exchanger network. These foulants are:
•    Iron Sulphide (FeS) and Iron Oxide (FeO, Fe2O3): Scales of iron oxide (FeO, Fe2O3) and iron sulphide (FeS) are generated as corrosion products within the unit itself but can also come from upstream units, intermediate storage and transport from well to refinery. Important corrosion sources are furnace tubes (hot sulphur corrosion), the CDU overhead condenser and the reactor effluent air cooler. Iron corrosion products in VGO’s are also associated with processing of naphthenic crudes.
•    Sodium (Na): Na can come from brackish or salty cooling water (i.e. leaking heat exchangers) or from processing water-containing slops or imported feeds. Sodium in combination with iron has been known to promote coke formation under conditions of high temperature and low pp H2. 
•    Coke (C): Coke fines can be entrained from VBU’s and cokers, which cause mainly fouling of the feed side of the feed/effluent heat exchangers, furnace tubes and the top beds of reactors. VGO hydrotreaters might also experience coke formation due to a poor separation in the upstream HVU. 

2.    Organic deposits: The organic foulants are primarily gums formed as a result of processing cracked material and accelerated if the material is exposed to oxygen at any time. The types of foulants are:
•    High di-olefin Content: Di-olefins (molecules containing (multiple) double bonds) are mainly found in product streams from (thermal) cracker units but can be present in other streams as well. At the right temperature level (~400-500°F) they polymerize readily to form gum-like substances often showing up as greyish flakes on the FEHXers. The rate of this reaction increases at higher temperatures, making the FEHX especially vulnerable. Elimination of feed streams with high di-olefin content is an easy way to reduce fouling but may not be preferred economically. If long periods of operation with high di-olefins content are expected, an option to avoid high fouling rates could be the installation of a low temperature di-olefin saturation reactor.
•    Oxygen in Feed: Oxygen can form a range of different molecules when it is dissolved in a hydrocarbon stream (e.g. peroxides, carboxylic acids, aldehydes and other oxygenated compounds). Amongst other problems, these molecules can initiate polymerization reactions to form gums. Oxygen can enter a feed stream in several ways, including but not limited to air-breathing storage tanks, marine or surface transport vessels, leaks in equipment operating in sub-atmospheric pressure or faulty pump seals. One common point for oxygen to enter feedstock is during storage. Feed streams can be routed directly from unit to unit to avoid intermediate tank storage. Another option is to put in a bypass jump-over on the tank such that only the extra feed goes to tankage and the rest will bypass and go directly to the hydrotreater. All tanks used for storing hydrotreater feed should be nitrogen blanketed, also for straight-run feed. If this is not done, gum formation and other side reaction might happen in the tank itself. Other mitigation is to run it through a stripper, fractionator or distillation unit before introducing it to the unit in order to strip away both dissolved oxygen and the oxygenated compounds, like peroxides. A last option that is used to avoid oxygen related fouling is by injecting anti-oxidants into the feedstock. These compounds are only effective when they are injected prior to the feed stream coming into first contact with the oxygen. So, it should be ideally be injected at the source prior to transportation to site or send to storage. Also, good mixing of the anti-oxidant with the hydrocarbon streams is essential. Only then can they prevent gum formation during storage. 

3.    Ammonium salts: There are several types of salts that can formed in the effluent exchanger.
•     If ammonia and HCl are present, ammonium chloride may deposit directly from the gas phase. The sublimation point in the process depends on the operating conditions (pressure and temperature profile) but also on the concentrations of ammonia (generated from hydrogenation of nitrogen compounds) and chlorides. Main locations of deposition are the feed/effluent heat exchangers (effluent side), the air cooler and recycle gas compressor valves. In the dry state this salt is not corrosive, but in areas where the water dew point is approached, the deposited NH4Cl salt will become moist and can be very corrosive (NH4Cl salts are hygroscopic, therefore the stream temperature must be maintained 15-20 C above the water dew point to assure dryness). Furthermore, apart from the danger of excessive corrosion, NH4Cl deposition can drastically increase the pressure drop and to decrease the effective duty of feed-effluent heat exchangers.
•    Like Chloride, organic bromide will convert to hydrogen bromide after hydrotreating and then react with ammonia to form ammonia bromide salt in the exchanger. 
•    Ammonium bisulphide (NH4HS) hydroprocessing reactors convert sulphur and nitrogen compounds in the feed to H2S and NH3. On cooling, these two compounds react to form NH4HS. In the absence of water, NH4HS deposits to form a crystalline solid that can cause plugging of the reactor effluent air cooler. This will occur at relatively low temperatures, 10-30°C. If the dew point of water is reached in the effluent air cooler (or water cooler) or if insufficient wash water is injected.

