Question 20: What is the industry experience with managing cracked feedstocks to avoid hydrotreater problems such as fouling, and pressure drop?
OHMES (KBC Advanced Technologies, Inc.)
Again, a lot of these comments are in the Answer Book, so I will skip through part of my remarks. People often forget about the refinery slop processing system. Typically, refiners understand where their cracked stock goes; but if it is typically or abnormally routed to the crude tanks or units, one will have cracked stock throughout the refinery, which can cause problems.
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Some of the mitigation options can be summarized in four areas. The first is feed logistics. We would say that if you are going to handle this, you need to do as much direct feeding as possible. Again, it is a refinery; so, it is really not practical to be 100%. If you are going to use storage, all of the tank streams will need to be under a nitrogen blanket or a floating roof tank utilized, even if these are virgin streams that will mix with cracked stock. You will obviously need to check your tanks during regular maintenance and cleaning because, again, the primary causes and ingredients of this foulant formation are the olefins, diolefins, the iron scale around it that acts like a catalyst, and the oxygen.
Once that is under control, another option would be chemical treatment. Some people do put in oxygen scavengers; they will work. Others put in different inhibitors to stop the polymerization from forming, particularly on the initial exchangers. Again, there are mixed results in the plants. Some work; others do not. It is probably worth a trial to see if it fixes your situation.
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An important area of consideration is the unit configuration. The traditional example is determining how to handle coker naphtha. Some people will have the low temperature diolefin saturation reactor upfront, which can very effective. For those who do not, sometimes putting this coker naphtha in other spots (the response given earlier about putting it in the ULSD or the gas oil unit) is an option. In Asia, many refiners actually put in oxygen strippers to help manage this issue. They will bring in fuel gas, hydrogen, nitrogen, or other media to strip out the tankage feed to remove the oxygen.
The final recommendation concerns the catalyst; and, again, I will defer to the catalyst vendors. There are a lot of different options for catalytically avoiding the polymerization from forming at the top of the reactor. Some refiners include specialized iron trap catalysts which keep that from causing those formations.
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LEICHTY (Chevron USA, Inc.)
Chevron has a number of units that store cracked feeds ahead of hydroprocessing units. When doing so, care must be taken to ensure that the tanks are free of oxygen to mitigate unwanted polymerization and gum formation. With the more reactive stocks like coker naphtha, it may be necessary to inject an antioxidant into the rundown and an antifoulant in the hydrotreater feed to mitigate exchanger fouling.
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UNIDENTIFIED SPEAKER (Foster Wheeler)
Regarding the storage of cracked products, what is the maximum duration that we can normally store the cracked products? Can we take the cracked products from storage directly back into the hydroprocessing or should they go back to coker or somewhere else?
OHMES (KBC Advanced Technologies, Inc.)
As long as it is under a nitrogen blanket or floating roof, you should be able to store it for quite a while. I do not really have a good rule of thumb on storage time limits, but we have seen refiners go between turnarounds storing cracked stocks. But if you do not have that capability, the cracked products will quickly form polymers. You will then need to either try to rerun the material elsewhere or just take the material, knowing that it will cause problems in the units. Most refiners just get ready for it.
OHMES (KBC Advanced Technologies, Inc.)
Cracked stocks can come from conventional sources, such as coker, FCC, and visbreaker. However, cracked stocks can be found in unexpected streams, such as virgin stream where cracked refinery slop has been introduced to the upstream crude and vacuum unit. The hydroprocessing units most commonly impacted by cracked stock, from a fouling and pressure drop perspective, are the naphtha and diesel hydrotreaters. Some of the typical symptoms of fouling and pressure drop (DP) problems when running cracked stock are:
• Loss of heat recovery in the preheat train, such that the reactor charge furnace has to fire harder,
• Reduced hydraulic capacity, such that flow control valves operates further open or near full-open, and
• High pressure drop across the reactor or first bed, depending on indication points.
The normal cause of fouling and DP due to cracked stock is polymer formation, which is driven by the olefins and diolefins inherently in cracked streams, oxygen contamination in the feed, and catalyzed by iron contamination. The four management options are feed logistics, chemical treatment, unit configuration, and catalysts.
Feed Logistics
The primary cause of fouling from cracked stocks is O2 contamination. Therefore, all feeds to a hydrotreating unit processing cracked stocks should be sent from the source unit directly to the hydrotreater. If feed from tankage is sent to the hydrotreater, whether that feed is virgin or cracked, the tank should have a nitrogen blanket or should be a floating roof. Regular maintenance on the blanketing systems and floating roof seals are especially important for these services. Using the tanks in a “once-through” mode also assists in minimizing fouling.
As highlighted above, iron can serve as a catalyst to drive polymer formation. Some logistics mitigation options are to use feed filters to remove iron scale and to complete regular tank maintenance to remove corrosion products from streams coming from tankage.
Finally, ensuring that refinery cracked stocks are properly managed and processed is important. Several instances of high fouling have been traced to cracked slop being reprocessed in crude units, thereby “contaminating” virgin streams with cracked material. Normal practice is to segregate virgin and cracked slop and process cracked slop in the coker or send to fuel oil.
