The Global - Grace · The Global Hydroprocessing ......
Transcript of The Global - Grace · The Global Hydroprocessing ......
Advanced Refining Technologies (ART) and Chevron Lummus
Global (CLG) together offer refiners the leading global source for
hydroprocessing from concept to commercial operation.
Worldwide provider of complete range of
hydroprocessing catalysts including CLG
hydrocracking and lubes processing
catalysts.
World-class technology development,
licensing, design and revamp for
hydroprocessing.
The GlobalHydroprocessingPuzzle is Now Complete
www.artcatalysts.com www.chevronlummus.com
Catalagram®
ISSUE 113, Spring 2013
Editor:Rosann Schiller
Contributors:Kenneth Bryden
E. Thomas Habib, Jr.
Charles Olsen
Brian Watkins
Gordon Weatherbee
Guest Contributors:Jimmy Crosby
Allen Hansen
Gautham Krishnaiah
Barry Speronello
Please address your comments to:[email protected]
Grace Catalysts Technologies7500 Grace DriveColumbia, MD 21044410.531.4000
© 2013 W. R. Grace & Co.-Conn.
3 Flexible Pilot Plant Technology for Evaluation ofUnconventional Feedstocks and Processes
By Kenneth Bryden, Gordon Weatherbee, and E. ThomasHabib, Jr., Grace Catalysts Technologies
This article includes comparisons of DCRTM pilot plant results to commercial
FCC units for petroleum derived gas oil and resid feeds and also describes
application of the DCRTM pilot plant to a variety of alternative feedstocks and
process designs. Testing experiences with vegetable oil, pine-derived pyrol-
ysis oil, and straight run shale oil are described, highlighting the utility of the
DCR unit in evaluating these feedstocks and understanding their effects on
yields and operation. Furthermore, applications of the DCR in studying new
high temperature cracking processes designed for high light olefins yields
and processing very light feeds in a circulating fluidized bed are described.
22 Rive Molecular HighwayTM Catalyst Delivers Over$2.50/bbl Uplift at Alon’s Big Spring, Texas Refinery
By Gautham Krishnaiah, Barry Speronello and Allen Hansen,Rive Technology, Inc. and Jimmy Crosby, Alon USA
Last year Rive successfully trialed the first generation of Molecular High-
way™ technology on a paraffinic VGO feed in the CountryMark refinery
FCCU. This year Rive has successfully demonstrated the second generation
of its technology, on a resid feed, at the Alon USA FCCU in Big Spring, TX.
3.15.13
33 Custom Catalyst Systems for Higher Yields of Diesel
By Brian Watkins, Manager, Hydrotreating Pilot Plant andTechnical Service and Charles Olsen, Director, Distillate R&Dand Technical Service, Advanced Refining Technologies
Both the hydrotreating catalyst system and the operating strategy for the
ULSD unit are critical to providing the highest quality products. Driving the
hydrotreater to remove sulfur and PNA's improves product value, but this
needs to be balanced against the increased costs of higher hydrogen con-
sumption. Use of tailored catalyst systems can optimize the ULSD hy-
drotreater in order to produce higher volumes of high quality products while
balancing the refiners available hydrogen.
2 Issue No. 113 / 2013
EditorialLast issue I reflected on the anniversary of cat cracking; an amazing
achievement for a process originally expected to run for only five years or
until World War II was over. Yet 70 years later, the FCC process chugs
along, albeit with changing expectations. Feedstocks have become either
much heavier or much lighter – both presenting their own set of unique
challenges. Product yields and quality have continually improved as
regional demand patterns changed and clean fuels regulations took effect.
Catalysts have also evolved over the years to be more active, more
selective, and drive the right conversion to meet the changing market
demands.
In this issue, we feature three AFPM papers that will help you prepare and plan for the future. Now is an exciting
time of change for refining. We see new feed sources and novel processing schemes being considered for
existing process units. Renewable fuels regulations are also impacting operating strategies. The first paper
discusses the essential role of pilot plant testing to evaluate alternate processing schemes and feedstocks.
