Catalagram 110 - Grace

No. 110 / Fall 2011 / www.grace.com
Catalagram
®
A Refining Technologies Publication
Hydroprocessing Catalysts from
The Chevron &GraceJoint Venture

Happy 10th 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 10th anniversary of the formation of Advanced Refining Technologies LLC,
the hydroprocessing joint venture between Grace and Chevron Products Company. The venture was created
in 2001 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

2 Hydrocracking
6 The
9 StART
12 Controlling
14 2011
Pretreat Catalyst
Development
By Jifei Jia and Dave Krenzke,
Advanced Refining Technologies
and Theo Maesen, Chevron Lummus
Global
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
Feedstock Contaminants
in Diesel
Hydrotreating Operations
By Dave Krenzke, Advanced Refining
Technologies
NPRA Q&A Answers
No. 110 / Fall 2011 / www.grace.com
®
Catalagram
A Refining Technologies Publication
Catalagram 110
ISSUE 2011
Managing Editor:
Scott K. Purnell
Technical Editor:
Charles Olsen
Contributors:
Geri D’Angelo
Jifei Jia
Dave Krenzke
Charles Olsen
Greg Rosinski
Brian Watkins
Guest Contributors:
Theo Maesen
Please address your comments to:
betsy.mettee@grace.com
Advanced Refining
Technologies
7500 Grace Drive
Columbia, MD 21044
410.531.4000
Hydroprocessing
Catalysts Caa
talysts
fr from om
The Chevron
& Grace
J JJoint
oint
Venture
© 2011 Advanced Refining
Technologies, LLC
10.7.11

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
2 SPECIAL EDITION ISSUE No. 110 / 2011
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 10 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 2001, 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.

Catalyst
Support
Development
Optimizing the catalyst support
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.
The physical properties of
the optimized support deter-
mine 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
590DX 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
Relative Activity
140
130
120
110
100
90
80
70
60
Small pores result higher surface area giving a higher
Large pores result in lower surface area, but offer better
Decreasing access to active sites
Increasing access to active sites
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 MoS2 slabs
before the Ni ions are sulfided.
Once Ni is released by
the chelate during sulfidation
it moves to the edges of the
MoS2 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
+10
Base
Catalyst Without
Surface Modification
HDN HDS
Catalyst With
Surface Modification
ADVANCED REFINING TECHNOLOGIES CATALAGRAM ® 3
120
110
100
90
80
70
60
50
Relative Activity
+6
Base
FIGURE 2: VGO HDN/HDS Performance of Catalysts
With and Without Surface Modification
HDN/HDS Activity Improvement, ˚F
+10
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

ART 590DX Commercial
Performance
590DX was commercialized in the middle of 2009. By mid 2010 it
had already been selected by seven refineries for catalyst refill. Figure
4 shows the HDN performance comparison between 590DX and
the previous generation HCR pretreat, NDXi, in one of the Chevron
refineries. 590DX is 20ºF (11ºC) more active than NDXi after 150
days on stream.
In one of the early uses of 590DX, 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 590DX. Despite the
compromised start-up, the 590DX catalyst continued to activate, and
became as active as the NDXi catalyst in the prior run at 2400
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 refrac-
tory than in the current run, it is clear that 590DX is a remarkably ro-
bust catalyst. The commercial results are summarized in Figure 5.
As a point of comparison, Figure 6 shows a different commercial op-
eration for 2 cycles using the same catalyst grade. One cycle
started-up normally while the other cycle had a significant upset dur-
ing sulfiding. This data demonstrates a more typical result for a sig-
nificant 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 deacti-
vation rate. This highlights the robust response of 590DX to opera-
tional upsets even during the critical start-up phase.
Summary
The formation of the ART joint venture in 2001 has resulted in a sig-
nificantly increased rate of innovation for HCR pretreat catalysts as
can be seen in Figure 7.
ART’s NDXi, 590DX 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 40
times for catalyst refills and start-ups by many refineries since its
commercialization in 2006. 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 AT580. The base metal loading and alumina support is
similar for AT580 and NDXi which confirms the advantage of
chelate technology to more efficiently utilize the active metals.
4 SPECIAL EDITION ISSUE No. 110 / 2011
WABT, ˚F
+80
+60
+40
+20
Base
-20
-40
NDXi
0 50 100 150 200 250
Days on Stream
590DX
300
+44
+33
+22
Base
FIGURE 4: HDN Performance Comparison Between
NDXi and 590DX
WABT, ˚F
+80
+60
+40
+20
Base
-20
-40
590DX
0 2000 6000 8000 10000 12000
Total Feed MBBLS
NDXi
+11
-11
-22
+44
+33
+22
+11
Base
FIGURE 5: Commercial HDN Performance Comparison
Between NDXi After Smooth Start-up and 590DX
After Compromised Start-up
-11
-22
WABT, ˚C
WABT, ˚C

