Catalagram. A Refining Technologies Publication. No. 110 / Fall 2011 / Hydroprocessing Catalysts from The Chevron &Grace Joint Venture

Transcription

1 No. 11 / Fall 211 / Catalagram A Refining Technologies Publication Hydroprocessing Catalysts from The Chevron &Grace Joint Venture

2 Happy 1th Anniversary, ART The ART leadership team (left to right): Scott Purnell, Ryan Heaps, Bruno Tombolesi, Chris Dillon, Woody Shiflett, John Creighton, Kristen Kopp, Chuck Olsen, Lauren Blanchard, Darryl Klein, Cecilia Radlowski, Nathan Carpenter, Babu Patrose, Mark Peterson, Kaidong Chen, Charles Wear, Balbir Lakhanpal and Maureen Birnbaum. This past spring, we celebrated the 1th anniversary of the formation of Advanced Refining Technologies LLC, the hydroprocessing joint venture between Grace and Chevron Products Company. The venture was created in 21 to develop, market and sell a comprehensive line of hydroprocessing catalysts. ART combines Grace s material science, manufacturing, marketing and sales strength with Chevron s extensive experience in operating its own refineries and leadership in technology, design and process licensing. Through new product innovation and business development, ART has grown significantly since we were formed. We acquired the Orient Catalyst Company s hydroprocessing technologies and the HOP catalyst product line; we have become the largest shareholder in Kuwait Catalyst Company; and we have acquired and integrated new catalyst technologies from several companies, including Crosfield and Japan Energy. But most of all, thanks to you, our customers, we have become a leading global supplier of hydroprocessing catalysts with a complete line of products designed for processing resid feedstocks in fixed-bed, moving bed and ebullated bed units. We also offer a full line of distillate catalysts, some of which are represented in the pages of this publication, most notably the SmART Catalyst System for ULSD processing. We look to the future with enthusiasm and optimism. Our foundation will continue to be our commitment to the global refining industry to provide proven industry-leading catalysts and technical service. Sincerely, Scott K. Purnell Managing Director Advanced Refining Technologies

3 2 Hydrocracking Pretreat Catalyst Development By Jifei Jia and Dave Krenzke, Advanced Refining Technologies and Theo Maesen, Chevron Lummus Global No. 11 / Fall 211 / Catalagram A Refining Technologies Publication Hydroprocessing Catalysts from The Chevron & Grace Joint Venture 6 The 9 StART Challenges of Processing FCC LCO By Charles Olsen, Brian Watkins and Greg Rosinski, Advanced Refining Technologies Catalyst System Success Story By Geri D Angelo, Advanced Refining Technologies Catalagram 11 ISSUE 211 Managing Editor: Scott K. Purnell Technical Editor: Charles Olsen Contributors: Geri D Angelo Jifei Jia Dave Krenzke Charles Olsen Greg Rosinski Brian Watkins 12 Controlling Feedstock Contaminants in Diesel Hydrotreating Operations By Dave Krenzke, Advanced Refining Technologies Guest Contributors: Theo Maesen Please address your comments to: betsy.mettee@grace.com Advanced Refining Technologies 75 Grace Drive Columbia, MD NPRA Q&A Answers 211 Advanced Refining Technologies, LLC

4 Hydrocracking Pretreat Catalyst Development Jifei Jia Lead Research Engineer Advanced Refining Technologies Richmond, CA, USA Theo Maesen Team Leader Distillates Chevron Lummus Global Richmond, CA, USA Dave Krenzke Regional Manager Hydroprocessing Technical Service Advanced Refining Technologies Richmond, CA, USA Introduction Demands for a cleaner environment have led to more stringent global fuel specifications. The primary target has been a reduction in sulfur content. Currently on-road diesel fuel must contain no more than 1 parts per million by weight (wppm) sulfur in both the United States and the European Union. This standard also exists in a number of Asia Pacific countries. Because the production of ultra low sulfur fuels is primarily met by hydroprocessing and hydrocracking, continuous improvement in catalyst technology is needed. In 21, Grace and Chevron combined resources to form Advanced Refining Technologies (ART) in order to provide world class hydrotreating catalyst technology to the refining industry. Chevron also has a long standing partnership with Lummus, a leading international engineering company, called Chevron Lummus Global LLC (CLG). ART and CLG offer a full product line of premium catalysts for upgrading products varying from heavy oil to distillates. Distillate hydrotreating operations include ultra low sulfur diesel (ULSD) and FCC and hydrocracker pretreat. Refiners, such as Chevron, use these catalysts to remove sulfur and other contaminants from petroleum to economically produce more environmentally friendly transportation fuels. As the most completely integrated source for hydroprocessing technologies and services, ART and CLG can provide incremental efficiencies at every step in a project. This paper focuses on ART and CLG s latest hydrocracker (HCR) pretreat catalyst development. Hydrocracking Pretreat Catalyst Development Good hydrodenitrogenation (HDN) activity is the primary function of the HCR pretreat catalyst because organic nitrogen compounds are detrimental to the performance of the HCR catalyst downstream. The rate limiting step in the HDN reaction pathway is aromatic ring saturation. This is because the most refractory nitrogen molecules are compounds like substituted carbazoles in which the nitrogen atom is incorporated into the aromatic ring at a fairly inaccessible position. The development of an improved catalyst therefore needs to focus on the catalyst properties that enhance ring saturation. The two critical components for optimizing catalyst performance are the support properties and active metals deposition technology as discussed in Reference 1. 2 SPECIAL EDITION ISSUE No. 11 / 211