Grace’s DCR circulating riser has been proven to be an excellent tool to understand process effects and gauge
risk. The second paper, describes the second commercial application of Rive’s Molecular Highway FCC catalyst
technology. The last paper from my colleagues at Advanced Refining Technologies (ART), summarizes catalytic
options and operating strategies to increase middle distillate yields with minimal investment.
And speaking of ART, this month ART signed an agreement with Chevron Lummus Global (CLG) that gives ART
the exclusive right to sell CLG's hydrocracking and lubes hydroprocessing catalysts to CLG's licensees and other
petroleum refiners for unit refills. The agreement will streamline hydroprocessing catalyst supply and improve
technical service for refining customers by establishing ART as the single point of contact for all their
hydroprocessing catalyst needs. Through the agreement, ART and CLG can provide their customers broader
service and more advanced catalytic materials to improve the competitiveness and profitability of their refineries.
Sincerely,
Rosann K. SchillerSenior Marketing ManagerGrace Catalysts Technologies
Rosann K. SchillerEditor
Grace Catalysts Technologies Catalagram® 3
Kenneth BrydenManager, FCCEvaluations Research
Gordon WeatherbeePrincipal Engineer
E. Thomas Habib, Jr.Director CustomerResearch Partnershipsand DCR LicensingManager
Grace CatalystsTechnologiesColumbia, MD, USA
AbstractThe fluid catalytic cracking process has been in commercial practice for 70+ years. Feedstocks and
process designs have evolved greatly over this period. Today is a time of exciting change in the FCC
world. New feedstocks such as bio-oils (vegetable and pyrolysis), and straight run shale oils are being in-
vestigated by refiners. New FCC designs such as High Severity Fluid Catalytic Cracking (HS-FCC) and
Deep Catalytic Cracking (DCC) have been developed. On-purpose olefins manufacturing processes, such
as ExxonMobil PCCSM and KBR Superflex™, and use of FCC type processes for very light feeds (includ-
ing gases and light alcohols) are being proposed. These options represent significant change, and there-
fore significant risk. One way to minimize the risk associated with these opportunities is to conduct
realistic pilot plant testing prior to commercial implementation. One pilot unit that has gained wide accept-
ance in mimicking commercial FCC operation is Grace's DCR™ pilot plant. Including the two DCR pilot
plants operated by W. R. Grace & Co., a total of 26 licensed DCR pilot units have been constructed
throughout the world. This paper includes comparisons of DCR pilot plant results to commercial FCC
units for petroleum derived gas oil and resid feeds, and also describes application of the DCR pilot plant
to a variety of alternative feedstocks and process designs. Testing experiences with vegetable oil, pine-
derived pyrolysis oil, and straight run shale oil are described, highlighting the utility of the DCR unit in
evaluating these feedstocks and understanding their effects on yields and operation. Furthermore, appli-
cations of the DCR in studying new high temperature cracking processes designed for high light olefins
yields and processing very light feeds in a circulating fluidized bed are described.
IntroductionFluid catalytic cracking is one of the most flexible processes in a refinery. It can readily adjust to changes
in feed quality through modifications to catalyst and operating conditions. The FCC unit is one of the few
units in a refinery that can handle a variety of feedstocks, including highly impure feedstocks. FCC feed-
stocks have changed over the 70+ years of commercial application, evolving from light gas oils feeds (31°
API) in the 1940’s, to a variety of streams in the present day which may contain resid, syncrude, as-
phaltenes, and hydrotreated feedstocks1. The flexibility of the FCC unit is of great interest to refiners in
utilizing unconventional feedstocks. A variety of unconventional feedstocks are under consideration for
Flexible Pilot Plant Technology forEvaluation of UnconventionalFeedstocks and Processes
4 Issue No. 113 / 2013
motor fuels production. Government mandates on renewable fuel
standards have resulted in interest in co-processing vegetable oils
and pyrolysis oils in refineries2. New technologies are being devel-
oped to convert waste plastics to synthetic crude oil3. The introduc-
tion of new drilling and extraction technologies such as horizontal
drilling and hydraulic fracturing has resulted in large quantities of
shale oil becoming available4.