WABT, ˚F
+100
+80
+60
+40
+20
Base
Base
0 100 200 300 400 500
Days on Stream
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
ART is now commercializing 591DX, a wider pore version of 590DX,
to handle VGO feeds with higher endpoints. The modified support
properties will provide equivalent activity to 590DX but enhanced
stability with heavier feeds.
References
1. Krenzke, Dave; Vislocky, Jim. Hydrocarbon Engineering
(2007), 12(11), 57-58
+55
+44
+33
+22
+11
WABT, ˚C
HDN Activity Improvement, ˚F
+40
+30
+20
+10
Base
AT580
NDXi
591DX
590DX
1972 1988 1996 2003 2006 2009 2011
FIGURE 7: Hydrocracker Pretreat Catalyst
Generations
ADVANCED REFINING TECHNOLOGIES CATALAGRAM ® 5
+22
+17
+11
+6
Base
HDN Activity Improvement, ˚C

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
6 SPECIAL EDITION ISSUE No. 110 / 2011
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 unit1 .
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 30% LCO. It is clear that even
low levels of LCO impact catalyst activity. At higher product sulfur about 20°F (11°C) higher temperature is
required for 15% LCO and this increases to 40°F (22°C) higher temperature for 30% 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 30% LCO blends relative to the straight run feed. The endpoints
for the 2 LCO’s differ by just over 40°F (22°C). This difference corresponds to an additional 800 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 20°F (11°C) higher temperature relative to the lower endpoint feed at higher product sulfur, and nearly
30°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.

Required Temperature Increase, ˚F
120
100
80
60
40
20
Straight Run: 37.7 API & 1.10 wt.% Sulfur
15% LCO blend: 34.0 API & 1.10 wt.% Sulfur
30% LCO blend: 30.7 API &1.08 wt.% Sulfur
0
0
0 50 100 150 200 250 300 350 400 450
Product Sulfur, wppm
Straight Run 15% FCC LCO 30% FCC LCO
Figure 8: Impact of LCO Content on Hydrotreater
Performance
Required Temperature Increase, ˚F
120
100
80
60
40
20
Straight Run: D2887 endpoint 747˚F (297˚C)
Low EP LCO blend: D2887 endpoint 755˚F (402˚C)
High EP LCO blend: D2887 endpoint 774˚F (412˚C)
0
0
0 50 100 150 200 250 300 350 400 450
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
9.5
9.0
8.0
7.5
7.0
6.5
6.0
5.5
5.0
10.0 20.0 30.0 40.0 50.0 60.0 70.0 80.0
Percent LCO in Feed
Cetane Index Increase API Increase
Figure 10: Cetane Increase Observed in a Commercial
ULSD Unit
67
56
44
33
22
11
67
56
44
33
22
11
Required Temperature Increase, ˚C
Required Temperature Increase, ˚C
A survey of commercial operating units shows there are a number of
operating parameters which influence cetane improvement and vol-
ume swell in a diesel hydrotreater, most notably hydrogen partial
pressure, LHSV and feed API gravity (i.e. amount of LCO). Gener-
ally speaking, as LHSV decreases, the potential cetane improve-
ment 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 10 numbers are
achievable, provided the H 2 pressure is high enough when process-
ing LCO blends. At higher LHSV’s (greater than about 1.7 hr -1 ) the
potential cetane improvement decreases to about 4 numbers or less.
Not surprisingly, higher pressure units tend to achieve much larger
cetane increases. It has been observed that the cetane uplift is typi-
cally less than 6 numbers when the unit pressure is less than 1000
Psig (69 Barg) while the cetane uplift increases to 8-10 numbers as
pressure increases beyond 1000 Psig (69 Barg). A more detailed
discussion can be found in Reference 2.
Figure 10 is a summary of commercial data from a ULSD unit using
a SmART Catalyst System ® which operates at about 1300 Psig (90
Barg) and slightly under 1 LHSV. The feed varies from 100%
straight run to 70%+ 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 ap-
proaches 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
850 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 consump-
tion is about 1150 SCFB (194 Nm 3 /m 3 ), an increase of about 30%
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 re-
sponsible for color formation in distillates are polynuclear aromatic
(PNA) molecules. Some of these PNA’s are green/blue and fluores-
cent in color, which is apparent even at very low concentrations.
Certain nitrogen (and other polar) compounds have also been impli-
cated as problems for distillate product color and product instability.
Work conducted by Ma et.al. 3 concluded that the specific species re-
sponsible for color degradation in diesel are anthracene, fluoran-
thene 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