Chemical Treatment
Though chemical treatment is an option, the success rate is varied and depends on the severity of the fouling. For oxygen contamination, some refiners have used oxygen scavengers to contain the oxygen contamination. Other refiners have used polymerization inhibitors, filming antifoulants, and dispersant antifoulants to minimize the impact of fouling on the preheat exchangers. For reactors with pressure drop problems, chemicals are available for use to help break up the foulant and provide a temporary reprieve on the pressure drop. For these options, the best option is to discuss with your site chemical vendor to determine options and solicit positive result case studies or contacts.
Unit Configuration
Obviously, one option is to remove the cracked materials from the unit feed, but that option is normally not practical. However, an option may exist to process some or all of selected cracked streams in other hydrotreating units. For instance, some refiners process coker naphtha in the diesel or gas oil hydrotreater, as opposed to processing in a dedicated coker naphtha hydrotreater or mixing in the virgin naphtha hydrotreater, such that the material can be diluted in a larger feed pool.
Processing coker naphtha in conventional naphtha hydrotreating units is possible and common. However, the successful units include a diolefin saturation reactor. This reactor is loaded with low activity catalyst and operated at a relatively low temperature, as the diolefins do not require normal reactor temperature not normal activity catalyst to saturate. This configuration option allows the refiner to saturate the diolefins at low severity, thereby reducing the propensity to form polymers.
From a heat recovery standpoint, several options are available to minimize fouling impacts. First, ensuring some hydrogen is available in each of the preheat exchangers helps reduce polymer formation as the liquid feed is heated to reactor conditions. Most modern designs already have this feature, but older or revamped units can have liquid-only exchangers in the hot circuit that have a tendency to foul. Some licensors and refiners have installed the capability to take exchangers out of service while the unit is in operation to allow online cleaning or have sacrificial exchangers in the preheat train. Avoid using plate heat exchangers, which are prone to fouling. Finally, as a reminder for units which will operate in 100% vapor phase, make sure that the dry point does not occur in the charge furnace, so that deposits are not placed on the furnace tubes.
A final configuration options that some refiners have used is an oxygen stripper on the fresh feed. This option involves installing a low pressure stripper tower that uses hydrogen, natural gas, or fuel gas to strip out any oxygen contamination. However, this configuration option can be capital intensive and requires regular loss of off gas to fuel gas or flare to function properly.
Catalysts
Catalysts have long been the primary reactor pressure drop control mechanism. Most refiners now employ a catalyst grading system on the top bed of the reactor to provide a mechanism to trap contaminants and manage pressure drop escalation during a cycle. Specifically for fouling due to cracked stock, several considerations are suggested:
• In addition to catalyst size grading and progression, also ensure that catalyst activity is being progressed from low to high activity. This approach is similar to the discussions above on diolefin management.
• Special grading material is available that helps better trap and isolate iron scale. Doing so will reduce the tendency for iron catalyzed polymer formation. Discussions with your catalyst vendor about catalytic options should be part of your regular catalyst selection process.
ESTEBAN (Suncor Energy, Inc.)
Hydroprocessing of cracked feedstocks is very common due to the sulfur content of coker and FCC products. These feed streams can often be sent “hot” (directly) to a downstream unit, which prevents potential polymerization reactions due to oxygen ingress at storage tanks. However, storing these materials in feed tanks can result in reactor fouling and pressure drop increase.
Blanketing of tanks that store cracked feedstocks with nitrogen or some other gas is the primary control mechanism to avoid fouling and pressure drop problems resulting from the polymerization of cracked materials when oxygen is present. In some of Suncor’s facilities, cracked feedstocks are stored along with straight run feedstocks in nitrogen-blanketed tanks. In order to ensure proper operation the control instrumentation is inspected and maintained routinely to prevent fouling issues. In our experience, we have identified a malfunctioning N2 blanketing system as a contributing cause to exchanger fouling in a diesel unit. The malfunction had allowed oxygen to enter the tank and react with the cracked material in the tank.
A secondary means of preventing fouling issues is by the use of polymerization inhibitors available from chemical vendors. These are injected into the stream entering the feed tank. Most are approved for use, by licensors and catalyst manufacturers’. However, Suncor does not have experience with these products.
ROGER METZLER (Baker Hughes)
Avoid tankage, if possible. Direct hot rundown to the HDS unit. If intermediate tankage is necessary, nitrogen or fuel gas blanketing should be employed to minimize oxygen intrusion. Treatment with Baker Hughes antioxidants and/or polymerization inhibitors can be effective when the specific polymerization mechanism is identified in order to select the appropriate inhibitor chemistry.
Question 84: What measures can be taken to reduce the residence time in the FCC main fractionator to offset slurry circuit fouling? What basis is used to determine the residence time in the main fractionator bottoms?
BROOKS (BP Refining)
I would like to start with a little background on bottoms coke to present some clarity on the basis for my answer to this question. Bottoms coke sources could come from multiple places. But to address the question with respect to residence time, our impression was that we are mostly looking at the thermal cracking and decomposition reactions in the main fractionator bottoms as opposed to asphaltene precipitation or reactor transfer line coke generation and carryover.
