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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.  
 

JOE RYDBERG (CITGO)
At CITGO Lemont Refinery, the base practice is to install post weld heat treated carbon steel piping for the main run of the pipe upon initial fabrication. For a set length downstream of control valves, 304 stainless steel is installed to try and mitigate the corrosive effects of flashing across the valve. However, in recent years, there has been a notable increase in the amount of corrosion that has been observed on the carbon steel lines. In areas that have been found to be corroding rapidly, an installation of 304 stainless steel has been put in place of the carbon steel. The concern of chloride content and its cracking potential in the circulating stream has been weighed against the risk of carbon steel corrosion. Thus far, the very localized corrosion of carbon steel and its difficulty of detection by common NDE methods has been deemed riskier than chloride cracking concerns, and the installation of 304 stainless steel has taken place.

The reason for choosing 304 stainless steel over 316 stainless steel is the damage that has been observed is mostly localized corrosion at heat affected zones (welds).  While 316 is generally better for pitting resistance in aqueous systems, the benefits it has over 304 in a general wall loss scenario are not significant enough to warrant the additional cost. We have found 304 stainless steel to be adequate in resisting corrosion (as we have a long history of its use just downstream of control valves in this very environment) to an acceptable degree. That is the primary driver behind the choice to use 304 stainless steel. 

JIM JENKINS (Shell Global Solutions)
For “normal” amine service, carbon steel (CS) is the material of choice.  However, hydrotreater applications that operate at higher pressure will have higher acid gas pick up in the amine.  This causes higher temperatures in the contactor.  For applications where temperatures exceed about 90⁰C (194⁰F), austenitic stainless steel (SS) is normally selected for the bottom 1/3 or ½ of the tower.  
When chlorides are present, SS can experience chloride-induced stress corrosion cracking (CSCC) at temperatures above 60⁰C (150⁰F).  Alternate materials not subject to CSCC include alloy 825, alloy 625 and alloy 2205 (duplex).  
The industry generally limits the amount of chlorides in the circulating amine system.  When high pressure hydroprocessing units require SS contactors, the chloride content of the amine is generally limited to 250 ppm (maximum).  
In summary, recommended materials of construction for amine systems that contain chlorides:
 

CHRIS WOZINAK (Honeywell UOP)
Historically, UOP has designed Amine Scrubbers with Killed Carbon Steel that is post weld heat treated (PWHT) and has additional corrosion allowance.  Due to the corrosive nature of the rich amine, the trays are upgraded to 304SS and the mesh blanket is typically 316SS.  While most customers have seen success with this design, even with feed sulfur levels exceeding 3%, there are some who have seen accelerated corrosion on the wall opposite the rich gas inlet line.  Typically this location will receive weld buildup, followed by application of stainless steel patch plates or weld overlay, followed by PWHT.  If the unit contains chlorides, UOP is concerned about the potential for chloride stress corrosion cracking (Cl-SCC) of any austenitic stainless steels (300 series) and so we will upgrade the 304SS and 316SS to Alloy 400 (also known  as Monel), which has good resistance to chlorides and H2S at moderate levels.
 

JOE RYDBERG (CITGO)
The biggest key factor in amine contactor operation in hydroprocessing units is controlling the lean amine flow and the lean and rich amine loading.  We have seen corrosion in rich amine piping likely due to elevated H2S loadings and higher velocities (both lean and rich), especially when unit rates are changed quickly.
 

Adjust amine circulation to maintain desired rich amine loading (Typically 0.3-0.45 mol total acid gas / mol of amine loading).  Ensure sufficient amine by:
o    Sampling rich amine and determining H2S loading for optimization (reduce amine circulation) 
o    Controlling lean amine flow based on temperature rise of amine coming in or out of contactor (relatively new )
o    Set minimum amine circulation based on maximum H2S recycle content
o    When pushing unit rates, amine loading, fluctuations in pressures and flows can “overwhelm” the amine and H2S can breakthrough.
o    Develop calculation tools to estimate sulfur load and set lean flow rates accordingly.
 