oiling points in the diesel (>480°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 de-
hydrogenation reaction becomes more and more favorable. At some
combination of ‘low’ hydrogen partial pressure and ‘high’ tempera-
ture, PNA’s actually begin to form (or reform) resulting in a degrada-
tion 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 730°F (388°C). The feed to this unit contained
30-50% LCO and it was operated at 1.0 LHSV and 850 Psig (59
Barg) inlet pressure.
Figure 13 summarizes some pilot plant data which was generated as
part of a larger color study using spent CDXi, a premium CoMo cata-
lyst for ULSD 5 . The figure shows a comparison of the diesel product
color achieved from a SR feed and a 30% FCC LCO blend at con-
stant 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 730°F (388°C) at 800 Psig
(55 Barg) while at 1200 Psig (83 Barg) the temperature can exceed
760°F (404°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 challeng-
ing 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, 2009 NPRA Annual Meeting, paper
AM-09-78
2. G.Rosinski and C.Olsen, Catalagram ® 106, Fall 2009
3. X. Ma et. al., Energy and Fuels, 10, pp 91-96 (1996)
4. T.Takatsuka et.al., 1991 NPRA Annual Meeting, Paper AM-91-
39
5. G.Rosinski, B.Watkins and C.Olsen, Catalagram ® 105, Spring
2009
8 SPECIAL EDITION ISSUE No. 110 / 2011
Increase in Cetane Index or API Gravity
H2 Consumption, Nm
135
9.5
145 155 165 175 185 195
3 /m3 9.0
8.5
8.0
7.5
7.0
6.5
6.0
5.5
5.0
800 850 900 950 1000 1050 1100 1150 1200
H 2 Consumption, SCFB
Cetane Index Increase API Increase
Figure 11: H 2 Consumption Increases with Cetane
Uplift
Product Color, ASTM
304 316 327 338 349 360 371 382 393 404 416
3.5
3.0
2.5
2.0
1.5
1.0
0.5
Outlet Temperature, ˚C
0.0
580 600 620 640 660 680 700 720 740 760 780
Outlet Temperature, ˚F
Figure 12: ULSD Product Color From a Commercial
ULSD Unit
Product Color, ASTM
338 349 360 371 382 393 404 416 427
4.0
3.5
3.0
2.5
2.0
1.5
1.0
0.5
Straight Run
30% FCC LCO
Temperature, ˚C
0.0
640 660 680 700 720 740 760 780 800
Temperature, ˚F
Low Pressure
High Pressure
High Pressure
Low
Pressure
Figure 13: Comparison of Product Color for SR and
30% LCO