5 Catalyst Support Development 14 Small pores result higher surface area giving a higher Large pores result in lower surface area, but offer better 12 Optimizing the catalyst sup port starts with the relationship between the physical structure of the support and the catalytic activity for the hydrocarbon stream being processed. Figure 1, from Reference 1, demonstrates this relationship. Relative Activity Decreasing access to active sites Increasing access to active sites Relative Activity The physical properties of the optimized support determine the maximum useful metal loading and consequently the maximum potential activity. Subsequent steps in the catalyst development process seek to utilize as much of this potential activity as possible. One aspect of this process is to look at the metal-surface interaction. Altering the alumina surface chemistry with inorganic additives can have a significant effect on active site formation and has been the subject of an ongoing investigation by ART and CLG scientists. ART s 59DX has benefited from this study and utilizes a proprietary inorganic promoter to facilitate the formation of active sites during activation. An example of this work is shown in Figure 2. Metals Deposition Technology The current generation of HCR pretreat catalysts relies on the formation of DX active sites during manufacturing and activation to improve catalyst activity. The DX technology utilizes chelates to optimize metal function FIGURE 1: Relationship Between Pore Size and HDN Activity for VGO and consequently the activity and number of active sites. The chelate binds preferentially to the transition metal ion, Ni, and controls the sequence of metal ion adsorption during the impregnation step. This allows the Mo to adsorb first, followed by Ni, and reduces the chance of the Ni interacting with the alumina support. The Ni associated with the alumina will ultimately form NiS which is inactive for HDN. During activation, the chelate delays the sulfidation of Ni relative to Mo. This allows the complete formation of the MoS 2 slabs before the Ni ions are sulfided. Once Ni is released by the chelate during sulfidation it moves to the edges of the MoS 2 slabs to form the highly active Type II sites. There are many chelates available and they yield different performance benefits for the resulting catalyst. The ultimate choice is based on a combination of catalyst activity and manufacturing compatibility. Figure 3 demonstrates the impact of the chelating agent on catalyst performance. HDN/HDS Activity Improvement, F +1 Base Catalyst Without Surface Modification HDN Catalyst With Surface Modification HDS +6 Base FIGURE 2: VGO HDN/HDS Performance of Catalysts With and Without Surface Modification HDN/HDS Activity Improvement, F +1 Base Catalyst with Organic Chelating Agent B Catalyst with Organic Chelating Agent A Catalyst without Organic Chelating Agent HDN HDS FIGURE 3: Chelating Agent Effect on HDN/HDS Performance of VGO +6 Base HDN/HDS Activity Improvement, C HDN/HDS Activity Improvement, C ADVANCED REFINING TECHNOLOGIES CATALAGRAM 3

6 ART 59DX Commercial Performance 59DX was commercialized in the middle of 29. By mid 21 it had already been selected by seven refineries for catalyst refill. Figure 4 shows the HDN performance comparison between 59DX and the previous generation HCR pretreat, NDXi, in one of the Chevron refineries. 59DX is 2ºF (11ºC) more active than NDXi after 15 days on stream. WABT, F NDXi 59DX WABT, C In one of the early uses of 59DX, there was an upset during the start-up which resulted in an extensive period without hydrogen flow. This compromised the start-of-run activity of 59DX. Despite the compromised start-up, the 59DX catalyst continued to activate, and became as active as the NDXi catalyst in the prior run at 24 MBBLs on stream. Considering that the start-up of the prior run was not compromised, and that the feed in the prior run was less refractory than in the current run, it is clear that 59DX is a remarkably robust catalyst. The commercial results are summarized in Figure 5. Base Days on Stream 3 Base FIGURE 4: HDN Performance Comparison Between NDXi and 59DX As a point of comparison, Figure 6 shows a different commercial operation for 2 cycles using the same catalyst grade. One cycle started-up normally while the other cycle had a significant upset during sulfiding. This data demonstrates a more typical result for a significant upset during start-up. As can be seen in Figure 6, the major upset during start-up resulted in a significant permanent loss of activity and a much higher deactivation rate. This highlights the robust response of 59DX to operational upsets even during the critical start-up phase. WABT, F Base 59DX NDXi Base WABT, C Summary The formation of the ART joint venture in 21 has resulted in a significantly increased rate of innovation for HCR pretreat catalysts as can be seen in Figure Total Feed MBBLS -22 ART s NDXi, 59DX and the prototype for 591DX are the catalysts in this series which utilize DX technology. NDXi is ART s previous generation catalyst for HCR pretreat applications. The catalyst is also used in ULSD applications. It has been selected more than 4 times for catalyst refills and start-ups by many refineries since its commercialization in 26. Many of the applications were selected through competitive pilot plant testing. NDXi has demonstrated a significant HDN activity advantage over the earlier generation catalyst, ART s AT58. The base metal loading and alumina support is similar for AT58 and NDXi which confirms the advantage of chelate technology to more efficiently utilize the active metals. FIGURE 5: Commercial HDN Performance Comparison Between NDXi After Smooth Start-up and 59DX After Compromised Start-up 4 SPECIAL EDITION ISSUE No. 11 / 211

7 WABT, F Base Base Days on Stream WABT, C HDN Activity Improvement, F Base AT58 NDXi 591DX 59DX Base HDN Activity Improvement, C Upset During Sulfiding Normal Start-up FIGURE 6: Comparison of Typical Commercial Performance Between Normal Start-up and One With a Major Upset During Sulfiding FIGURE 7: Hydrocracker Pretreat Catalyst Generations ART is now commercializing 591DX, a wider pore version of 59DX, to handle VGO feeds with higher endpoints. The modified support properties will provide equivalent activity to 59DX but enhanced stability with heavier feeds. References 1. Krenzke, Dave; Vislocky, Jim. Hydrocarbon Engineering (27), 12(11), ADVANCED REFINING TECHNOLOGIES CATALAGRAM 5

8 The Challenges of Processing FCC LCO Charles Olsen Director, Distillate R&D and Technical Services Brian Watkins Manager, Hydrotreating Pilot Plant & Technical Service Engineer Greg Rosinski Technical Service Engineer Advanced Refining Technologies Chicago, IL, USA FCC LCO has long been a common component of feed to diesel hydrotreaters. More recently, there has been greater interest in processing higher quantities of LCO due to economic considerations and to meet the market demand for ULSD products. LCO has a number of impacts both on the performance of the hydrotreater and on the resulting ULSD product properties. The extent of the impact depends upon a number of factors including the amount of LCO in the feed and the catalyst system used in the ULSD unit 1. It is generally accepted that the addition of LCO to the diesel feed decreases the feed reactivity and requires an increase in reactor temperature in order to meet the product sulfur target. Figure 8 summarizes pilot plant data demonstrating this effect. The figure shows the required temperature increase relative to the straight run feed as a function of product sulfur for feeds containing 15 and 3% LCO. It is clear that even low levels of LCO impact catalyst activity. At higher product sulfur about 2 F (11 C) higher temperature is required for 15% LCO and this increases to 4 F (22 C) higher temperature for 3% LCO relative to the straight run feed. As shown in Figure 8, the activity difference is even greater at lower product sulfur. In addition to the amount of LCO, the endpoint also has a significant influence on the reactivity of the feed. Figure 9 compares the performance of two 3% LCO blends relative to the straight run feed. The endpoints for the 2 LCO s differ by just over 4 F (22 C). This difference corresponds to an additional 8 ppm of hard sulfur in the feed for the higher end point LCO, in addition to a 15% increase in PNA s. The LCO endpoint effects are apparent even when blended into the straight run feed. The higher endpoint feed requires about 2 F (11 C) higher temperature relative to the lower endpoint feed at higher product sulfur, and nearly 3 F (17 C) higher temperature for ULSD sulfur levels. Processing LCO also has a major impact on the product quality of the hydrotreated diesel. A significant problem relates to the aromatics content, as LCO s tend to have very high concentrations of naphthalene type aromatic species which have very low cetane numbers causing the LCO to have relatively low cetane. The high PNA content also has an impact on diesel product color. This becomes important as end of run (EOR) is approached since ULSD units processing LCO blends will have product go off color at a lower temperature relative to a SR feed. 6 SPECIAL EDITION ISSUE No. 11 / 211