The flexibility of the fluidized catalytic cracking process, where a
circulating fluidized bed provides excellent heat and mass transfer,
and where a reaction step can be coupled with a catalyst regener-
ation step, have resulted in adoption of FCC-type processes for
applications outside of the conventional molecular weight reduction
of the heavy fraction of crude oil to produce motor fuels. New de-
signs for high temperature cracking to produce light olefins from
heavy feedstocks have been developed, such as High Severity
Fluid Catalytic Cracking (HS-FCC)5, and Deep Catalytic Cracking6.
FCC-type processes such as ExxonMobil PCCSM7 and KBR Super-
flex™8 are designed to crack naphtha range feedstocks preferen-
tially to light olefins. Circulating fluidized bed processes have
been proposed for converting biomass to motor fuels9 and biomass
to benzene, toluene and xylene10,11. FCC-type processes have
also been developed for propane dehydrogenation12,13 and for con-
verting methanol to olefins14,15. Clearly, circulating fluidized beds
are a versatile technology and are not limited to converting gas oil
to motor fuels.
New feedstocks and process designs represent significant change,
and therefore significant risk. Understanding the potential yields
and performance is vital in assessing the economic viability of
feedstock and process changes. One way to minimize the risk as-
sociated with these opportunities is to conduct realistic pilot plant
testing prior to commercial implementation. As Leo Baekeland, an
entrepreneur and pioneer in the plastics industry, famously spoke
to the importance of lab and pilot plant testing when he stated in
his 1916 Perkin Medal acceptance speech- “The principle: ‘Com-
mit your blunders on a small scale and make your profits on a
large scale,’ should guide everybody who enters into a new chemi-
cal enterprise.”16 Conducting testing before commercial implemen-
tation reduces risk for a refiner or petrochemical manufacturer.
Examples of questions that can be answered via testing include:
What will be the effect of a potential feedstock change on yields
and product quality?
What will be the effect of a new feedstock on operating conditions?
What are the optimum process conditions to maximize desired
yields?
Description of DCRTM Pilot PlantPerformance testing of FCC catalysts can be done by either bench
scale testing or pilot plant scale testing. Examples of bench scale
testing equipment include fixed bed microactivity testing (MAT)17
and fixed fluidized bed testing, one example of which is the ACE™
catalyst evaluation instrument marketed by Kayser Technology18.
Several pilot plant designs are in operation throughout the world
and include both once-through and circulating designs. The most
common is the Grace-developed DCRTM pilot plant. Table I pro-
vides a comparison of the conditions in these test units to commer-
cial operation.
MAT and ACE testing have the advantages that they are easy to
set up and require small amounts of material. However, these
units cannot provide the detailed product analysis or feedback on
extended operation that pilot scale units can. Larger scale test
equipment such as a pilot unit can provide sufficient liquid product
for distillation and detailed analysis (such as API gravity and aniline
point on LCO produced, viscosity of bottoms, octane engine testing
of gasoline, etc.) and can provide information on continuous opera-
tion. Additionally, compared to bench scale units, the DCR pilot
plant has the advantage that it mimics all the processes present in
commercial operation and it can operate at the same hydrocarbon
partial pressure as a commercial unit. The continuous catalyst re-
generation in the DCR allows for the measurement of regenerator
SOx and NOx emissions and testing of environmental additives,
experiments which cannot be done in a batch unit.
The continuous nature of the DCR and the fact that it represents
the commercial FCC process is particularly important when evalu-
ating new process designs based on FCC. A 2007 study by Inde-
pendent Project Analysis19 that examined the success of 850
TABLE I: Comparison of Test Units to Commercial Conditions
MAT/ACE Circulating Riser Commercial
Nature of Operation Unsteady State Steady State Steady State
Catalyst Contact Time 12 to 150 secs 2 to 5 secs 2 to 8 secs
Temperature Range 930 to 1100°F 930 to <1100°F 980 to 1030°F
Hydrocarbon Partial Pressure ~12 psia 20-45 psia 20-50 psia
Catalyst Inventory 5 to 10 grams 2 to 3 kg 100 tons
Advantage Easy to set up Mimics commercialoperation
Grace Catalysts Technologies Catalagram® 5
capital projects involving new technology found that “An integrated
pilot run for an extended period of time can dramatically improve
the early operability of new technology processes.” They found
that revolutionary new technology projects that had a pilot facility to
provide basic operability data averaged 79 percent of design ca-
pacity seven to twelve months after startup, while comparable
processes for which pilot facilities were not built only achieved 30
percent of design capacity seven to twelve months after startup.