In summary, foulants are typically found on the feed side of the preheat exchangers include various gums or polymers, iron sulfide, and salts. The organic fouling due to gums and polymers results from the polymerization of unstable species in the unit feed. Therefore, in order to determine the risk of organic fouling for a particular feed stream, detailed analysis of the feed is required to determine the problematic species in order to evaluate the fouling propensity and mitigation strategies. Another key factor to consider is the oxygen content of the feed stream as this can promote the polymerization of various unstable compounds, particularly olefins. Therefore, it is a good practice to exclude oxygen from feed storage tanks using a nitrogen blanket. However, this method is ineffective with streams already exposed to oxygen. The inorganic fouling is mainly caused as a result of iron sulfide that can either be carried from upstream units or generated in-situ in the preheat exchanger network.  Identifications of the contaminants source and mitigations are key to eliminate the inorganic foulants.

SERGIO ROBLEDO (Haldor Topsoe, Inc.)
To answer this question, we need to differentiate between feed-side fouling and effluent-side fouling. Potential causes and mitigations will depend on which side is experiencing the fouling.
Feed-side fouling in your F/E exchangers can be the result of:
•    Olefins
•    Oxygen
•    Particulates 
Olefins/Oxygen
Olefins are normally introduced with cracked stocks in the feed. Typically, olefin gumming happens at lower temperatures (300 – 350 °F). Gums formed from peroxides, as a results of oxygen contamination of straight run feed, usually occurs at >400 °F. 
In the case of coker naphtha, conjugated diolefins are present which are highly reactive species.  In the presence of very small amounts of oxygen, or at elevated temperatures above 450 °F, these molecules will radically polymerize to form gum that can foul exchangers causing poor heat transfer as well as high pressure drop. If the feed contains significant quantity of coker naphtha then these Diolefins must be removed to prevent gum formation.  
The coker naphtha should preferentially be sent to directly from the coking unit to the hydrotreater to prevent contamination with oxygen. Even straight run stock, which may be part of the feed component, must be prevented from contacting oxygen by storing the feed in a nitrogen blanketed storage tank.  
Even with strict adherence to avoid feed contact with oxygen, the diolefins in the coker naphtha can polymerize at elevated temperatures.  A dedicated saturation reactor operating in the range of 300 °F to 450 F will ensure that these highly reactive species are removed from the feed before polymerization can take place.  Once the diolefins are removed from the feed then the feed can be heated to the required temperature for the required operating scenario.
Keep in mind that even though cracked stocks are not fed directly to a unit, there is potential of introducing cracked stock in sites that process slop in their Crude unit.
Particulates
Particulates, at high enough concentrations, in conjunction with low tube velocities, can result in these particulates settling out and plating on the surface. If no filter is present, then plans should be made to engineer and install a filter system to reduce the amount of particulates present in the feed. There are also companies that offer tube inserts to reduce the likelihood of particulates settling out and plating on the tube surface, preventing a loss of heat transfer.
Even with filters and tube inserts, if any gumming is taking place in the exchangers, then any small amount of particulates present will be picked up by the gums formed.
Examples in industry where fouling occurred on the feed side are:
•    Cracked stocks blended with Canadian crude coming down the pipeline. These formed gums with oxygen and fouled the exchangers.
•    Crudes from Venezuela blended with cracked stocks.
•    Virgin jet was exposed to oxygen and fouled in exchangers operating >400 °F (this happened in multiple units).
o    Similar example with natural gasoline.
As mentioned before, preventing oxygen ingress via direct, hot feed of cracked stocks to the unit, along with floating roof and/or nitrogen blanket on tanks is imperative. Tube inserts are also a viable option to prevent fouling where tube velocities are low enough allowing particulates to drop out in the exchanger. Most importantly, quality control is paramount in preventing this, or reducing further loss in performance. Examples of actions are:
•    Notify the shipper and have the diluent stream changed.
−    What is the crude source? Are there potential cracked stocks coming down the same line? How about Canadian crudes?
•    Notify crude supplier about the poor quality.
•    Install an oxygen stripper.
−    Done for the virgin jet example.
•    Install nitrogen blanket on feed tank or change to floating roof tank.
•    Inject antioxidants into the tanks (mixed results).
•    Bypass tank and go hot (direct) to the unit.
As for effluent side fouling, this is typically the result of NH3Cl (salts), which are the result of high levels of Cl in the feed. A water wash should be installed to remove these salts, and continuous is recommended versus intermittent.
•    A licensor can help calculate where the water should be injected.
–    Need to inject enough water and at the right spot to keep it as a liquid and not vaporize
–    Liquid water will wash out the salts while steam will not
–    Licensor can calculate where the dry point will occur and how much water needs to be injected
•    Many examples of where water wash helps
–    If done at the right spot with the correct amount of water and at the right temperature
•    Boiler fouling
–    Water treatment company can help with this
There is also a very good P&P this year covering reactor effluent diseases jointly presented by Flint Hills Resources and Marathon. Please plan to attend to learn more.
 