9 A survey of commercial operating units shows there are a number of operating parameters which influence cetane improvement and vol- Required Temperature Increase, F Straight Run: 37.7 API & 1.1 wt.% Sulfur 15% LCO blend: 34. API & 1.1 wt.% Sulfur 3% LCO blend: 3.7 API &1.8 wt.% Sulfur Required Temperature Increase, C ume swell in a diesel hydrotreater, most notably hydrogen partial pressure, LHSV and feed API gravity (i.e. amount of LCO). Generally speaking, as LHSV decreases, the potential cetane improvement and volume swell increase. Commercial ULSD experience shows that for LHSV s around 1 hr -1 or less, cetane increases (as measured by cetane index, ASTM D-976) of about 1 numbers are achievable, provided the H 2 pressure is high enough when processing LCO blends. At higher LHSV s (greater than about 1.7 hr -1 ) the potential cetane improvement decreases to about 4 numbers or less Product Sulfur, wppm Straight Run 15% FCC LCO 3% FCC LCO Figure 8: Impact of LCO Content on Hydrotreater Performance Not surprisingly, higher pressure units tend to achieve much larger cetane increases. It has been observed that the cetane uplift is typically less than 6 numbers when the unit pressure is less than 1 Psig (69 Barg) while the cetane uplift increases to 8-1 numbers as pressure increases beyond 1 Psig (69 Barg). A more detailed discussion can be found in Reference 2. Required Temperature Increase, F Straight Run: D2887 endpoint 747 F (297 C) Low EP LCO blend: D2887 endpoint 755 F (42 C) High EP LCO blend: D2887 endpoint 774 F (412 C) Product Sulfur, wppm Straight Run High EP LCO Low EP LCO Figure 9: LCO Endpoint Effect on Catalyst Activity Increase in Cetane Index or API Gravity Percent LCO in Feed Cetane Index Increase API Increase Figure 1: Cetane Increase Observed in a Commercial ULSD Unit Required Temperature Increase, C Figure 1 is a summary of commercial data from a ULSD unit using a SmART Catalyst System which operates at about 13 Psig (9 Barg) and slightly under 1 LHSV. The feed varies from 1% straight run to 7%+ FCC LCO. As the figure shows, the percent LCO in the feed has a big effect on the cetane upgrade. At high LCO levels the cetane index increase achieved in this unit approaches 9 numbers, compared to less than 7 numbers for lower LCO levels. As might be expected, there is a significant cost in hydrogen to achieving very high cetane increases from feeds containing LCO. Figure 11 shows how the H 2 consumption increases with increasing cetane uplift from the same commercial ULSD unit. For cetane number increases of 6-7 numbers the H 2 consumption is just over 85 SCFB (143 Nm 3 /m 3 ). Notice, however, that achieving increases greater than about 8 numbers come at a very large increase in H 2 consumption; at 9 numbers of cetane improvement the H 2 consumption is about 115 SCFB (194 Nm 3 /m 3 ), an increase of about 3% compared to a 6-7 number increase. Note the similar behavior for API increase as a function of LCO content and H 2 consumption. Another product quality issue when processing LCO containing feeds is the diesel color. It is generally accepted that the species responsible for color formation in distillates are polynuclear aromatic (PNA) molecules. Some of these PNA s are green/blue and fluorescent in color, which is apparent even at very low concentrations. Certain nitrogen (and other polar) compounds have also been implicated as problems for distillate product color and product instability. Work conducted by Ma et.al. 3 concluded that the specific species responsible for color degradation in diesel are anthracene, fluoranthene and their alkylated derivatives. Other work completed by Takatsuka et.al. 4 demonstrated that the color bodies responsible for diesel product color degradation were concentrated in the higher ADVANCED REFINING TECHNOLOGIES CATALAGRAM 7

10 boiling points in the diesel (>48 F or >249 C) suggesting that color can be improved by adjusting the diesel endpoint. PNA s such as these are readily saturated to one and two ringed aromatics under typical diesel hydrotreating conditions at start of run (SOR), but as the temperature of the reactor increases towards EOR, an equilibrium constraint is reached whereby the reverse dehydrogenation reaction becomes more and more favorable. At some combination of low hydrogen partial pressure and high temperature, PNA s actually begin to form (or reform) resulting in a degradation of the diesel product color. Figure 12 summarizes data from a commercial ULSD unit using ART catalysts. The data show that the product color exceeded 2.5 ASTM, the pipeline color specification for diesel, at reactor outlet temperatures above 73 F (388 C). The feed to this unit contained 3-5% LCO and it was operated at 1. LHSV and 85 Psig (59 Barg) inlet pressure. Increase in Cetane Index or API Gravity H 2 Consumption, Nm 3 /m H 2 Consumption, SCFB Cetane Index Increase API Increase Figure 11: H 2 Consumption Increases with Cetane Uplift Figure 13 summarizes some pilot plant data which was generated as part of a larger color study using spent CDXi, a premium CoMo catalyst for ULSD 5. The figure shows a comparison of the diesel product color achieved from a SR feed and a 3% FCC LCO blend at constant H 2 /Oil ratio and two pressures. Processing the SR feed results in acceptable color over a wide range of temperatures and for both pressures shown. The product from the LCO blend, on the other hand, goes off color (>2.5 ASTM) above 73 F (388 C) at 8 Psig (55 Barg) while at 12 Psig (83 Barg) the temperature can exceed 76 F (44 C) before going off color. This clearly demonstrates the significant impact that H 2 partial pressure has on diesel product color when processing LCO containing feeds. Processing LCO as part of the feed to a ULSD unit can be challenging since the quality, quantity and endpoint of LCO affect catalyst activity and product properties. These challenges can be overcome with proper choice of catalyst system and an understanding of the impact LCO has on both unit performance and ULSD product quality. Advanced Refining Technologies is well positioned to provide assistance on how best to maximize unit performance and to take advantage of opportunities to successfully process more LCO into ULSD. References 1. B. Watkins and C.Olsen, 29 NPRA Annual Meeting, paper AM G.Rosinski and C.Olsen, Catalagram 16, Fall X. Ma et. al., Energy and Fuels, 1, pp (1996) 4. T.Takatsuka et.al., 1991 NPRA Annual Meeting, Paper AM G.Rosinski, B.Watkins and C.Olsen, Catalagram 15, Spring 29 Product Color, ASTM Outlet Temperature, C Outlet Temperature, F Figure 12: ULSD Product Color From a Commercial ULSD Unit Product Color, ASTM Straight Run 3% FCC LCO Temperature, C Temperature, F Low Pressure High Pressure High Pressure Low Pressure Figure 13: Comparison of Product Color for SR and 3% LCO 8 SPECIAL EDITION ISSUE No. 11 / 211