Adjust amine regenerator to maintain desired lean amine loading to meet FG H2S specifications.  

There have been upsets in the sulfur unit due to hydrocarbon carryover into the amine.  Typically this is caused by large hydrocarbon carryover events (Loss of HPS levels, upstream Hydrocarbon fractionators). Hydrocarbons cause foaming, solids will contribute to that as well.  The amine system has to be kept clean (filtered).  
•    Amine temperature typically controlled to 120-130F for H2S control (less critical for MDEA based systems)
•    Properly designed wash water trays designed and installed to minimize entrainment / carryover of amine into recycle gas compressor (4 trays, with a water circulation to provide adequate tray loading)
•    Process inlet temp should be maintained 10 degF above amine temp for vapor hydrocarbon systems to prevent condensing
o    Less of a risk of condensing Hydrocarbons in Recycle H2 amine contactors
 

ROBERT STEINBERG (Motiva Enterprises)
Amine absorbers that remove H2S from a gas phase stream need to be operated with the proper amount of amine, at the right temperature and without liquid hydrocarbons.

Amine rates need to be high enough to remove the H2S in the sour gas and to keep the H2S concentration in the rich amine low enough to avoid corrosion. In low pressure service, such as a stripper offgas going to fuel gas, only a limited amount of H2S can be captured by the amine and the amine rate needs to be adjusted to control the H2S content of the sweet gas.

In a higher pressure contactor, such as in the recycle gas loop of most hydrotreaters, high concentrations of H2S can be captured by the amine. Allowing the rich amine to be saturated with H2S in a Recycle Gas Scrubber would lead to excess corrosion in carbon steel lines and equipment. Normally, the amine rate has to be adjusted to maintain the H2S loading in the rich amine within acceptable limits. Typically, at these amine rates there is minimal H2S remaining in the sweet gas and changes in the amine rate have negligible impact on the remaining H2S in the sweet gas. The maximum H2S loading in rich amine is normally expressed in units of moles H2S per mole of amine, the maximum acceptable limit depends on the type of amine used.

Liquid hydrocarbons should be avoided in gas-phase amine contactors as they tend to cause foaming. Liquid hydrocarbons are prevented by:
 

•    Avoid carryover of liquid hydrocarbons in the sour gas to the scrubber. This requires proper sizing of the upstream separation vessel like a Cold Separator with good internals. As a minimum there should be some sort of inlet device to help separate gas and liquid in the Cold Separator and a mesh pad below the vapor outlet nozzle. In some cases, particularly if the unit charge rate has been increased, more sophisticated devices like a cyclone separator in the Cold Separator are needed. If the Cold Separator is not large enough, an additional knockout vessel on the vapor line between the Cold Separator and the scrubber may be helpful, such a vessel normally has internal cyclones to remove liquid carried over from the Cold Separator.
•    Keep the lean amine to the scrubber hotter than the sour gas to the scrubber to prevent condensation of hydrocarbons. A minimum 10°F margin is typically used but in some cases such as very high pressure units, margins of 20-30°F may be needed. It is often a good practice to provide a lean amine heater to heat the amine up to required temperatures as there may be limited ability to provide additional cooling of the sour gas in the upstream Reactor Effluent Air Condenser (REAC). A lean amine heater would generally be provided upstream of the Lean Amine Pump so it can be a low pressure exchanger. Low pressure steam is normally an adequate heating mechanism. Avoid using higher temperature heating sources that would result in high skin temperatures and potentially degrade the amine – limiting maximum skin temperatures to less than 260°F or simply keeping temperatures lower than in the Amine Regenerator Reboiler is a good practice.
•    Provide facilities to skim hydrocarbons from the sump at the bottom of the scrubber. Even with good upstream separation and high lean amine temperatures, some hydrocarbons will often accumulate. These hydrocarbons are generally insoluble in aqueous solutions and form a separate liquid phase, since the hydrocarbons have a lower density than amine solutions they tend to accumulate on top of the rich amine. Level control devices using a dP transmitter will give a false low level if a significant hydrocarbon layer is present. Measuring the hydrocarbon layer can be difficult as the amount of hydrocarbons in an external level gauge will not be the same as in the tower. Using a gauge with a small level range by first lowering the level until it is drained and then raising the level to the middle of the gauge will give the same amount of hydrocarbon in the gauge as in the vessel, this information can be used to decide how much material to skim. Alternatively, an overflow device can be provided in the vessel sump at around the 50-60% level to continuously skim whatever liquid is at this elevation.