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, is-
sues 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 300 users
worldwide in both straight run and coker naphtha service. In addition to the active catalyst, ART supplies ac-
tive 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

pipeline companies are using silicon containing additives injected
into the crudes for both flow enhancing performance and foaming is-
sues. 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 2002 as a way for refiners to extend the cycle length of
their naphtha treaters, while at the same time facing increased inci-
dence 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 ac-
cepted 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 30% higher Si capacity.
Figure 15 shows a commercial example of the how the StART sys-
tem 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 12-18 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 2500 ppm and 140 ppm respectively, and the product speci-
fications for sulfur and nitrogen are

vation was consistent with the estimated Si levels discussed previ-
ously, and both indicated the lead reactor was approaching the cata-
lyst Si capacity.
The performance of the unit was outstanding and easily exceeded
the expected 24 month run length. After 40 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 40 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 0.3-0.5°F/month (0.2-0.3°C/month).
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 re-
actor. 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 20 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 lit-
tle 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.
The use of ART’s StART catalyst system in this unit was a tremen-
dous success. The cycle length was increased from 12-18 months
to over five years, with full unit turnaround expected in two more
years.
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 0.3-0.5°F/month (0.2-0.3°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
2013 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.
Percent Delta Temperature Rise
100%
90%
80%
70%
60%
50%
40%
30%
20%
10%
0%
Top Mid 1
Days on Stream
Mid 2 Mid 3 Bottom
Figure 17: Reaction Exotherms Move Through The
Catalyst Bed
Wt% Si on Catalyst
25%
1.0%
Silicon 0.9%
20%
AT724G
Arsenic 0.8%
0.7%
15%
0.6%
0.5%
10%
0.4%
AT535
0.3%
5%
0.2%
0.1%
0%
0%
0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0
Fraction of Lead Reactor Catalyst Volume
Figure 18: Silicon and Arsenic Profiles Through the
Catalyst Bed
ADVANCED REFINING TECHNOLOGIES CATALAGRAM ® 11
Wt% As on Catalyst

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
12 SPECIAL EDITION ISSUE No. 110 / 2011
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 100 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, 3600 ppm N, etc.) that U.S. Gulf Coast refiners began processing
back in the 1980s are finding a new market in emerging refining regions such as India. These new highcomplexity
facilities may nonetheless be unfamiliar with their contaminants.

Contaminant Feed Guidlines Common Source Remedies
Si < 1.0 wppm
The contaminants found in these feeds can be particularly harmful to
hydrotreating catalyst activity. In anticipation of the challenges in-
volved in processing these types of feedstocks, there has been a no-
ticeable 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 tar-
gets, for example, with higher catalyst volume and lower space ve-
locity.
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 1980s, that are relatively low in sulfur, may again
become marketable in large quantities. However, they are high in As
Anti foam from
delayed cokers
Na, Ca < 0.5 wppm Sea water; caustic
As < 250 ppb
Pb, P < 0.5 wppm
Ni + V
Fe
< 1.0 wppm total
for Ni + V + Fe
< 1.0 wppm total
for Ni + V + Fe
Crudes from W. Africa,
Russia, synthetic crudes
Gasoline slop tanks,
imported feeds
Resid; heavy feeds
Soluble: corrosion; insoluble:
unfiltered particulates
C7 insolubles MCR < 100 wppm < 0.5 wt% Resid; heavy feeds
Table 1: Feed Contaminants Commonly Found in Crudes
Guard catalysts
Improved desalting;
Guard catalysts;
Don’t send spent
caustic to feed
tanks or units
High Ni guard catalysts
Don’t process feeds
containing Pb or P
Better feed distillation;
Guard Catalysts;
Bed grading
Inert guard material
with high void
space; Fe-traps (soluble);
Top bed skimming
Better feed
distillation;
Guard catalysts;
Bed grading
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