BP has done a bit of internal research based on a 1988 paper by TOTAL that was presented in an AKZO catalyst seminar. We looked at the pool residence time, pool temperature, and asphaltene content. Based on that paper, we took those three pieces and combined them with some internal data to regress an equation. We then backed out a coking point temperature (based on about 0.45 wt% coke) as our coke formation point. We ended up with a coking point temperature, and that is what BP targets when we watch our main fractionator for bottoms coking. We do not focus as much on the residence time as we do on this coking point temperature. We have noticed that in this equation, when you look at the different variables of temperature, feed paraffinicity, and residence time, feed paraffinicity and temperature have much greater effects in our equation than residence time, which is why we focus on coke point temperature.
For most of our units, we use the same coke point temperature equation and track a certain temperature for the main fractionator bottoms pool. We will try to keep it about 10°F to 30°F below the calculated coke point temperature. That will then be one of our KPIs, or monitoring points, for our bottoms to help prevent coking in that circuit. Basically, to help control this, we focus on removing heat through the pumparound system. Also, 10 of our 11 BPoperated FCCs have an internal slurry quench system, which includes a slurry pumparound distribution ring near the top of the pool and a stripping steam ring near the bottom of the pool to provide extra agitation. Both of those rings will provide cooling and help us maintain the main fractionator bottoms’ pool temperature below the coking point temperature that I mentioned earlier.
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To go back and address the question, though, since it was not really about temperature, BP calculates our residence times based on the main fractionator bottoms’ head volume, and then we divide it by the DCO product flow rate out of the unit. By that definition, BP’s residence times for most of our units are in the hour to hour and a half range. We feel that this is a fairly good proxy for the residence time in the circuit without having to go develop detailed calculations around the volume in the piping or in the exchangers.
If you are considering reducing the residence time, then it is conceivable for you to do this by dropping the liquid level in your main fractionator bottoms, which would then lower the residence time in your main fractionator. And if you are doing any major modifications, you can adjust the volume of the main fractionator head, although that is typically not cost-effective.
When looking at residence time as opposed to temperature, consider that as you drop the time in this pool, you will have a greater chance of building up coke into the bottom of the pool, which will result in extra clogging problems in the main fractionator bottoms circuit. You should always consider your pumparound rates and the minimum pump NPSH (net positive suction head) on your bottoms pumparound circuit. Lowering your residence time could reduce the amount of the NPSH available. But typically, BP focuses more on the temperature and less on the residence time.
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PIMENTEL (CITGO Petroleum Corporation)
I have a little to add to what Halle said. CITGO’s basis for residence time calculation is the total volume from the normal fractionator operating level all the way to the inlet passes of the slurry feed heat exchanger divided by both the slurry and recycle. The sum of these two flows represents what is being purged out of the system. Again, there are not many options to reduce residence time without sacrificing the cycle oil yield than operating at your lowest possible fractionator level.
As for temperature control, we do find significant temperature variances between different points in the fractionator bottoms compartment after injecting the quench oil. So, I would recommend that you monitor temperature not only by the temperature of the product leaving the fractionator, but you should also take some readings around the bottoms to find the significant temperature variances.
WILLIAMS (KBR)
KBR designs for a one-minute holdup between the maximum and minimum levels based on the total slurry draw-off rate. Operating the tower at a lower level will reduce residence time within the column.
AVERY (Albemarle Corporation)
I agree with the comments made by Halle, Sergio, and Jesse. My detailed answer is in the Answer Book.
NAVEEN DIMRI (Reliance Industries Limited)
What is the typical CSO (clarified slurry oil) asphaltene content? Is there any limit for the fouling control?
BROOKS (BP Refining)
The equation we use does not include the asphaltene level specifically. The feed quality parameter is around the feed paraffinicity, which is obviously not really the asphaltene. We have that built in to account for feed quality. I have seen some other general references for the industry around feed quality. A lot of people will look at their minimum bottoms temperature and then subtract 20ºF or 25ºF off of that as an extra cushion around feed quality differences, if they have particularly nasty feed.
PIMENTEL (CITGO Petroleum Corporation)
We do not measure asphaltenes, but we do control the API. We have found that when we operate below -3 API or -4 API, the coking tendency increases significantly. So, we have set that as a limit for a minimum API of your slurry oil.
NAVEEN DIMRI (Reliance Industries Limited)
At BP, what is the asphaltene content you normally create?
BILL WILSON (BP Products North America Inc.)
Just to clarify, in this coking calculation, we actually looked at the paraffinicity of the feed. So that is the parameter. Normally, we do not actually measure the asphaltenes in the slurry or, in fact, even in the feed.
BROOKS (BP Refining)
Coke in the main fractionator bottoms section can originate from multiple sources. The key typical sources are coke formed in the reactor transfer lines, thermal cracking and decomposition in the main fractionator, and precipitation of asphaltenes. High catalyst fines in the bottoms can act as nucleation sites for the asphaltene precipitation and thus result in higher bottoms coke formation. Catalyst and coke entrainment can be considered as determined by factors outside the main fractionator. Thus, when focusing on controlling main fractionator bottoms coking, we focus on the key factors influencing thermal cracking and asphaltene precipitation: bottoms pool residence time, pool temperature, and asphaltene content. To avoid coking in the main fractionator bottoms section, we typically do not focus on the residence time as much as we focus on maintaining lower temperatures.