They concluded “With a pilot facility, operating conditions can be
fully explored and optimal operating ranges established.”
Figure 1 is a schematic drawing of the DCR. The range of typical
operating conditions of the DCR is shown in Table II. The system
consists of three main units - a riser, a stripper, and a regenerator.
Both the regenerator and the stripper are equipped with slide
valves for control of catalyst circulation rate. The DCR riser is typi-
cally operated in adiabatic mode, where changing feed preheat or
regenerator temperature will result in a change in catalyst circula-
tion to maintain reactor outlet temperature, the same process con-
trol strategy used in many commercial FCC units. The catalyst
circulation and thus, the catalyst to oil ratio, is varied by changing
the feed preheat temperature. During operation of the DCR, a me-
tering pump precisely controls the feed rate as feed is pumped
from the load cell through a preheater. Nitrogen and steam, in-
jected through a separate preheater/vaporizer, are used as a feed
dispersant. Catalyst and product pass from the riser to the stripper
overhead disengager. Products exit the disengager through a re-
frigerated stabilizer column to a control valve, which maintains unit
pressure at the desired level. A section of the stripper-regenerator
spent catalyst transfer line consists of a shell and tube heat ex-
changer. The rate of heat transfer across this exchanger provides
a precise and reliable method to calculate the catalyst circulation
rate. The stabilizer column, also called the debutanizer column, is
ControlValveMeter
Feed Pump
Feed Preheater
Dispersant Steam Stripping Steam
Liquid Product Receivers
FeedStorageTank #1
FeedStorageTank #2
FeedTank
FeedTank
Scale Scale
Reg
ener
ato
r
Ris
erR
eact
or
Hea
tE
xch
ang
er
Str
ipp
er
Co
nd
ense
r
Sta
bili
zer
Co
lum
n
MeterControlValve
FIGURE 1: Schematic Diagram of Grace DCRTM Pilot Plant
Operating Condition Min/Max Range Typical for FCC-type operations
System Pressure max. 45 psig (4atm)* 25 psig (2.7 atm)
Catalyst Charge 1.5 – 3.5 liters 2.0 liters
Catalyst Circulation Rate 2500 – 15,000 grams/hr 4,000 – 9,000 grams/hr
Feed Rate 300 – 1500 grams/hr 1000 grams/hr
Feed Pre-heat Temperature 150 – 750°F 300 – 700°F
Riser Temperature 500°F -1300°F* 970 – 1000°F
Regenerator Temperature max. 1375°F* 1300°F
Stripper Temperature max. 1300°F* 950°F
Stabilizer Column Temperature minimum -30°F 0°F
* note: higher maximum pressures and temperatures can be achieved by constructing the unit with specialized alloys
TABLE II: DCRTM Pilot Plant Operating Ranges
6 Issue No. 113 / 2013
operated to separate C4 minus from the liquid product, which is
condensed and collected. The collected liquid is analyzed by GC to
determine its composition. Product can also be collected for sub-
sequent physical analysis. (Under typical FCC operating condi-
tions, approximately one liter of liquid product is generated per
hour.) The gaseous products are metered and batch collected for
subsequent analysis by GC. The carbon on regenerated catalyst
can be maintained at various levels by controlling the regenerator
operating conditions. The continuous nature of the DCR and the
circulation of catalyst between riser and regenerator make it well-
suited to study the effect of process conditions and additives on
fuel sulfur20, and air pollutant emissions such as SOx21 and NOx22.