11 StART Catalyst System Success Story Geri D Angelo Senior Technical Service Engineer Advanced Refining Technologies Chicago, IL, USA Hydrotreating coker naphtha poses some unique challenges for the refiner. All coker naphthas contain a high concentration of olefins/diolefins, and naphthas from delayed cokers also contain silicon. In addition, these stocks have a significantly higher concentration of sulfur and nitrogen than straight run naphthas. Even though the typical processing scheme involves blending the coker stock with straight run naphtha, issues of high exotherms, pressure drop increase and silicon poisoning often control the cycle length. Advanced Refining Technologies (ART) has long had a strong position in naphtha hydrotreating technology. AT535 is one of the keys to the technology with an outstanding commercial track record with over 3 users worldwide in both straight run and coker naphtha service. In addition to the active catalyst, ART supplies active silicon guard materials to economically supplement the silicon capacity and provide activity grading. Activity grading is an important aspect of coker naphtha processing. The high heat release resulting from olefin saturation can cause polymerization and a subsequent pressure drop problem. By grading hydrogenation activity from low to high (active guard to catalyst) the temperature rise is spread out over a larger portion of the catalyst bed and the potential for polymerization is mitigated. As previously mentioned, units that process coker naphtha streams typically contain silicon which comes from the delayed coker. In order to suppress foaming at the coker, an anti-foam agent such as silicone oil (i.e. polydimethylsiloxane) is used. The silicone oil breaks down during the coking process to lower molecular weight fragments consisting of modified silica gels. These fragments end up primarily in the naphtha range fraction, although small quantities are also found in the kerosene and diesel fractions. Silicon can also be found in synthetic crudes because the process of making synthetic crude often involves a coking step. Many crude suppliers have also used additives containing silicon in drilling processes, and ADVANCED REFINING TECHNOLOGIES CATALAGRAM 9

12 pipeline companies are using silicon containing additives injected into the crudes for both flow enhancing performance and foaming issues. As a result, even refiners who don t have coking capabilities may run into Si contamination issues. ART introduced the StART TM (Silicon tolerance by ART) Catalyst System in 22 as a way for refiners to extend the cycle length of their naphtha treaters, while at the same time facing increased incidence of Si contamination. This technology combines a state of the art silicon guard material, AT724G, along with the active HDS and HDN catalyst AT535. The StART technology has been widely accepted in the industry for the combination of high Si capacity and high activity in naphtha service. Figure 14 compares the Si capacity of AT724G and AT535 with some competitor catalysts. AT535, by itself, has essentially the same Si capacity as the competitor catalysts while AT724G has over 3% higher Si capacity. Wt. % of Fresh Catalysts Si SiO 2 Competitor A Competitor C AT724G AT535 Figure 14: Comparison of Si Capacity Figure 15 shows a commercial example of the how the StART system can increase the cycle length over competitive silicon tolerant catalysts. In this case, the StART system more than doubled the cycle length over several previous cycles that used catalyst from competitor A. This refiner was processing coker naphtha and living with very short cycles of between 7 and 13 months. After installing the StART System the cycle length more than doubled. Based on this significantly improved performance this refiner continues to use a StART system in their coker naphtha unit. Another recent success story involves a naphtha hydrotreating (NHT) unit in the Gulf Coast region which had been experiencing cycle lengths of between months due to silicon poisoning of the catalyst. ART worked with the refiner to design a catalyst system which would extend the cycle length to 24 months or more. The NHT unit consists of two reactors in series and produces the feed for a downstream reformer. The feed sulfur and nitrogen get as high as 25 ppm and 14 ppm respectively, and the product specifications for sulfur and nitrogen are <.5 wppm. The unit processes a blend of straight run and coker naphthas, and typically contains between 2 and 8 wppm of silicon. The feed was not regularly analyzed for Si content, so this was estimated from the antifoam addition at the coker. Figure 16 summarizes the performance of the unit after installing the StART catalysts. Figure 17 shows how the reaction exotherms moved through the bed as the catalyst was poisoned with Si during the cycle. There were thermocouples installed throughout the lead reactor catalyst bed from top to bottom. At start of run (SOR), most of the reaction was occurring at the top of the bed as indicated in the figure. Over 5% of the delta temperature occured between the TC s at the top and Mid 1. As the cycle progresses, the top of the bed was poisoned and the exotherm shifted lower in the bed. Eventually over 7% of the exotherm occured in the bottom half of the bed (Mid 3 and Bottom TC s). After about 3 months on stream, the temperature rise at the bottom of the bed increased significantly indicating that the catalyst in the top half of the bed was poisoned. This obser- Cycle Length, Days Cycle 1 Cycle 2 Cycle 3 StART System Figure 15: StART System Doubles Cycle Length WABT, F Days on Stream Figure 16: Performance of the StART System WABT, C 1 SPECIAL EDITION ISSUE No. 11 / 211

13 vation was consistent with the estimated Si levels discussed previously, and both indicated the lead reactor was approaching the catalyst Si capacity. The performance of the unit was outstanding and easily exceeded the expected 24 month run length. After 4 months on stream, the estimated Si loading on the catalyst combined with the exotherms shown in Figure 17 indicated that it was time to change out the lead reactor. The lead reactor catalyst was replaced with fresh AT724G and AT535. The NHT was re-started with the new catalyst in the lead reactor and the AT535 in the lag reactor which had already been in use for 4 months. At the time of the lead reactor catalyst change, the NHT was still producing on spec product and the deactivation rate of the catalyst was.3-.5 F/month (.2-.3 C/month). Percent Delta Temperature Rise 1% 9% 8% 7% 6% 5% 4% 3% 2% 1% % Days on Stream Top Mid 1 Mid 2 Mid 3 Bottom ART received spent catalyst samples from the lead reactor which were lab regenerated. These were analyzed for a variety of poisons including silicon. The analysis showed that the catalyst picked up significant amounts of silicon as well as arsenic. Figure 18 shows the silicon and arsenic profile through the catalyst bed in the lead reactor. The analysis shows that silicon and arsenic got well into the catalyst bed, consistent with the observation of the exotherm shifting to the lower part of the bed. The results show that AT724G picked up 2 wt% silicon at the top of the catalyst bed and around 15 wt% Si at the bottom of the AT724G bed. This indicates there was a little more silicon capacity in the AT724G, but Si was definitely breaking through to the active catalyst. The amount of silicon in the spent AT535 samples was 8 wt% near the top of the catalyst bed and < 2 wt% at the bottom of the bed. The spent catalyst data show that little if any Si and As made it all the way through the lead reactor into the lag reactor. While there may have been some capacity left to trap more of both Si and As, the right decision was made to change out and protect the lag reactor catalyst from poisoning. Figure 17: Reaction Exotherms Move Through The Catalyst Bed Wt% Si on Catalyst 25% 1.% Silicon.9% 2% AT724G Arsenic.8%.7% 15%.6%.5% 1%.4% AT535.3% 5%.2%.1% % % Fraction of Lead Reactor Catalyst Volume Wt% As on Catalyst The use of ART s StART catalyst system in this unit was a tremendous success. The cycle length was increased from months to over five years, with full unit turnaround expected in two more years. Figure 18: Silicon and Arsenic Profiles Through the Catalyst Bed The SOR temperature after the lead reactor catalyst change-out was the same as the first StART loading despite the fact that the lag reactor was not changed out. The catalyst deactivation in the current cycle continues to run between.3-.5 F/month (.2-.3 C/month) and has currently been on stream for 18 months since the changeout of the lead reactor. The unit is on track to run through the end of 213 when it is anticipated to shut down due to silicon poisoning. At the end of the current cycle the lag reactor will have been on line for 7 years. ADVANCED REFINING TECHNOLOGIES CATALAGRAM 11