ALFREDO VILLA (Haldor Topsoe, Inc.)
Several operating conditions should be maintained and evaluated to determine the condition of the amine system. A properly operating amine contactor will yield a clean recycle hydrogen stream and minimize impact to downstream amine regeneration equipment. 
Operating with a clean amine should mitigate potential foaming, corrosion, and heat stable salt build-up in the absorber system. Monitoring the concentration and circulation rate of the amine used is crucial to maintaining good operation in the absorber. These parameters will allow for adequate hydrogen sulfide absorption and maintain target recycle gas concentration. An adequate concentration of amine will reduce the potential of corrosion in the system by maintaining low rich amine loadings. A target amine strength will depend on the type of amine circulated and should be discussed with the amine vendor. Periodic visual inspection of the amine should be conducted to note changes in color and solids content. Typically, lean amine has a pale-yellow color and should be relatively free of solids, a change in color could indicate an increase in corrosion rates. For example, a change from pale-yellow to dark green could indicate a change in the iron sulfide content of the amine and an increased solids content. Solids in the system could lead to erosion of the iron sulfide scale protecting the inside of the piping. Removal of this scale exposes unprotected metal, which could lead to further corrosion. Along with the removal of the iron sulfide scale, solids in the system could stabilize foam. 
Temperature control is of importance in any amine contactor system as it can lead to poor acid gas absorption. The loading capacity of the amine is directly impacted by the lean amine temperature, the cooler the lean amine the greater its H2S removal capacity. However, it is recommended that the lean amine temperature is maintained warmer than the gas feed. Lean amine temperatures which are too low, increase the potential for condensed hydrocarbons, which could lead to foaming in the contactor. Feed gas temperature control is also recommended as a high feed gas temperature leads to a higher lean amine temperature, which reduces the capacity of H2S removal.

GARY BOWERBANK (Shell Global Solutions)
Solvent hygiene is high on the list.  Poor solvent quality, which is often measured in terms of high degradation products (Heat Stable Salts) or high suspended solids, can lead to both corrosion and fouling of the system but also increase the risk of foaming events.  These foaming events are the most common issue to impact operations of any amine contactor, which can result in losses of solvent (if carried over beyond KO vessel) or off specification product.  Sites may often focus on treating the symptom by dosing an anti-foaming agent, however we prefer to understand the root cause and tackle that.  If not linked to solvent quality; then entrainment or condensation of hydrocarbons in the contactors is the most common root cause.  So, having high efficiency gas/liquid separators upstream and maintaining lean solvent at least 5oC above the gas temperature are critical.

ERIC LIN (Norton Engineering Consultants, Inc.)
Other than the type of amine used, the most important operational factors for amine gas contactors are temperature and pressure drop.  Lean amine coming to the column should be at least 10°F (6°C) hotter than the sour gas to prevent condensation of hydrocarbons and cause possible foaming.  Sometimes a LP Steam heater may be necessary if the lean amine is too cold from the ARU.  On the other hand, if the lean amine is too hot (~25°F greater than the sour gas), there exists the possibility of appreciable amounts of amine carryover into the sweet gas.  Contactors should also have alarms indicating high pressure drop normally caused by foaming.  Most contactors should have the capability of periodically skimming the oil from the top of the rich amine in order to proactively prevent this from happening.