BP has done a significant amount of work internally around the factors that influence main fractionator bottoms coking. Based on a 1988 paper by Total at an AKZO catalyst seminar, the thermal decomposition of DCO is directly related to both residence time and pool temperature. BP used data from this study and additional internal data to regress a relationship between weight percent coke formation in the bottoms, residence time, and pool temperature. Additional studies followed which indicated that FCC feed paraffinicity also has a significant effect on these main fractionator coking reactions. Thus this factor was also added to our regression equation. For our purposes, BP assumes the critical coking point in the main fractionator bottoms at 0.45 wt% coke. The regression equation is solved for temperature at the 0.45 wt% coke level. This temperature is our “coking temperature.” The BP units regularly track the coking point temperature and work to maintain main fractionator pool temperatures 10°F to 30°F below this coking point temperature to prevent main fractionator bottoms coking. Typically, refiners have addressed this by specifying a maximum allowable bottoms temperature which is reduced by 20°F to 25°F for highly paraffinic feeds.
A holistic look at this equation suggests that the effects of temperature and feed paraffinicity on bottoms coking is much greater than the effect of residence time. Typically our main controls around temperature are to install slurry pool quench systems and focus on heat removal in the pump around circuit. Ten of our eleven FCCUs in the BP system have slurry pool quench systems to maintain low pool temperatures. These quench systems typically include a distribution system to deliver cooled slurry pumparound into the base liquid as evenly as possible. This stream should be submerged in the liquid, near the top of the pool. The system also typically includes a steam distributor near the base of the slurry pool to cool the bottom of the pool and aid in mixing the cooled slurry into the pool. This steam ring can have multiple benefits including:
- improving the stripping of the DCO and thus may also improve the DCO flashpoint,
- cooling the slurry in the base of the fractionator and thus reducing coking tendency, and
- providing agitation to help mix the quench and liquid in the fractionator base.
To focus on the base question around residence times, BP calculates the residence time by calculating the main fractionator bottom head volume based on liquid levels as the pool volume and dividing this by the slurry product rate. BP’s typical residence times are between one and 1.5 hours for this system. Some of our units run with residence times over two hours due to large main fractionator bottoms volumes and minimum bottoms make operations.
It is important to note that reducing the main fractionator bottoms volume in an effort to reduce residence time can have the potential to cause issues around build-up of coke shed from the transfer line. Without major equipment modifications, a site could conceivably reduce residence time by minimizing bottoms liquid levels and maximizing the bottoms pump around rates consistent with the bottoms heat removal required to maintain the desired fractionator heat balance. If major modifications or revamps are made to the tower (including a full tower replacement), and a goal is to reduce the residence time in the bottoms pool, modifications can be made to the shape of the bottom head of the fractionator. Changing the shape to a cone or other lower volume head will reduce the residence time. This is typically not a cost-effective modification in most situations which is why we focus on maintaining lower slurry pool temperatures (as mentioned before). If residence time is reduced, it is also important to review the pumps and revamp as necessary for lower NPSH required due to lower liquid head volumes.
PIMENTEL (CITGO Petroleum Corporation)
The basis for the residence time calculation are: the slurry oil flow (as produced from the unit plus flow recycled to riser) and the volume of the main fractionator bottoms compartment from the normal operating level plus the volume of the piping from the bottom of the tower to the inlet of the first set of heat exchangers. There are not many options to reduce residence time other than lower the main column bottoms level. We have found that maintaining a continuous flow of cold slurry quench into the bottom head of the main column reduces coke formation and subsequent plugging of the slurry pump screens or pumparound exchangers. This practice minimizes hot spots in the slurry liquid in the bottom of the tower where coke can form even if the bottom outlet temperature is at target (690°F max). If the bottom temperature drops below our target when the quench flow is at minimum, we reduce slurry pumparound duty at the steam generator. We have experienced slurry exchanger fouling due to the presence of excessive PNAs (polynuclear aromatics) and find that this often correlates with very low slurry API. This can be mitigated somewhat by sacrificing LCO recovery and dropping LCO into slurry.
AVERY (Albemarle Corporation)
Including the composition of the main column bottoms oil and the liquid temperature, an important variable impacting coking in the main column bottoms circuit is the hydrocarbon residence time in the bottom of the fractionator. Excessive time and temperature can lead to coke formation. The oil residence time can be determined from the volume of oil in the bottom of the tower and the bottoms flow rate leaving the tower. The residence time should be minimized by maintaining a minimum bottoms liquid level and maximizing the bottoms pumparound rate consistent with the bottoms heat removal requirement necessary to maintain the desired fractionator heat balance.
SUBHASH SINGHAL (Kuwait National Petroleum Company)
Lower fractionation, bottoms level, and optimum temperature (less than 365°C) are a few considerations to avoid slurry circuit fouling.