14 Controlling Feedstock Contaminants in Diesel Hydrotreating Operations Improving Unit Performance Requires Strategies for Avoiding Rapid Deactivation of Hydrotreating Catalyst from Contaminants Dave Krenzke Regional Manager Hydroprocessing Technical Service Advanced Refining Technologies Richmond, CA, USA Higher feedstock volumes contaminated with catalyst poisons such as those listed in Table 1 are being processed in high complexity refining facilities. Many of these hydrocarbons are from sources recently introduced into the global crude market in significant quantities over the past five years. For example, Canadian Synbit and Dilbit crudes will come to make up a significant fraction of feedstocks to North American refineries. North American refiners and their technology partners are just now discovering the challenges encountered with upgrading these cheap crudes to ULSD specifications. In other producing regions such as deep offshore Brazil, preliminary reports indicate that hydrocarbons from the Tepi oil field, which has been called the greatest oil discovery in the past 1 years, may be relatively high in nitrogen content. What has also been noticeable are the new challenges certain refiners are facing with the redistribution of certain crudes. For example, the heavy Venezuelan crudes (e.g., Merey Crude: 16 API, 11.5 UOP K-factor, 2.7wt% S, 36 ppm N, etc.) that U.S. Gulf Coast refiners began processing back in the 198s are finding a new market in emerging refining regions such as India. These new highcomplexity facilities may nonetheless be unfamiliar with their contaminants. 12 SPECIAL EDITION ISSUE No. 11 / 211

15 Contaminant Feed Guidlines Common Source Remedies Si < 1. wppm Anti foam from delayed cokers Guard catalysts Na, Ca <.5 wppm Sea water; caustic Improved desalting; Guard catalysts; Don t send spent caustic to feed tanks or units As < 25 ppb Crudes from W. Africa, Russia, synthetic crudes High Ni guard catalysts Pb, P <.5 wppm Gasoline slop tanks, imported feeds Don t process feeds containing Pb or P Ni + V < 1. wppm total for Ni + V + Fe Resid; heavy feeds Better feed distillation; Guard Catalysts; Bed grading Fe < 1. wppm total for Ni + V + Fe Soluble: corrosion; insoluble: unfiltered particulates Inert guard material with high void space; Fe-traps (soluble); Top bed skimming C7 insolubles MCR < 1 wppm <.5 wt% Resid; heavy feeds Better feed distillation; Guard catalysts; Bed grading Table 1: Feed Contaminants Commonly Found in Crudes The contaminants found in these feeds can be particularly harmful to hydrotreating catalyst activity. In anticipation of the challenges involved in processing these types of feedstocks, there has been a noticeable trend towards designing new diesel hydrotreating units at higher hydrogen partial pressures to compensate for catalyst activity loss from these poisons. Instead of operating at higher H 2 partial pressures, older units can be modified to meet Euro 4 or ULSD targets, for example, with higher catalyst volume and lower space velocity. Contaminant Control In general, silicon (Si) carryover from Si-based antifoam agents used in delayed coker operations will plug up catalyst porosity. Sodium (Na) and calcium (Ca) can originate from seawater exposure, poor desalting or caustic sources in the refinery. Arsenic (As) is becoming more of a common issue as more synthetic crudes and crudes from Africa and Russia are processed. It is present at low levels throughout the whole boiling range for synthetic crudes, which is why arsenic traps may need to be employed. Depending on the future cost of crudes, the shale oil based hydrocarbons from Colorado that first drew interest in the 198s, that are relatively low in sulfur, may again become marketable in large quantities. However, they are high in As and Ni. Also, phosphorous (P) is becoming a problem with some feeds. The presence of nickel (Ni) and vanadium (V) in heavier feeds has always been a significant concern to hydrotreating operations. Their presence may also mean entrainment is occurring. These Ni + V contaminants are found in significant concentrations in deep-cut VGO streams (e.g., 1-2 ppm), which is why end-point control is critical with heavy feedstocks. In hydrotreating operations, either one of these elements will behave like coke and plug up the catalyst. Si will also plug up catalyst pores. Although Fe in naphthenic crudes may not be as significant a problem as Ni + V poisoning of hydrotreating catalyst, it nonetheless is a corrosion precursor and leads to FeS formation. Soluble Fe generated from naphthenic or acidic-based crudes can lead to corrosive products entering catalyst beds. ADVANCED REFINING TECHNOLOGIES CATALAGRAM 13