PRASAD HARDIKAR (Honeywell UOP)
The primary purpose of an amine contactor /amine scrubber is to remove H2S from the circulating recycle gas stream or remove H2S from hydrotreating unit off gases before blending hydrogen rich gas with fuel gas.

Here are key factors to be considered in various stages:
Design/ Commissioning Stage:
•    Remove oil or rust inhibitors present from construction
UOP always recommend degreasing of new/modified amine column with 2 wt% soda ash neutralization solution before putting it in service. Degreasing removes grease or protective oil layer on amine contactor /scrubber trays and internals which can contribute to foaming. Hydrocarbon layer is meant for protecting trays/internals from rusting during transportation and storage.
Operation:
•    Limit hydrocarbons entering with recycle gas:
•    Keep lean amine 3-5°C (5-9°F) warmer than inlet gas to prevent hydrocarbon condensation and consequent foaming
•    Ensure mesh blanket in upstream knockout drum is working 
•    Lean and Rich amine loading:
Amine loading is expressed as mole loading of H2S per mole of Amine. Lean amine loading indicates efficiency of amine regeneration unit and capacity of lean amine to absorb H2S. Rich amine loading sets maximum H2S capacity for rich amine. 
Higher rich amine loading increases potential for erosion in the rich amine lines due to two phase flow in the rich amine lines especially from downstream of rich amine control valve.  This can have negative effect on the operation of the amine regeneration system as the lean amine will eventually become more contaminated with iron particles and have deficient performance. It is recommended to maintain rich amine loading < 0.4 mol acid gas/mol MDEA with KCS metallurgy lines.
It is quite tempting to reduce lean amine flow during operation from an optimization perspective. However, such reduction in flow with same acid gas (H2S) content increases rich amine concentration well beyond acceptable rich amine loading as amine will continue to absorb acid gas till saturation. Hence monitoring rich amine loading is a critical factor in amine contactor operation.  
•    Amine Appearance and Quality 
Amine appearance and quality is one of the critical aspects to monitor. 
It is observed that increase in total dissolved solids (TDS), total suspended solids (TSS), heat stable salts (HSS) enhances foaming tendency in lean amine. TSS are expected to be NIL and HSS at < 0.5% for lean amine solutions. These parameters are better controlled with mechanical filtration (with activated charcoal) and it may require replacement if foaming issue persists.  
Lean amine displaying a green color is indicative of the presence of solubilized iron sulfides in the amine due to corrosion or issues with amine system filtering.  Lean amine with a slight green color will continue to remove H2S from recycle gas adequately, but if soluble iron sulfide buildup in the amine is not eliminated, foaming could be a potential concern. 
•    Amine Color Guideline:
– Light straw = amine is in good condition
– Black or dark green = corrosion is taking place
– Light green rich amine = small particulates present
•    Avoid excessive foaming inhibitor as this with increase foaming tendency
•    As an operational task, regularly skim out hydrocarbon layer before it starts to build up

AMIT KELKAR (Shell Catalysts & Technologies)
The primary source of Silicon is the anti-foam injection at the coker. Silicon is also present in a wide variety of crudes including Maya, Canadian syncrudes and Venezuelan. It may show up in straight run naphtha due to use of Silicon based additives in upstream crude production and pipeline operations. 

One way to monitor breakthrough is to track cumulative Silicon content of feed compared to the design Silicon capacity of the installed catalyst system. This requires regular feed analysis. A reasonable estimation of feed silicon can be obtained by detailed spent catalyst analysis from prior cycles. Monitoring the progression of dT across the catalyst bed is a qualitative way to monitor Si breakthrough. The catalyst at the top of the reactor gets saturated with Si first resulting in shift of dT to the lower portion of the bed. This Si “wave” can be tracked by means of a plot of percent dT for each individual bed or even sections of the same bed provided adequate thermometry is available.
Another way to monitor Si is by means of a rigorous kinetic model while splits the catalyst bed into multiple “slices” and estimates an overall Si profile for different cycle lengths. 
Si deposition preferentially poisons the HDN function of the NHT catalyst. In some cases, product nitrogen slip precedes Si slip. Nitrogen specification for reformer feed is 0.1 – 0.5 ppm. Increased N slip strips the chloride off the reforming catalyst and affects the metal / acid balance. It leads to fouling of recycle compressor blades, stabilizer trays, feed / effluent exchangers and overhead coolers due to deposition of ammonium chloride (NH4Cl).
High silicon (target < 0.1 ppm) in the reformer feed blocks the active sites and reduces regeneration efficiency by impacting Pt dispersion and Cl adsorption. 