Answers to the 2011 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 sili-
con 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 to-
gether 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 instru-
ment/method can measure very low silicon concentrations.
Fouling/Poisons
4. What has been your experience with antimony and phospho-
rous 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.
14 SPECIAL EDITION ISSUE No. 110 / 2011
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 10°F (6°C) loss in activity. ART recommends
that a feed content of < 0.5 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 ob-
served 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 sev-
eral months the unit required an unplanned turnaround and fresh
catalyst was installed. The spent catalyst analysis indicates the cat-
alysts 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 cata-
lyst in the top half of the bed had lost over 60°F (33°C) of activity.
Sulfiding
6. What is the minimum hydrogen sulfide required in the recy-
cle 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 2S
concentration during the process cycle.
Na, wt% P, wt%
GSK-9 1.71 4.43
GSK-6A 0.21 2.29
Top Catalyst 1.66 2.29
Table 2: Commercial Example of Phosphorus
Poisoning

Start-Up and Shut-Down
7. Is there any harm adding cracked stocks too quickly after
break-in following catalyst activation? What is typical introduc-
tion rate?
Ben Sim
Introducing cracked stocks too early after sulfiding will cause notice-
able 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 passi-
vated 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.
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 particu-
lates 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.
Kerosene/LCO Processing
10. In treating kerosene, what factors play into the decision to
use hydrotreating versus sweetening processes such as caustic
treating?
Dave Krenzke
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 there-
fore the sulfur content of the product is only limited by the process
conditions and catalyst activity. The sweetening process only re-
moves mercaptan sulfur so the product sulfur is limited to the non-
mercaptan 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
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 summa-
rizes 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 100 ppm sulfur down to 10 ppm sulfur.
The 20% LCO blend requires almost 20°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 pol-
yaromatic 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 com-
pleted 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 40°F
(22°C). The decrease in endpoint lowers the total sulfur by almost
1000 ppm and total nitrogen decreases by 129 ppm.
A comparison of activity on the two LCO feeds blended at 30% by
volume into SR feed is shown in Figure 20. Over 30°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

The addition of LCO has a major impact on activity for both the low
and high endpoint LCO materials. The required temperature in-
crease for ULSD in going from 0 to 30% LCO for the lower endpoint
material is about 1.2°F (0.7°C) per percent LCO. Processing the
higher endpoint LCO increases the required temperature to about
1.4°F (0.8°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 ef-
fects are not exactly linear with increasing LCO content. The first
15% LCO has a larger impact on activity than the next 15%.
Volume Gain/Conversion
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?
Charles Olsen
There are a number of parameters which influence volume gain in a
ULSD unit. Hydrogen partial pressure and LHSV are two key oper-
ating conditions which have a large effect on the product volume in-
crease. 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 func-
tion of unit LHSV. Generally speaking, as LHSV decreases the po-
tential volume swell increases. At a LHSV around 1 hr -1 or less,
total product volume increases of 6-7% or more are achievable (pro-
vided 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)
16 SPECIAL EDITION ISSUE No. 110 / 2011
LCO
(high FBP)
API 18.31 15.31
Sulfur, wt% 0.948 1.041
Nitrogen, ppm 708 837
Aromatics, lv% 66.86 68.81
mono-,lv% 22.65 18.44
poly-, lv%
Dist., D2887, ˚F/C
44.21 50.37
IBP 249/121 256/124
10% 425/218 432/222
50% 531/277 550/288
70% 600/316 620/327
90% 677/358 699/371
FBP 772/411 812/433
Table 3: Comparison of Boiling Point Reduction on
LCO
Required Temperature Increase,˚F
70
60
50
40
30
20
10
0
0
0 50 100 150
Product Sulfur, ppm
SR
LCO
Figure 19: Activity Comparison on SR and Blended
SR/LCO
Required Temperature Increase,˚F
160
140
120
100
80
60
40
20
0
0
0 100 200 300 400 500 600
Product Sulfur, ppm
SR 30% Hi FBP LCO
30% Lo FBP LCO
Figure 20: Impact of Endpoint Reduction on
Hydrotreating Performance
WABT Increase to Achieve
Product Sulfur, ˚F
50
45
40
25
35
20
30
25
15
20
15
10
10
5
5
0
0
0 5 10 15 20 25 30 35
% LCO
Hi EP LCO Lo EP LCO
Figure 21: Activity Comparisons at different LCO FBP
and Concentration
35
30
25
20
15
10
5
80
70
60
50
40
30
20
10
Required Temperature Increase,˚C
Required Temperature Increase,˚C
WABT Increase to Achieve
Product Sulfur, ˚C