16 Answers to the 211 NPRA Q&A FCC Questions Contaminants/Analytical 2. Are there any standard sampling and analytical methods that can be used in the refinery labs to accurately determine the silicon content in the feed to the coker naphtha hydrotreater? Geri D Angelo Accurately measuring silicon in naphtha streams can be done but it takes a bit of work to get a representative sample of the naphtha. The silicon in the coker naphtha depends on the type and amount of antifoam chemical at the delayed coker unit. Delayed cokers have cycles ranging anywhere between 8-24 hours. The coker unit is continually producing a coker naphtha stream during these cycles which is typically being sent from the fractionator straight into the naphtha hydrotreater feed drum. The antifoam chemical is usually not added for the entire coker cycle. This means that the silicon in the naphtha stream will vary with the timing of the coker cycle. In order to get a representative amount of silicon in the coker naphtha stream a composite should be made of hourly samples mixed together for the time of the cycle. For example, for an eight hour cycle, eight samples would be mixed and the composite sample analyzed for silicon. To measure the silicon an ICP-MS (Inductively Coupled Plasma Mass Spectrometry) instrument can be used. This instrument/method can measure very low silicon concentrations. Fouling/Poisons 4. What has been your experience with antimony and phosphorous poisoning on hydrotreating catalyst performance? What is the maximum level? Charles Olsen Phosphorous (P) contamination in oil has been traced to frac fluids that are often used in crudes from the Western Canadian Sedimentary Basin. The source is diphosphate esters which are soluble in the crude oil. Refineries that run large percentages of light Western Canadian crude have reported crude column and crude furnace fouling for many years. Improvements made to crude columns to minimize fouling have transitioned the depositing of phosphorous to the downstream hydrotreaters. Other sources of phosphorous include gasoline slop tanks, imported feeds and lube oil wastes. If phosphorous does manage to make its way into the hydrotreater it will poison the active sites of the catalyst causing a loss in activity. A level of 1 wt% of phosphorous on the catalyst results in roughly 1 F (6 C) loss in activity. ART recommends that a feed content of <.5 wppm be maintained whenever possible, as well as the use of feed filters to assist in trapping of phosphorous sediment. Historically, phosphorous contamination has not been very common but, with the increasing use of opportunity crudes, it is being observed more frequently. A recent example is summarized in the table below which shows the results of some spent catalyst analysis from a diesel unit. This unit experienced extremely rapid catalyst deactivation shortly after start up. It was so severe that within several months the unit required an unplanned turnaround and fresh catalyst was installed. The spent catalyst analysis indicates the catalysts were exposed to high levels of several poisons including sodium and phosphorous. The contaminants penetrated well into the catalyst bed. The level of contaminants indicates that the catalyst in the top half of the bed had lost over 6 F (33 C) of activity. Sulfiding 6. What is the minimum hydrogen sulfide required in the recycle gas for units with low sulfur feed? Do refiners inject sulfur compounds to maintain a minimum concentration? Gordon Chu There is no minimum hydrogen sulfide requirement as long as the feed contains some sulfur as the sulfided catalyst is very resistant to sulfur loss under normal process conditions. We are not aware of any refiners adding sulfur compounds to maintain a minimum H 2 S concentration during the process cycle. Na, wt% P, wt% GSK GSK-6A Top Catalyst Table 2: Commercial Example of Phosphorus Poisoning 14 SPECIAL EDITION ISSUE No. 11 / 211

17 Start-Up and Shut-Down 7. Is there any harm adding cracked stocks too quickly after break-in following catalyst activation? What is typical introduction rate? Kerosene/LCO Processing 1. In treating kerosene, what factors play into the decision to use hydrotreating versus sweetening processes such as caustic treating? Ben Sim Dave Krenzke Introducing cracked stocks too early after sulfiding will cause noticeable loss in activity. Coke precursor molecules in cracked feeds will have a tendency to form coke over the fresh and highly active sites on the catalyst. Delaying the introduction of cracked stocks for at least 3 days after sulfiding will allow the catalyst activity to be passivated which helps to minimize these effects. After running for three days on straight run the cracked material should be added to the feed stream gradually. ART typically recommends adding the cracked feed in small increments every shift making sure the reactor exotherm remains under control and within acceptable limits before increasing the cracked feed amount any further. The decision to use hydrotreating or a sweetening process depends on the types of sulfur in the kerosene and the product sulfur target. Hydrotreating can remove all types of sulfur compounds and therefore the sulfur content of the product is only limited by the process conditions and catalyst activity. The sweetening process only removes mercaptan sulfur so the product sulfur is limited to the nonmercaptan sulfur in the feed. 11. What LCO 95% distillation point do you target for optimizing ULSD production? Do you see a significant catalyst life penalty with increased LCO cut point? Brian Watkins Loading and Unloading 9. Have you successfully dumped, screened and reloaded spent hydrotreating or hydrocracking catalyst without regeneration during a turnaround? Can you share any best practices during this operation to avoid problems on restart? Greg Rosinski Spent hydroprocessing catalyst is pyrophoric due to small particulates of iron sulfide scale that are present, so care must be taken to minimize the exposure of the spent catalyst to air. In addition, spent sulfided catalyst has some coke on it and it will slowly oxidize in air. If the spent catalyst is exposed to air, it will slowly heat up and, if iron sulfide is present, it will combust which may ignite the coke or other residual hydrocarbon on the catalyst. The key to this procedure is to have competent and experienced personnel performing the required tasks. The reactor must be thoroughly swept of hydrocarbons, and a nitrogen purge should be kept on the reactor at all times. During the unloading, the screener and the dump nozzle should be continuously purged. The containers that will hold the catalyst during unloading should be blanketed with nitrogen or have dry ice placed inside until ready for loading. The containers should not be open to the atmosphere. The loading should be done under inert conditions with experienced personnel. When preparing your procedure, make sure to involve your refinery EH&S group and give careful consideration to all aspects of the process to ensure that you take all the precautions necessary. The addition of LCO to a ULSD hydrotreater has several effects such as increased hydrogen consumption, higher required reactor temperatures and possibly shorter cycle time. Figure 19 summarizes some of pilot plant data comparing a SR and a SR/LCO feed blend. It shows that the SR diesel requires a 43 F (24 C) increase in temperature to go from 1 ppm sulfur down to 1 ppm sulfur. The 2% LCO blend requires almost 2 F (11 C) higher temperature to achieve the same product sulfur relative to the SR feed. The product from the LCO blend also has a 2 to 3 number lower API compared to the SR product, and hydrogen consumption increases significantly for the LCO blend due to saturation of additional polyaromatic compounds found in the LCO. These latter consequences set limits on the amount of LCO which can be processed and still meet product cetane specifications and also meet hydrogen availability constraints. One option to re-gain some of the lost activity in adding additional LCO is to change the end point of the LCO in the feed. ART completed pilot plant testing on an LCO from the same FCC which had been cut at two different end points. Table 3 lists the analysis of the two LCO feeds and shows that the end points differed by about 4 F (22 C). The decrease in endpoint lowers the total sulfur by almost 1 ppm and total nitrogen decreases by 129 ppm. A comparison of activity on the two LCO feeds blended at 3% by volume into SR feed is shown in Figure 2. Over 3 F (17 C) higher temperature is required to treat the higher endpoint feed to meet the ULSD specification. This difference in activity corresponds to a significant decrease in the hydrotreater cycle length. ADVANCED REFINING TECHNOLOGIES CATALAGRAM 15