JOE RYDBERG (CITGO)
Silica capacity can become a run limiter especially on units that process Heavy Coker Naphtha.   Typically the reactors in these units will contain a significant amount of high surface area Silica Guard catalyst in addition to the high active main bed catalyst.  Silica is analyzed in the feed to ensure the unit is “on pace” with meeting its cycle length target.   Silica pickup is temperature dependent.  Work with your catalyst supplier to provide a total Silicon “trap capacity” number for the reactor and individually for the silica trap material, mainbed catalyst, and so on.  At the end of the run, the Silica lab numbers are totalized and compared to the weight of Silica the spent catalyst “captured”.  Each layer needs to be sampled and analyzed. Typically these compare pretty closely (10-20% or so).  Need to make sure not to confuse Silica with Silicon.    Work with your catalyst suppliers to potentially add additional Silica Guard catalyst and extend the next cycle.  Tracking Silica can be an effective method to predict and schedule catalyst changes several years out.   Looking at the Silica levels in catalyst samples at the bottom of the bed will also ensure that Silica hasn’t broke-through into the reformer. Typical consequences of Silica (which is a permanent poison) breaking through into the downstream reforming unit is loss of catalyst activity and yield

KA LOK (Honeywell UOP)
Regular monitoring of silicon levels in naphtha feed and reload of the hydrotreater catalyst as needed are tools to safe guard downstream catalytic units, primarily catalytic reformers.    Generally silicon in naphtha hydrotreater feed is from byproducts of delayed coker antifoam.  Analysis of the feed for silicon content should be on a composite sample collected over a coke drum cycle to determine a representative average concentration.  Cumulative mass of silicon throughout the hydrotreater cycle can be tracked against the silicon capacity of the reactor catalyst inventory represented by the catalyst supplier.  Although contaminant silicon will deposit in a relatively narrow band progressively through the reactor, a safety margin to account for catalyst bed bypassing and the length of the silicon removal zone should be included in estimating silicon breakthrough out of the hydrotreater.  Periodic sampling of reactor effluent for silicon analysis near estimated end of run can be used to verify the estimate. 

A hydrotreating process unit is commonly used to remove metals in the naphtha prior to feeding to a catalytic reforming unit, such as a UOP PlatformingTM unit. Reforming catalyst is very sensitive to silicon contamination.  Silicon is an irreversible support poison. It reduces chloride retention and accelerates platinum agglomeration. At elevated levels this could lead to poor catalyst performance, increase in cracking, lower C5+ yield and lower activity.  In addition, regular analysis of reforming catalyst for silicon concentration provides additional protection against poisoning.

SERGIO ROBLEDO (Haldor Topsoe, Inc.)
As stated in previous AFPM Q&A’s answer books (see 2014 and 2016), silicon will impact HDN activity to a greater extent than HDS (see Figure 1). Therefore, a good way to measure silicon pick-up is by tracking the feed and product nitrogen. The %HDN, and normalized WABT HDN, can then be calculated and the decline over time can be tracked.

Fig. 1 – Silicon impact on catalyst activity.

It is recommended that feed and product silicon be measured.  This will allow the pounds of silicon being placed on the catalyst to be calculated and help determine when a Si breakthrough will occur. Alternatively, antifoam usage at the coker unit can be used to estimate pounds of silicon entering the naphtha unit. Check with your chemical supplier as to 1) how much pounds of silicon are present in each gallon of antifoam and 2) what percentage of that ends up in your naphtha cut, based on cutpoint.