Total Product Volume, % of Fresh Feed
110
109
108
107
106
105
104
103
102
101
100
0.00 0.50 1.00 1.50 2.00 2.50 3.00 3.50 4.00 4.50
Refiner A Refiner B Refiner C Refiner D
Refiner E Refiner F
LHSV, hr1
Figure 22: Effect of LHSV on Volume Swell in
Commercial Units
Total Product Volume, % of Fresh Feed
27.6
110
41.4 55.2 68.9 82.7 96.5 110.3 124.1 137.9 151.7
109
108
107
106
105
104
103
102
101
Refiner A Refiner B Refiner C Refiner D
Refiner E Refiner F
Reactor Inlet Pressure, Barg
100
400 600 800 1000 1200 1400 1600 1800 2000 2200
Reactor Inlet Pressure, Psig
Figure 23: Effect of Unit Pressure on Volume Swell
Total Product Volume, % of Fresh Feed
110
109
108
107
106
105
104
103
102
101
0.934 0.924 0.914 0.904 0.894 0.884 0.874 0.864 0.854 0.844 0.834 0.824 0.814
100
18.0 20.0 22.0 24.0 26.0 28.0 30.0 32.0 34.0 36.0 38.0 40.0 42.0
Refiner A Refiner B Refiner C Refiner D
Refiner E Refiner F
Feed Specific Gravity
Feed API Gravity
Figure 24: Feed Gravity Has a Significant Impact on
Volume Swell
Of course LHSV is not the only parameter which can influence the
volume swell. H 2 partial pressure also has a significant effect. Fig-
ure 23 summarizes the total product volume yield as a function of
unit pressure for the commercial units shown in Figure 22. Not sur-
prisingly, 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 1000 Psig (69
Barg). The total volume swell increases to 4-7% as the unit pres-
sure increases beyond 1000 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 pres-
sure at similar LHSV.
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.
Total Product Volume, % of Fresh Feed
105.0
104.5
104.0
103.5
103.0
102.5
102.0
101.5
0.50 0.70 0.90 1.10 1.30 1.50 1.70 1.90 2.10 2.30
LHSV, hr -1
NiMo CoMo
Figure 25: The Catalyst Type Effects Volume Swell
ADVANCED REFINING TECHNOLOGIES CATALAGRAM ® 17

13. What considerations are being given to include mild hydroc-
racking in your high pressure ULSD unit?
Robert Wade
Many refiners consider complimenting their existing ULSD HDT cat-
alyst with a hydrocracking catalyst to improve cold flow properties.
This is accomplished through end point reduction. The hydrocrack-
ing 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 consid-
erably higher for a VGO service. In addition, the capital expense re-
quired 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 re-
duction 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 cata-
lysts 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.
18 SPECIAL EDITION ISSUE No. 110 / 2011
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 ApARTTM 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 re-
finers 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 antici-
pated 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 accommo-
date downstream processing. These upstream processes include
coking or an ebullating bed resid operation, followed by a hydrotreat-
ing or hydrocracking operation in order to produce a lighter grade
material. These hydrocracking units tend to operate at severe con-
ditions in conjunction with high hydrogen partial pressures. At these
conditions, the removal of all the easy, less refractory sulfur is read-
ily achieved, and the majority of the multi-ring aromatics are satu-
rated. 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 surpris-
ingly 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.