Question 22: Can you elaborate on the benefits, drawbacks, and trade-offs of liquid phase vs. gas phase catalyst activation in a hydrocracking unit? Is there an activity or yield difference of liquid phase or gas phase activation of 2nd stage catalyst in a 2-stage hydrocracking unit? Consider both catalyst formulation and operational factors (MPT, excursion risk, etc.)
JEFF CATON (Axens)
Activation (or sulfiding) of hydrotreating and hydrocracking catalysts can be done either ex-situ or in-situ. In both methods, the objective is to convert all the catalyst’s metal oxide sites to active metal sulfides (Refer to Figure 1 - Catalyst Sulfiding/Activation). A poor or insufficient activation will result in reduced catalyst activity and cycle length. This is applicable for nearly all hydrotreating and hydrocracking applications, including 1 and 2 stage hydrocrackers. Axens’ recommendations do not significantly differ when considering activation (or sulfiding) a hydrocracker 2nd stage over a hydrotreater or hydrocracker 1st stage.
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For ex-situ catalyst activation (commonly referred to as pre-activation), the catalyst is fully activated via conversion of metal oxides to active metal sulfides. Note, this is not the same as ex-situ pre-sulfiding, where the catalyst is more simply loaded with sulfur, but still requires an in-situ activation, or conversion of the metal oxide sites to active metal sulfides using the pre-loaded sulfur. Pre-activated catalyst is delivered to the site in drop-in, ready to use form. In addition to standard pre-activation, the catalyst can be additionally pre-treated to allow for immediate processing of cracked feedstock, and if applicable, can be pre-passivated in the case of hydrocracking catalysts. Additional treatment to allow for atmospheric loading of the catalyst is also available. The disadvantage to pre-activation of catalysts is relatively limited to the cost of pre-activation, however, this cost is usually offset by the decreased loss-of-production time realized through elimination of in-situ activation (sulfiding) steps, which can commonly take 24-36 hours to complete. Axens’ experience with the use of pre-activated catalysts has been exceptionally good, with most start-ups progressing quickly and smoothly followed by catalyst performing as expected.
In terms of in-situ catalyst activation (sulfiding), the sulfiding can be done in liquid phase or in gas phase. Further, the sulfiding can be done through injection of a sulfurous agent (e.g., DMDS, Sulfrzol, etc.) or via native feed/sour gas sulfiding, assuming a high enough sulfur content in the feed or sour gas. All these methods are possible, and when completed according to recommended guidelines and best industry practices, with the proper safety and emergency scenarios considered, the catalyst activity and catalyst cycle length has achieved the expectations.
Axens’ position is that liquid phase sulfiding is more controllable and robust, therefore less prone to unforeseen issues that might impact the overall catalyst activity and stability when compared with gas phase catalyst sulfiding. The primary benefit of liquid phase sulfiding is the controllability of delta temperature (dT) caused by the exothermic sulfiding reactions. Early in the in-situ sulfiding process, if the feed sulfur concentration is too high, the exotherm could result in a premature temperature increase creating a risk of coking or metal oxides reducing to the metal state. If this were to occur, this could substantially impact the catalyst activity and stability, ultimately leading to poor yields and/or a shortened cycle length. With liquid phase sulfiding, this risk is better managed as the liquid acts as a heat sink, thus dampening temperature spikes caused by the exothermic reactions. Similarly, Axens favors injection of a sulfiding agent versus native feed/sour gas due to the ability to better control the sulfur content going to the reactor through control of the sulfiding agent injection rate. If the dT is too high and excessive hot spots occur, it is relatively manageable to decrease the sulfiding injection rate to decrease dT. Conversely, with native feed/sour gas sulfiding, it is not as straightforward to monitor or control the amount of sulfur going to the reactor, so the dT is less manageable. Further, based on Axens’ experience, sulfiding in liquid phase with injection of a sulfiding agent has generally led to shorter duration of the sulfiding procedure.
A risk during in-situ sulfiding is the loss of sulfur to the reactor caused by loss of sulfiding agent injection. If there is a loss of sulfiding agent injection during the higher temperature step, there could be a resultant reduction of metal sulfides back to metal oxides and then further to the metal state. With liquid phase sulfiding, it is relatively more manageable to cool the reactor temperature to a lower temperature where the risk of reduction is minimized. With gas phase sulfiding, it is more difficult to decrease temperature of the reactor in a timely manner. As a precaution to counter the risk of loss of sulfiding agent injection, Axens recommends that the oil and/or makeup hydrogen used for sulfiding contain a minimal amount of sulfur and that the recycle gas amine scrubber be idled throughout the sulfiding process. These precautions will ensure continuous sulfur levels in the system thus minimizing the risk of catalyst reduction. Axens recommends that the recycle gas H2S content not fall below 1,000 vppm H2S when the reactor temperature is above 400 oF.
Like the ability to cooldown the reactor, it is also found to be more manageable to heat-up with liquid phase sulfiding. This can save time during the initial heat-up of the reactor to the first temperature hold and subsequently to the second temperature hold. Even when the refiner is limited to complete sulfiding in gas phase, Axens would recommend circulation of treated diesel throughout the unit to ensure wetting of the catalyst pores and to provide heat sink as discussed above.