Another important factor is to understand the silicon capacity of the installed catalyst system and its pick-up profile. A standard/max theoretical pick-up number should not be used to estimate breakthrough. Silicon capacity is dependent not only on the catalyst installed, but feed and operating conditions of the unit it is installed in. Maximum silicon pick-up in a given unit is dependent not only on catalyst type, size and shape but also on 1) silicon content in the feed, 2) feed boiling point range and 3) reactor operating temperature (Figure 2).
 

Fig. 2 – Max silicon capacity.

How sharp the pick-up profile is, and as such what percentage of maximum silicon pick-up breakthrough will occur, depends not only on catalyst type, size and shape but also on 1) feed silicon content and 2) liquid hour space velocity, or LHSV (Figure 3). Please check with your catalyst supplier for the expected pick-up tailored to the respective unit conditions.

Fig. 3 – Silicon laydown profile.

Please see Haldor Topsoe’s previous AFPM responses on the subject of silicon poisoining for mitigation steps including proper catalyst design when processing silicon containing streams.
As for consequence to the downstream reformer, the silicon is a permanent catalyst poison that cannot be removed via regeneration.
 

GRANT YOKOMIZO and JOHN PETRI (Honeywell UOP)
A hydroprocessing unit can be designed to handle all potential feedstocks when there is a detailed understanding of the feedstocks’ characteristics. Specifically feedstock selection regarding incompatibility limits needs to be considered as a broad and complex balance inherently linking with design choice for hydrogen addition in distributed product streams. In reality the selection is subject to choice of reactor configuration (e.g., fixed bed, fluidized or slurry), process target (cracking conversion, or product heteroatom content) as well as process conditions (catalyst usage or life cycle for fixed bed, pressure levels). Frequently incompatibility can be more limiting depending on the objective of a process, e.g., deeper cracking conversion is frequently found more limiting when hydrogenation is aggressively sought for cleaner and more saturated product as design priority, exemplified by fixed bed resid hydrotreating and fluidized bed resid hydrocracking scenarios. Bitumen residues may be potentially high in particulates based on the process of removal from tar sands.  
 
In any residue conversion process, asphaltene precipitation can occur at higher conversions when the components that keep asphaltenes in solution are saturated and/or converted to lower boiling points.  In addition, the cracking of asphaltenes makes them smaller and more polar increasing the propensity to precipitate.  The feed selection in higher conversion processes needs to keep this phenomenon in consideration.  More aromatic feeds will help keep asphaltenes solubilized. Process conditions (severity, catalyst type and use) related to hydrogen addition to bulk product affect reactor solvency thus stability. Clarified slurry oils or other highly aromatic feeds can be added into residue processes with significant conversion to improve the solubility of asphaltenes and enable higher conversion with reduced precipitation.  However, clarified slurry oils may also contain significant concentration of FCC fines.  

 Bottoms sediment and water (“BSW”) analysis can be used to identify the potential range of concentration of particulates.  Laser refractometry can be used to identify the particle size distribution.   For example, a sample can be ashed, particulates suspended and sonicated, followed by particle size distribution measurement.  The pressure drop media can be selected and the pressure drop mitigation system can be designed to remove particulates externally or internally to the hydroprocessing unit.  Additional media ought to be carefully considered, pending characterization of other contaminants in the feed components, e.g. Si, As, Ni, V, alkali metals, etc.  Other feeds having a relatively higher Total Acid Number (“TAN”) can internally create metal carboxylates and internal particulate generation after conversion to the metal sulfide.
 
Bulk properties of vacuum residues do not necessarily indicate the entirety of coking potential and deactivation.  NMR and high resolution mass spectrometry of deasphalted oils and vacuum residue can be used to indicate the fused aromatic ring size and the extent of alkylation. The larger fused rings have a higher coking potential.