18 The addition of LCO has a major impact on activity for both the low and high endpoint LCO materials. The required temperature increase for ULSD in going from to 3% LCO for the lower endpoint material is about 1.2 F (.7 C) per percent LCO. Processing the higher endpoint LCO increases the required temperature to about 1.4 F (.8 C) per percent LCO. Figure 21 demonstrates this more clearly in the form of a plot of the required temperature increase as a function of LCO content. Notice from the chart that the activity effects are not exactly linear with increasing LCO content. The first 15% LCO has a larger impact on activity than the next 15%. Required Temperature Increase, F Required Temperature Increase, C Volume Gain/Conversion Product Sulfur, ppm SR LCO 12. What sets the volume gain in ULSD units? How much does lowering the space velocity increase the volume gain? How much volume gain can be expected for each feed component? Figure 19: Activity Comparison on SR and Blended SR/LCO Charles Olsen There are a number of parameters which influence volume gain in a ULSD unit. Hydrogen partial pressure and LHSV are two key operating conditions which have a large effect on the product volume increase. Catalyst selection also plays an important role since at higher pressures NiMo catalysts have a higher aromatics saturation activity compared to CoMo catalysts. Figure 22 shows the total volume yield on a fresh feed basis that has been achieved in several commercial diesel hydrotreaters as a function of unit LHSV. Generally speaking, as LHSV decreases the potential volume swell increases. At a LHSV around 1 hr -1 or less, total product volume increases of 6-7% or more are achievable (provided the H 2 pressure is high enough), while at a LHSV greater than about 1.7 hr -1 the total product volume increase is about 1-2%. Type LCO (low FBP) LCO (high FBP) Required Temperature Increase, F SR Product Sulfur, ppm 3% Hi FBP LCO 3% Lo FBP LCO Figure 2: Impact of Endpoint Reduction on Hydrotreating Performance Required Temperature Increase, C API Sulfur, wt% Nitrogen, ppm Aromatics, lv% mono-,lv% poly-, lv% Dist., D2887, F/C IBP 249/ /124 1% 425/ /222 5% 531/277 55/288 7% 6/316 62/327 9% 677/ /371 FBP 772/ /433 WABT Increase to Achieve Product Sulfur, F Hi EP LCO % LCO Lo EP LCO WABT Increase to Achieve Product Sulfur, C Table 3: Comparison of Boiling Point Reduction on LCO Figure 21: Activity Comparisons at different LCO FBP and Concentration 16 SPECIAL EDITION ISSUE No. 11 / 211

19 Total Product Volume, % of Fresh Feed LHSV, hr1 Refiner A Refiner B Refiner C Refiner D Refiner E Refiner F Of course LHSV is not the only parameter which can influence the volume swell. H 2 partial pressure also has a significant effect. Figure 23 summarizes the total product volume yield as a function of unit pressure for the commercial units shown in Figure 22. Not surprisingly, higher pressure units tend to achieve a much higher level of volume swell. In these examples, the volume increase is typically less than 3% when the unit pressure is less than 1 Psig (69 Barg). The total volume swell increases to 4-7% as the unit pressure increases beyond 1 Psig (69 Barg). The data in Figures 22 and 23 also suggest there is a practical limit to the volume swell achieved from typical hydrotreating. A comparison of the volume swell achieved by Refiners A and B shows they are roughly the same for both units despite the large difference in operating pressure at similar LHSV. Figure 22: Effect of LHSV on Volume Swell in Commercial Units Total Product Volume, % of Fresh Feed Refiner A Refiner B Refiner C Refiner D Refiner E Refiner F Reactor Inlet Pressure, Barg Reactor Inlet Pressure, Psig The volume swell also varies significantly with feedstock. Figure 24 summarizes how the total product volume yield correlates with the gravity of the feed. In general, the product volume swell increases as the feed API decreases or specific gravity increases. In other words, as more FCC LCO is added to the feed the potential volume swell from hydrotreating increases. As mentioned previously, the catalyst will also have an impact on the degree of volume swell achieved in a hydrotreater. It is well known that NiMo catalysts have higher aromatic saturation activity than CoMo catalysts, and therefore a NiMo catalyst is expected to deliver greater volume swell. Figure 25 summarizes pilot plant data which demonstrates this. These data were generated using a 25% LCO containing feed, and shows that the NiMo catalyst results in 1-2 numbers higher total product volume increase compared to the CoMo catalyst. Figure 23: Effect of Unit Pressure on Volume Swell Feed Specific Gravity Total Product Volume, % of Fresh Feed Refiner A Refiner B Refiner C Refiner D Refiner E Refiner F Feed API Gravity Total Product Volume, % of Fresh Feed LHSV, hr -1 NiMo CoMo Figure 24: Feed Gravity Has a Significant Impact on Volume Swell Figure 25: The Catalyst Type Effects Volume Swell ADVANCED REFINING TECHNOLOGIES CATALAGRAM 17

20 13. What considerations are being given to include mild hydrocracking in your high pressure ULSD unit? Robert Wade Many refiners consider complimenting their existing ULSD HDT catalyst with a hydrocracking catalyst to improve cold flow properties. This is accomplished through end point reduction. The hydrocracking catalyst used is usually the type that preferentially cracks large straight chain paraffins. For properly designed catalyst systems this will be economic as heavier feeds may be processed while meeting stringent diesel cold flow properties, and simultaneously not severely reducing diesel selectivity. Some refiners find it economic to retrofit their existing ULSD units so that they can run full range VGO. This option is very dependent on the design pressure of the reactors as the fouling rate will be considerably higher for a VGO service. In addition, the capital expense required to meet recovery section requirements may be a major consideration. Hydrogen Optimization/HPNA s 14. With limited hydrogen availability for desulfurization of diesel, what criteria influence the optimization of hydrogen consumption between the FCC Pretreat and ULSD units? What catalytic options exist to achieve the desired balance of consumption? Greg Rosinski For any given feed, hydrogen consumption is a function of hydrogen partial pressure, LHSV, H 2 /Oil and catalyst. For the most part, the first three variables are fixed for a given unit, since throughput reduction is not an economical choice. Thus, catalyst selection is one of the few variables which refiners are willing to consider changing. CoMo catalysts have lower hydrogen consumption than NiMo catalysts due to lower aromatic saturation activity. At equivalent product sulfur, using all CoMo catalyst in the FCC Pretreater will lower the hydrogen consumption with a longer cycle in terms of HDS activity, but at the cost of lower conversion in the FCC and higher LCO yields. Using all NiMo catalyst in the FCC Pretreater will result in higher FCC gasoline yields and lower LCO yields due to higher PNA saturation, but a shorter cycle life in terms of HDS activity. With regards to the ULSD units, if the unit is high pressure, using a NiMo catalyst will result in higher aromatic and PNA saturation. This may be beneficial if cetane upgrade is desired; however, there may be a diminishing return on hydrogen for the incremental cetane upgrade over a CoMo catalyst. ART can help optimize both FCC Pretreater and ULSD performance based upon the refiner s needs, including hydrogen consumption, cetane uplift, and cold flow properties. ART provides the ApART TM and SmART staged catalyst systems for FCC Pretreat and ULSD applications, respectively. ART has helped many refiners manage hydrogen consumption in both units by using staged catalyst systems utilizing NiMo, CoMo and NiCoMo catalysts optimized to enhance HDS, HDN or HDPNA activity for a given feedstock. Furthermore, ART s relationship with Grace Davison can enhance the unit optimization to include the FCC unit as well as the FCC pretreater and the ULSD unit. Utilizing the technical resources of both ART and Grace Davison, the refiner can gain a more comprehensive understanding of the interactions and dependence of these units on each other in terms of hydrogen consumption and product property enhancement. Hydroprocessing 16. Pre-hydrotreated feeds and crudes look easy to process on paper. Why is it more difficult than expected to process pre-hydrotreated feeds in a hydroprocessing unit? Brian Watkins Opportunity feedstocks, having already been processed through conventional refinery processes, pose unexpected challenges to refiners wishing to incorporate them into the distillate pool. Some of these streams have proven to be significantly more difficult to process underscoring the fact that it is important to understand the potential impacts of processing new feed streams in order to avoid unpleasant surprises. Significant differences in feed reactivity for various pre-processed feed components are not necessarily anticipated from observing the usual bulk feed analyses. When considering the use of synthetic crudes an understanding of the upstream processing is important. Production of synthetic fuels involves a combination of several processes in order to accommodate downstream processing. These upstream processes include coking or an ebullating bed resid operation, followed by a hydrotreating or hydrocracking operation in order to produce a lighter grade material. These hydrocracking units tend to operate at severe conditions in conjunction with high hydrogen partial pressures. At these conditions, the removal of all the easy, less refractory sulfur is readily achieved, and the majority of the multi-ring aromatics are saturated. This leaves a product that is relatively low in sulfur and PNA s and, when added to the feed to a ULSD unit, gives rise to a surprisingly difficult feedstock to process. These products are then blended in with other heavier materials as a diluting or cutting stock and sent downstream as synthetic crude. 18 SPECIAL EDITION ISSUE No. 11 / 211