AMIT KELKAR (Shell Catalysts & Technologies)
There is limited boiling point shift from feed to product in a typical diesel hydrotreater. The boiling point shift correlates strongly with H2 consumption which is dependent on feed properties and unit conditions. In our experience, boiling point shift varies from 5 – 10 oF for mostly straight run feed to 30 oF plus for highly aromatic feeds such as LCO and LCGO. The ASTM D975 specification for Ultra Low Sulfur Diesel is a maximum D86 T90 of 338 oC (640 oF). In most units the feed cut point is set so that it meets the final product distillation specification. Raising the cutpoint worsens the cold flow properties particularly for paraffinic straight run streams. Cut point for straight run streams is set to ensure product meets the cold flow specification.
In addition to worsening cold flow properties, raising the cut point brings in streams with more refractory sulfur species. Substituted dibenzothiophene is a classic example of this type of molecule wherein the sulfur atom is sterically hindered by the alkyl groups resulting in very low reaction rate. These types of species are more common in cracked feeds such as LCO and LCGO. Increasing the cut point for such feeds is likely going to require higher delta WABT compared to a similar change for straight run feeds.
T90 or T95 is best suited for unit monitoring. FBPs tend to have a lot of variability compared to T90 and T95 and are unsuited for use as a controlled variable. Repeatability and Reproducibility of T90 or T95 is much better compared to FBP for D2887.

SERGIO ROBLEDO (Haldor Topsoe, Inc.)
Both 90% BP and FBP should be monitored in an Ultra Low Sulfur Diesel unit. ASTM D-975 establishes the maximum 90% BP at 640 °F, based on ASTM D-86 distillation method. Therefore, the feed component(s) 90% BP should be controlled, such that based on their volume percentage, the blended feed meets ASTM D-975 specification. A diesel hydrotreater is not ideal to correct 90% BP. An increase in WABT of roughly 50 °F is required to drop the 90% BP by only 5 °F, and this is mainly via thermal cracking.

It should be noted, that the use of one of Haldor Topsoe’s very selective dewaxing catalysts can shift the T90 significantly. This will enable the refiner to increase the endpoint of the feed while still meeting the T90 spec, resulting in more diesel barrels. 

With respect to FBP, ASTM D-86 is a poor method to truly capture how big of a tail and as such, how much coke precursors are sent to the diesel hydrotreater. ASTM D-2887 (SimDist) is a much better method to understand how high the tail is to the hydrotreater. Typical off-set between both methods is usually 75 – 100 °F, but we have seen discrepancies as high as 600 °F in the most extreme cases. Needless to say, the amount of coke precursors, and as a result deactivation rate and cycle length were very different than expected from D-86 method. Therefore, before drawing harder from a specific stream, it is recommended to analyze the stream via D-2887 to have a baseline sample to compare to.

As for what level is acceptable for the tail, it depends somewhat on hydrogen partial pressure. A higher-pressure unit will suffer less coking than a lower pressure unit. The higher amount of poly-aromatics present, the higher the propensity for coking. For each changing aromatic ring class, the effect can be an additional 10-25% increase in deactivation. The actual increase will depend on the hydrogen partial pressure and ring-class type.

MICHAEL PEDERSEN and VERNON MALLETT (Honeywell UOP)
Foremost, feedstock boiling range must be selected to permit satisfying product specifications such as cold flow properties, gravity and distillation limits.  The most useful indicator of acceptable boiling range will depend on operating experience and the constraining specifications at each site.  For example, there may be more flexibility defining feedstock if the controlling specification is ASTM D-86 T90 than if the limit is true boiling end point.  Some boiling range reduction can be expected in a ULSD hydrotreater.  

The more aromatic the feed and the higher the unit operating pressure, the more significant the impact. 
As indicated in the question, boiling range limits are dependent on feed type.  Product cold flow properties often constrain maximum cut point for straight run streams.  Particularly for cracked stocks, sulfur and nitrogen content as well as amount of polyaromatics compounds increase rapidly with boiling point.  As an example, Light Cycle Oil polyaromatics content could increase from approximately 4 weight percent at a Final Boiling Point of 690°F, to 16 weight percent at a Final Boiling Point of 730°F.  With increasing end point the complexity of sulfur compounds also increases, including more dibenzothiophenes. This will result in higher operating temperatures to achieve product ULSD. Hydrogen consumption will increase accordingly.
                                                                                                                                                                                    Processing Coker Diesel material in a diesel hydrotreating unit also brings processing complexity and severity.  A mild hydrocracking operation is a possible option to allow processing higher distillation boiling range feeds. ULSD product quality can be achieved with minimal yield selectivity shift. There are several examples in which a straightforward revamp of a diesel hydrotreating unit enabled mild hydrocracking operation, as long as there is ample unit design pressure and hydrogen supply.  A revamp does require attention to a few critical details.
 