Light SR Gas Oil LCO EB Diesel Synthetic Diesel FB Diesel
Sulfur, wt% 1.11 0.17 0.017 0.07 0.006
Nitrogen, wppm 138 203 135 261 71
API Gravity 32.37 24.09 31.24 31.96 31.42
Aromatics, vol%
Total 24.18 64.61 41.26 36.58 44.83
mono 14.47 31.97 36.55 32.52 41.95
poly 9.71 32.64 4.71 4.06 2.88
Dist., D2887, ˚F/˚C
0.5 286/141 219/104 303/151 206 / 97 285 / 141
10 492/256 367/186 407/208 349 / 176 379 / 193
50 601/316 470/243 589/309 517 / 269 515 / 268
90 731/388 578/303 707/375 662 / 350 652 / 344
99.5 799/426 644/340 759/404 738 / 392 775 / 413
Sulfur Distribution
Thiophenes 8 0 0 0 0
Benzothiophenes (BT) 2 69 0 0 0
Substituted BT’s 2793 1083 0 26 0
DiBenzothiophenes (DBT) 222 110 0 7 0
Substituted DBT’s 3453 436 72 266 15
4,6 DM-DBT 199 0 78 29 43
C 3-DBT 4410 0 17 370 24
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 signifi-
cant 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 ap-
pearance to straight run (SR) materials. Table 4 lists the properties
for several of these diesel feeds including the diesel product frac-
tions 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

Figure 26 shows the activity difference between a SR and a blended
SR/Synthetic diesel. Note that at higher product sulfur, the two feed-
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.
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.
ART featured in August issue of
Hydrocarbon Engineering
The August 2011 issue featured ART on its cover and in the lead ar-
ticle “Yield of Dreams” by ART’s Brian Watkins and Chuck Olsen and
Grace’s David Hunt.
The collaboration discusses the key relationships between FCC pre-
treat and FCC unit operations and their corresponding catalyst sys-
tems in maximizing the distillate pool. Both processes must be
reoptimized as the refiner moves from gasoline to distillate produc-
tion 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
20 SPECIAL EDITION ISSUE No. 110 / 2011
HYDROCARBONENGINEERING HYDROCARBOON ENGINEERING
AUGUST2011 AU G GUST2011
www.energyglobal.com
www. energygloba .comm
Required Temperature Increase,˚F
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Volume 16 Number 8 - August 2011
SR Synthetic
Figure 26: Activity Comparison of the SR and
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Required Temperature Increase,˚C

-· < '
ART's 420DX Catalyst­
Leading the Pack in ULSD Processing
Confirmed by independent lab testing
According to independent lab testing*, ART's 420DX ranks highest in activity among six leading
ULSD catalysts. 420DX is the newest CoMo member of ART's DX® Catalyst Platform. Depending
on operating conditions and objectives, 420DX 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
7500 Grace Drive
Columbia, MD 21044 USA
v +1.410.531.4000
F +1.410.531.4540
W artcatalysts.com
*Independent lab results available upon request. Advanced Refining Technologies•

© 2011 Advanced Refining Technologies, LLC
SmART Catalyst System ® is a registered trademark in the United States and/or other countries, of W. R. Grace & Co.-Conn.
ApART ®, StART ® and DX ® are registered trademarks of Advanced Refining Technologies, LLC.
ART ® and ADVANCED REFINING TECHNOLOGIES ® are trademarks, registered in the United States and/or other countries,
of Advanced Refining Technologies, LLC.
This trademark list has been compiled using available published information as of the publication date of this brochure and may not
accurately reflect current trademark ownership.
artinfo@grace.com
www.e-catalysts.com
The information presented herein is derived from our testing and experience. It is offered, free of charge, for your consideration,investigation and
verification. Since operating conditions vary significantly, and since they are not under our control, we disclaim any and all warranties on the results
which might be obtained from the use of our products. You should make no assumption that all safety or environmental protection measures are
indicated or that other measures may not be required.