It is important to note that in-situ liquid sulfiding should be completed with a 100% virgin (straight run) feedstock. After full completion of in-situ sulfiding, it is recommended to operate the unit with 100% straight run feedstock for a period of 3 days and then gradually introduce cracked feed stock. Further, the initial operating temperature should be maintained below the expected SOR temperature and then slowly increased as required to meet product specifications. Operating the unit with cracked feedstock within the first 3 days, increasing temperature too rapidly, and over treating may result in premature catalyst coking and resultant loss of activity and cycle length.
NEIL HOWARD (Chevron Lummus Global)
For current generation catalyst CLG has experienced comparable catalyst performance with either gas phase or liquid phase sulfiding procedures. There are some units that require gas-phase sulfiding, mostly because of the mechanical configuration – high MPT reactors, inadequate heat input, inadequate recycle compressors, etc. Chevron’s experience suggests that gas-phase sulfiding can be done safely and timely and without causing any yield differences, even in a second-stage reactor. We have observed that some older catalysts give slightly worse start-of-run performance under gas-phase sulfiding, but many of the most modern catalysts show little or no reduction in activity with gas-phase sulfiding.
Liquid phase vs. Gas phase startup
In general liquid phase sulfiding is preferred for reasons already stated above. Some of these liquid phase advantages include:
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better heat sink during the sulfiding steps to manage temperature rise in the catalyst beds.
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faster response to temperature increases or decreases in the event of a need to suspend sulfiding activities in the course of the startup.
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faster heat-up to MPT on units where startup oil is introduced at low pressure under nitrogen.
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no prolonged cool-down to feed-in temperature after high temperature sulfiding
Some of the concerns with heat release with gas phase sulfiding mentioned above can be mitigated by:
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Maximizing recycle gas rate to improve distribution and carry heat from the catalyst beds during sulfiding
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Limiting initial DMDS injection to maintain H2S in the recycle gas to no higher than 0.5 vol%. The reactor temperature rise should be limited to 50 degF
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DMDS injection can be increased to 1 vol% H2S in recycled gas after the initial temperature wave passes through the reactor. But again, dT should be limited to 50 degF maximum.
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sulfiding can be conducted at low pressure to avoid delays due to heating up to MPT and going to full pressure. 400 psia minimum H2 partial pressure is recommended, or as high as allowed by MPT.
MAUREEN BRICKER and JOHN KULACH (UOP)
This question requires some background on catalyst manufacturing and studies that we have performed utilizing both activation procedures on various hydrocracking catalyst types. As catalyst performance is being optimized to enhance yield and targeted hydrogen saturation so too are the methods employed for catalyst formulation and manufacture. Each step of the process is studied and optimized for the exploitation of the catalyst performance. Optimization techniques in manufacture and the activation of the catalyst in-situ are linked to catalyst performance. While the gas phase activation tends to maximize the metal sulfide uptake there is a clear advantage with liquid sulfiding pertaining to maximizing the yield and in many cases the activity of the catalyst system. During liquid phase sulfiding, there is some modest coke formation even at mild liquid sulfiding conditions. The type of coke formation occurring during the activation phase likely prepares the catalyst better for initial design feed by preventing initial over activity and higher rates of coke formation during early stages of the cycle. It could also attenuate the most acidic sites on the catalyst which would deactivate immediately in process feeds.
Liquid sulfiding procedures may be more difficult to enact in older units where the minimum pressurization temperature (MPT) may be higher than the potential catalyst reduction temperature in hydrogen atmosphere. Since liquid phase sulfiding is performed under hydrogen atmosphere, there is concern that heating the catalyst under hydrogen to achieve minimum pressurization temperature (MPT) could result in catalyst reduction. This is not a concern with gas phase catalyst activation. UOP has developed procedures for those reactors requiring higher minimum pressurization temperatures to enable use of liquid phase catalyst activation.
Regarding temperature excursions: In gas phase and liquid phase sulfding, the sulfding reaction results in axial rise across each bed of catalyst. In liquid phase sulfding exotherms can be seen additionally from adsorption (wetting) and from cracking. Procedures must take these possibilities into account and provide avoidance or action to be taken.
Question 62: Do you know of factors that are likely to lead to deposit formation on power recovery turbine blades? Is there anything that can be done to prevent these deposits from laying down on the blades? Once the deposits have been formed, what are the consequences and is there any way to remove the deposits online?
BHARGAVA (KBC Advanced Technologies, Inc.)
In summary, it is all about turbine blade deposition in the expander on the flue gas. First, I will talk about the causes, catalyst loading being the number one and the only cause for most of the turbine blade deposition. The catalyst loading on the flue gas inlet is what determines the amount of deposition. From a mechanical standpoint, the blade deposition increases as the performance of a third-stage separator goes down – for whatever reason – or the loss in regenerator cyclone efficiencies. From a catalyst perspective, if you are trying to put in a new catalyst, that catalyst will have different attrition properties. Evidence of mechanical damage in the unit that results in more catalyst fines will have the same effect. Increasing fresh catalyst additions will produce the same result, because fresh catalyst contains more fines. Also, as you start processing resid or heavy metal gasoil and your sodium and vanadium levels go up, the catalyst will start to get stickier and will create eutectic mixtures at certain temperatures, resulting in more catalyst sticking on the blades. From an operational standpoint, if you increase flue gas rates when you increase air rates, the amount of catalyst loss through the cyclones will increase and result in more deposition.