For pollution control studies, the regenerator temperature can be
varied to match those found in a commercial FCCU. Excess air
and combustion products exit the regenerator through a pressure
control valve and are then metered and continuously analyzed for
O2, CO2, and CO and optionally for SO2 and NO. A more detailed
description of the unit is available in Reference 23.
The DCR's ability to match commercial yields is due in large part to
the adiabatic reactor operating system which controls the reactor
temperature and catalyst circulation rate in much the same manner
as most commercial FCCU's. In this mode, the reactor is setup
with insulation and heaters to prevent any heat from being added
to or lost from the sides of the reactor. The reactor temperature is
then controlled by the amount of hot catalyst added to the reactor
from the regenerator. This operating mode can successfully match
a commercial operation, not only in yields and conversion, but also
in key process variables like catalyst to oil ratio when operating at
the same reactor exit, feed, and regenerator (catalyst) tempera-
tures. While the DCR is significantly smaller than a commercial
unit, it closely matches the operation of commercial units. A study
by Independent Project Analysis on “Best Practices in Process De-
velopment” determined that “Research has shown that the scale
factor of the pilot to the commercial unit is less important than the
fact that the pilot truly represents the commercial facility”19. Several
studies have been done in the DCR by using commercial equilib-
rium catalysts, feeds, and operating conditions to compare yields
obtained from the DCR with commercial yields. Typically, the DCR
yields match very closely the commercial yields at similar condi-
tions. Sample data showing the close match between DCR data
and commercial data are shown in Table III for a gas oil feed and
Table IV for a resid containing feed. In both cases the coke yield
from the DCR is ~10-15% lower than the coke yield in the commer-
cial unit. This is because the DCR has excellent stripping due to
the small diameter of the stripper and the increased residence time
relative to a commercial unit (note that while the DCR reactor oper-
ates in adiabatic mode, the overall unit is not necessarily heat bal-
anced since the regenerator temperature can be controlled
independent of coke yield). When coke from unstripped hydrocar-
bons in the commercial unit is accounted for, the coke match of the
DCR to commercial becomes even better.
Due to its simplicity of operation and ability to match commercial
yields, the DCR has become the leading commercially available
technology for small scale FCC pilot units. There are currently 26
DCR technology licenses worldwide.
Pilot Plant Work withUnconventional FeedstocksThe DCR has been used to process a variety of petroleum based
feedstocks from hydrotreated VGO to resid feedstocks. The DCR
has been shown to routinely process most feedstocks containing
up to 5 wt.% Conradson carbon and limited success has been pos-
sible with feeds containing up to 9.7 wt.% Conradson carbon25.
In addition to conventional feedstocks, the DCR has been able to
process straight run crude oil, naphthas, gases, and feeds from
non-petroleum sources. Naphthas and gases require a modified
feed system but otherwise generally process similar to standard
feeds. Non-petroleum based feedstocks vary widely in their char-
acteristics and, while some are easily processed in the DCR, oth-
ers are extremely difficult to run, if they can be run at all. Three
illustrative examples of processing unconventional feedstocks are
given below.
Straight Run Shale OilThe introduction of novel drilling technologies has resulted in large
amounts of oil from shale becoming available in North America.