ROBERT STEINBERG (Motiva Enterprises)
Wash water is injected to hydrotreating and hydrocracking reactor effluent upstream of the Reactor Effluent Air Cooler (REAC) to wash out ammonium salts (NH4Cl and NH4HS) that would otherwise deposit. Such salts foul and plug up exchangers. If the salts are wet, they are also extremely corrosive. When injecting water, enough water needs to be used to limit the NH4HS concentration in the water phase downstream of the REAC and to keep at least 25% of the injected water in the liquid phase at the injection point. If all of the injected water were to vaporize, the water would start to condense in the REAC. The first drop of water that condenses has a high concentration of HCl and will be very corrosive. It is highly desirable to have all, or nearly all, of the HCl be dissolved in the water before the reactor effluent gets to the REAC. Any remaining HCl may condense as the effluent is cooled and corrode airfan tubes.
 

To get all of the HCl into the water phase at the injection point there needs to be good contact between the liquid water and the vapor. A good flow regime (i.e. – churn flow in a vertical upflow portion of the line) promotes such contact. But the best way to ensure good contacting is often to use a static mixer. While a static mixer can be used in most situations there can be issues with using one:
•    Static mixers are relatively expensive.
•    Static mixers can be bulky and take up more space than is available.
•    Static mixers increase pressure drop. If there is not sufficient pressure drop, they will not achieve full contacting of the water and vapor phases. The extra pressure drop means the Recycle Compressor needs additional head to maintain the desired recycle gas rate, some extra head may also be required for the Charge Pump, Wash Water Pump and Make-Up Compressor.
•    Static mixers can occasionally get plugged up, particularly if the wash water is dirty.

The following items can be considered when deciding if a static mixer should be used in a particular application:
•    Distribution of vapor and liquid phases at the REAC inlet. With poor distribution there are more likely to be tubes without adequate water where a first drop of water could condense.
•    Severe service such as a high chloride content. If there is little or no chlorides there is less need to ensure good contacting of water and vapor.
•    Metallurgy of the REAC and the piping upstream of the REAC. With corrosion resistant alloys there is less consequence if there is poor contacting of the water and vapor.
•    Flow regime downstream of the water injection point. If there will always be good mixing due to the flow rates and piping orientation there is no need to have a static mixer. However, even if the flow regime ensures good mixing at normal flow rates, turndown and variations in oil, water and gas flow rates should be considered as well.
•    Available pressure drop. If there is not enough pressure drop for good contacting in a static mixer, alternatives need to be considered.
•    Wash water injection equipment. If a full cone spray nozzle is used within its design operating range it can spray water across the full cross-sectional area of the pipe and get good contact between the water and the vapor. Without such an injection there is unlikely to be good contact immediately downstream of the injection point.
•    Amount of injected water remaining in the liquid water phase at the injection point. More water gives a better chance of contacting all the vapor and scrubbing out all the HCl.
•    REAC bundle arrangement. If there are multiple rows per pass in the REAC, the liquid will tend to go preferentially to the lower row and there may not be adequate water in the upper row. Even with a single row per pass the flow should be annular to avoid points where a first drop of water can condense.
 

RICHARD HOEHN (Honeywell UOP)
UOP’s experience has shown that a static mixer upstream of the effluent air cooler is not necessary if the unit is designed according to UOP practice.  On the downside, static mixers can trap debris.

LARS JORGENSEN( Haldor Topsoe)
All wash systems upstream of the REAC will be supplied with a spray system to ensure good contact between water and reactor effluent stream. In addition, a static mixer will be included for systems where the feed chloride content is high.

MAX LAWRENCE (Shell Global Solutions)
A static mixer is recommended upstream of the REAC in hydroprocessing services that require continuous wash water injection.  The static mixer is installed downstream of the wash water injection point to provide thorough mixing of the wash water and the process gas stream.   If the process gas stream includes HCl (or HF), the static mixer ensures that unvaporized wash water efficiently scrubs the halides from the process gas stream.     

SAM LORDO (Consultant)
In this service, using a static mixer may be require if there is an inadequate amount of washwater going to REAC section. Using a static mixer would enhance contact between water and hydrocarbon. This is not a normal operation.