At some sites, we found additional steam being injected upstream of the expander, a practice that is done for different reasons. One reason could be to maintain temperature and pressure on the inlet of the expander in order to keep up the efficiencies, but that is not a recommended solution because it actually makes the situation worse on the turbo expander. So, those were the causes.
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What is mitigation? It is important to understand the cause of the depositions. You want to make sure you analyze the deposits. If you do not have that luxury, then track the 20-microns or less size range, the catalyst’s physical properties, and monitor the e-cat metals. From a mechanical perspective, you want to do a routine monitoring of the bearing temperatures and vibration, and then check the process temperatures and pressures to identify if you are having a problem. One option to help reduce the turbine blade deposition is to run close to the design temperature and pressure.
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Consequences: These are obvious consequences. You lose power generation efficiency, but more serious are excessive depositions. If you have deposition on the blade and have uneven breaking off the deposits, the expander can become unbalanced, which goes back to the previous point about monitoring vibrations for early detection. How do people remove the deposits? First, you need to monitor the deposits via a viewport using a strobe light. This light will allow you to quickly detect the buildup of deposits. This monitoring is more important because if the deposit just builds up, it will be easier to remove the deposits. You can even do that with an online riser walnut shell cleaning. Some people have resorted to thermal cycling by bypassing the expander. Again, that is also a thermal shock to the unit, so we do not recommend it. Finally, if the deposits have been there for a long time, you do not have any choice but to shut down the expander.
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FEDERSPIEL (W.R. Grace & Co.)
Like what was previously mentioned, some of the hard deposits that may form on the expander blades might consist of fine catalyst particles but may be enriched with other contaminants like sodium, potassium, calcium or chlorides, vanadium, iron, and other trace elements. The theory is that they might form a eutectic that drops out on areas of high velocity and pressure drop. It is a little counterintuitive; but the expander, of course, is one such place. Further, if you build up the deposit to a sufficient thickness where it starts to cause friction on the machine, those deposits can then be enriched with the expander metallurgy. And if they get hot enough, those deposits could then sinter and be very hard to remove.
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So, one of the key takeaways is to send in the deposit for analysis. We can do chemical analysis and look at what material is actually there. There are more advanced techniques as well, like microprobe or line scanning, which can tell us where the different elements are lining up in that deposit or how they are formed over time. Even SEM (scanning electron microscopy) and X-ray diffraction can look at not just the shape of the deposit but can also identify crystal structure.
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PUI-NANG LIN [Flint Hill Resources (FHR)]
Another area we found very important for the expander fouling is the quality of your expander cooling and impingement steam. That is another source of sodium that can accelerate the blade deposits.
ALEXIS SHACKLEFORD (BASF Corporation)
Another element you should look for is sulfur. Sulfur is often enriched in these deposits. Occasionally, you may also see evidence of refractory in these deposits. Please see BASF’s response in the Answer Book that shows you what these deposits look like compared to e-cat and compared to fines samples.
MELIKE YERSIZ (Chevron U.S.A., Inc.)
How often do you recommend inspecting the blades with the strobe lights?
BHARGAVA (KBC Advanced Technologies, Inc.)
If you have a viewport in a strobe light, then you should be routinely monitoring the blades once every couple of weeks, depending on the severity of the situation. You can then increase or reduce the frequency, depending on how your deposition goes.
FEDERSPIEL (W.R. Grace & Co.)
When I worked at Hovensa, we would do it weekly. We often found out that staying ahead of the problem was a lot easier than trying to address it after it became an issue. So, if you have the facilities there that look to the viewport or take the pictures, it is probably easiest to set it up on a regular basis – like on Saturday – just to have the Inspection guys go out, take the pictures, and monitor it.
PHILLIP NICCUM (KP Engineering, LP)
I want to make a reference to a paper written by David Linden with Ingersoll Rand back in the 1980s. The topic of the paper was the composition of these deposits. It is a seminal work on this subject, and I recommend it. In the paper is a reference to tables of these eutectic mixtures, and I have some of them. It is pages and pages of eutectic mixtures with many elements from FCC with which you are very familiar. It is quite a useful reference.30
BOB LUDOLPH [Shell Global Solutions (US) Inc.]
Calcium and iron also play into those eutectics as well and can have a dramatic effect. As far as the sodium goes, make sure your desalter is being checked for its effectiveness, because a dramatic shift in its performance could really result in a much larger change in the expander operation.
SANJAY BHARGAVA (KBC Advanced Technologies)
Deposit formation is usually linked to catalyst depositing on the turbine blades. The deposits are mostly a function of catalyst loading of the inlet flue gas. An increase in catalyst loading could be due to several factors, including performance of the third-stage separator, loss in cyclone efficiency, change in catalyst attrition properties, excess fines in fresh catalyst, an increase in fresh catalyst additions, and/or an increase in flue gas rates due to higher air rates.