21 Light SR Gas Oil LCO EB Diesel Synthetic Diesel FB Diesel Sulfur, wt% Nitrogen, wppm API Gravity Aromatics, vol% Total mono poly Dist., D2887, F/ C.5 286/ /14 33/ / / / /186 47/ / / /316 47/ / / / / /33 77/ / / / /34 759/ / / 413 Sulfur Distribution Thiophenes 8 Benzothiophenes (BT) 2 69 Substituted BT s DiBenzothiophenes (DBT) Substituted DBT s ,6 DM-DBT C 3 -DBT Table 4: Diesel Feedstock Analysis Likewise, the use of diesel range products from an H-Oil, LC-FIN- ING unit or fixed bed residuum desulfurizer can also have a significant impact on downstream diesel catalyst activity for similar reasons. The general properties of these types of diesel feeds often indicate that they may be relatively easy to hydrotreat due to their low sulfur content and higher API gravity, which is often similar in appearance to straight run (SR) materials. Table 4 lists the properties for several of these diesel feeds including the diesel product fractions from an ebullating bed residuum (EB) unit, a fixed bed residuum (FB) unit and a diesel fraction from Canadian synthetic crude. ADVANCED REFINING TECHNOLOGIES CATALAGRAM 19

22 m Figure 26 shows the activity difference between a SR and a blended SR/Synthetic diesel. Note that at higher product sulfur, the two feed- 9 5 stocks respond similarly to each other. As the application becomes more demanding, the required reactor temperature increases dramatically for the synthetic diesel feed compared to the SR feed. The blended feed requires more than 25 F (14 C) higher temperature relative to the SR to achieve ULSD sulfur levels. It is reasonable to expect that the upstream hydroprocessing of the synthetic diesel material results in a feed that behaves similarly to other previously hydrotreated feedstocks like those from the EB and FB residuum applications. Required Temperature Increase, F Required Temperature Increase, C Advanced Refining Technologies can work closely with the refining technical staff to help plan for processing opportunity feeds such as those discussed above. One of the keys is being aware of the potential impacts processing certain feeds will have on unit performance. Feeds which have been previously processed present unique challenges and ART is well positioned with its experience at providing customized catalyst systems for ULSD applications. Opportunity feeds provide yet another objective to consider when designing the appropriate catalyst system to maximize unit Product Sulfur, ppm SR Synthetic Figure 26: Activity Comparison of the SR and Synthetic Diesel Blend 5 ART featured in August issue of Hydrocarbon Engineering The August 211 issue featured ART on its cover and in the lead article Yield of Dreams by ART s Brian Watkins and Chuck Olsen and Grace s David Hunt. The collaboration discusses the key relationships between FCC pretreat and FCC unit operations and their corresponding catalyst systems in maximizing the distillate pool. Both processes must be reoptimized as the refiner moves from gasoline to distillate production to ensure maximum profitability. The article emphasizes the need for refiners to follow an integrated approach to managing the catalysts and operation of the FCC pretreater and FCCU. The complex relationship between the FCC pretreater and the FCCU underscores the importance of working with a catalyst technology supplier that has the capability to understand the interplay between the hydrotreating performance of the FCC pretreater and the performance, yield structure and product sulfur distributions of the FCCU. For copies of the article, contact betsy.mettee@grace.com HYDROCARBONENGINEERING ON AUGUST211 Volume 16 Number 8 - August 211 The World s Leading Supplier of Hydroprocessing Catalysts Distillate Hydrotreating Fixed Bed Resid Hydrotreating Ebullating Bed Resid Hydrocracking Let ART be a part of your solution. For more informationon on our products, please contact us. Advanced Refining Technologies 75 Grace Drive Columbia, MD 2144 USA V F W artcatalysts.com 2 SPECIAL EDITION ISSUE No. 11 / 211

23 - < ' ART's 42DX Catalyst Leading the Pack in ULSD Processing Confirmed by independent lab testing According to independent lab testing*, ART's 42DX ranks highest in activity among six leading ULSD catalysts. 42DX is the newest CoMo member of ART's DX Catalyst Platform. Depending on operating conditions and objectives, 42DX can be applied in either standalone service or stacked with ART's NDXi, the high-performance NiMo analog, to achieve the additional benefits of ART's SmART Catalyst System technology. Our technical experts can discuss your specific ULSD operation and tailor a SmART Catalyst System to meet your deep treating requirements. Contact us to learn more about how to extend your cycle, improve upset recovery and optimize H 2 consumption in your ULSD unit with ART catalyst technology. Advanced Refining Technologies, LLC 75 Grace Drive Columbia, MD 2144 USA v F W artcatalysts.com *Independent lab results available upon request. Advanced Refining Technologies