In addition, KBC has seen locations where steam is introduced between the regenerator outlet and tunable diode laser spectroscopy third-stage separator. The addition of steam is used to control the temperature within the guidelines of the expander. The use of steam and the added “humidity” increase the deposition potential of the catalyst on the blades. Further, high levels of sodium and vanadium on equilibrium catalyst can also form a sticky, eutectic mixture which would tend to stick to the blades more easily and lead to accelerated deposition rates.
Rigorous monitoring of catalyst balance and fines generation – specifically, sub-20-micron particles – helps us understand the deposition rate on the blades. Minimizing deviations from design pressure and temperature also helps reduce the rate of deposition. Monitoring catalyst physical properties and Na/V (sodium/vanadium) to ensure good cyclone efficiencies is as important. Early or regular action to correct the deposition problem by routine monitoring of bearing temperatures and vibration, along with process temperature and pressure, will provide the required information for corrective course of action.
Deposits can be monitored via a view port with a strobe light to allow weekly photographing of the blades. This information can then be used to establish the frequency of regular online cleaning of the blades. Cleaning can be done with rice, walnut shells, or a less preferred method of thermal cycling of the turbine by partial bypassing of the flue gas to cool down the blades. If the deposits are allowed to build up over extended time periods, online cleaning is not recommended as chunks of deposits are likely to be removed non-uniformly, which can unbalance the blades and result in excessive vibration.
MICHAEL FEDERSPIEL (W.R. Grace & Co.)
The paper, “Catalyst Deposition in FCCU Power Recovery Systems” by David H. Linden31 at Ingersoll-Rand, refers to four types of deposits that occur in flue gas lines, equipment, and power recovery turbines. The first type (A) are powdery catalyst deposits that cling to surfaces in the flue gas train. The second type (B) occurs when those powdery catalyst deposits get wet and then harden after drying out. Upon analysis, these deposits appear similar in chemical makeup to equilibrium catalyst or third-stage separator fines.
The third type (C) of deposit is made up of very small catalyst particles, along with elevated levels of alkali metals (sodium, potassium, and calcium), chlorides, vanadium, and iron, as well as other trace contaminants from unique or challenging feedstocks. It is theorized that these contaminants form a low melting point eutectic. These deposits are very hard and are often found in areas of high gas velocity and pressure drop, and the expander is just such a place. If these deposits form along the expander blade tips to sufficient thickness to cause rubbing, the friction will increase the temperature enough to sinter the deposit into a new type (D) of deposit which has increased metals content from the expander metallurgy.
By reducing catalyst traffic to the flue gas and preventing condensation, these types of deposits can be minimized. Keeping the regenerator cyclones and third-stage separator mechanically healthy and operating within design specifications will help. Ensuring catalyst coolers are leak-free will prevent boiler feed water chemicals from further contaminating catalyst and eliminate an attrition source. Running the expander at design conditions can keep the flow path through the expander fully developed, thus mitigating eddy currents and dead spaces. Suppressing the contaminant metals (particularly those that accumulate on the surface of the catalyst particle and then abrade off) or using a catalyst with a low attrition tendency can help to reduce blade depositions. There are also several options for expander coatings which can reduce or prevent the accumulation of deposits.
Monitoring deposits during the run can help a refiner be prepared for work that needs to be done during the next available outage. Taking pictures through view ports and vibration monitoring are two common methods used to quantify and track deposit formation.
Once formed, deposits can cause loss of expander efficiency and threaten the mechanical integrity of the machine, as well as force a shutdown due to high vibrations. While these deposits can be removed through the injection of walnut shells or rice, a regular program aimed at preventing the formation of deposits is generally more effective at achieving longer run lengths than attempting to fix a vibration issue after it develops. Varying the size of the media can help reach the different places these deposits form. Other methods include thermal shocking or thermal cycling of the expander, which takes advantage of the different thermal expansion coefficients of the deposits and the turbine metallurgy to release accumulated deposits from the surface of the blades.
ALEX MANNION (BASF Corporation)
Below are deposit examples from a cyclone and expander blade. Typically, these deposits are rich in elements such as Fe (iron), Ca, Na, Mg and S. These elements can act as a “glue,” binding catalyst fines together.
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Factors that could lead to deposit formation include high catalyst attrition, high catalyst losses from cyclones, poor water quality being used from steam injection, water condensation and precipitation at cold spots, and high metals content in the FCC feed.
Several measures can be taken to avoid deposits from laying down on the blades. To minimize fines generation, an attrition-resistant catalyst can be used, and fresh catalyst additions should be minimized. Also, all steam ROs (restriction orifices) should be in place, and excess velocities should be avoided. High-quality water should always be used while minimizing water injection, when possible. “Cold spots” in the overhead line or expander should be avoided. Crude desalting additives can be used to minimize metals (e.g., Ca) in the feed. Finally, effective soot blowing of CO boiler tubes can mitigate deposit formation as well.
Deposits can lead to expander vibrations and blade erosion, potentially leading to catastrophic equipment failure. If deposits have begun to form, regular walnut shell cleaning can be conducted and thermal spalling if required.