While fluid catalytic cracking is typically done to reduce the molec-
ular weight of the heavy fractions of crude oil (such as vacuum gas
oil and atmospheric tower bottoms), in some cases refiners are
DCR FCCU
Riser Temperature, °F 959 959
C/O 6.6 5.9
Conversion, wt.% 67.2 66.2
Yields, wt.%
Fuel Gas 2.2 2.3
LPG 9.2 8.7
Light Gasoline (C5 – 302°F) 31.4 31.1
RON 93.3 93.1
MON 79.4 78.3
Heavy Gasoline (302-365°F) 7.2 6.4
Naphtha (365-500°F) 13.1 12.7
LCO (500-644°F) 11.3 13.3
HCO (644°F+) 21.4 20.4
Coke 3.9 4.5
TABLE III: Comparison of DCR to Commercial FCCUnit Run at Same Operating Conditions Using a GasOil Feed (from Reference 24)
Grace Catalysts Technologies Catalagram® 7
DCR Run 1 DCR Run 2 Refiner A
Rx Exit Temp, °F 1000 1000 1000
Regen Catalyst Temp, °F 1366 1366 1366
Feed Temp, °F 486 486 486.4
Rx Exit Pressure, psig 40.0 40.1 28.9
Rx Exit HC Pressure, psia 35.3 35.5 19.1
Riser Bot HC Pressure, psia 14.2 14.2 24.3
Cat/Oil Ratio 7.7 7.7 6.5
Conversion, wt.% 76.9 77.5 77.5
Kinetic Conversion 3.34 3.44 3.44
H2 Yield, wt.% 0.17 0.17 0.15
C1 + C2's, wt.% 4.1 4.1 3.9
Total C3, wt.% 6.8 6.8 6.1
C3=, wt.% 5.5 5.5 4.7
Total C4, wt.% 11.2 11.5 10.2
Gasoline, wt.% 50.0 50.1 51.6
RON (DCR - Est from GCON® software) (92.5) (92.6) 91.1
MON (DCR - Est from GCON® software) (80.0) (80.1) 81.6
LCO, wt.% 14.3 13.9 14.9
Bottoms, wt.% 8.8 8.7 7.6
Coke, wt.% 4.7 4.8 5.4
TABLE IV: Comparison of DCR to Commercial FCC Unit Run at Same Operating Conditions Using aResid-Containing Feed
charging whole shale oil as a fraction of their FCC feed. Also,
whole crude oil has been charged to FCC units when gas oil feed
is not available due to maintenance on other units in the refinery26,
and to produce a low-sulfur synthetic crude27.
As a model case to understand the cracking of whole crude oil in
the FCC and the effect of process conditions on yields, a straight
run shale oil was processed in the DCR at three riser outlet tem-
peratures: 970°F, 935°F, and 900°F. The whole crude oil was a light
sweet Bakken crude, with an API of 42°. The properties of the
crude were similar to those given in a publically published assay28.
Table V presents a comparison of the properties of the whole crude
used by Grace and the publically available assay data. Additionally,
the straight run Bakken sample was distilled into a 430°F minus
gasoline cut and a 430°F-650°F LCO cut and the properties of
these cuts were measured. Gasoline from the straight Bakken was
highly paraffinic and had low octane numbers (a G-Con® RON soft-
ware of 61 and MON of 58). The LCO fraction had an aniline point
of 156°F and an API gravity of 37.6°, resulting in a diesel index of
59. The catalyst used in the experiments was a high matrix FCC
catalyst, deactivated metals-free using a CPS type protocol. The
properties of the deactivated catalyst are given in Table VI.
8 Issue No. 113 / 2013
Bakken sampleused in Grace work
Published Bakken assaydata from Reference 28
API Gravity Degrees 41.9 >41
Sulfur wt.% 0.19 <0.2
Distillation Yield wt.% vol.%
Light Ends C1-C4 1 3
Naphtha C5-330°F 32 30
Kerosene 330-450°F 14 15
Diesel 450-680°F 25 25
Vacuum Gas Oil 680-1000°F 23 22
Vacuum Residue 1000+°F 5 5
Total 100 100
Conradson Carbon Residue wt.% 0.78
Gasoline Fraction PropertiesG-CON® RON software 60.6
G-CON® MON software 57.6
LCO Fraction (430°F - 650°F)properties Aniline point (˚F) 155.9
API Gravity 37.6
Diesel Index 58.6
TABLE V: Properties of Straight Run Shale Oil Feed Used by Grace Compared to Publically Published Assay Data
C/O Ratio
C5+ Gasoline, wt.% LCO (430-650˚F), wt.%
Dry Gas, wt.% Coke, wt.%
Bottoms (650˚F+), wt.%
Conversion, wt.%
75.0 80.0 85.0 75.0 80.0 85.0 75.0 80.0 85.0
10.0
8.0
6.0
70.0
62.5
4.0
65.5
67.5
60.0
1.50
1.25
1.00
0.75
0.50
20.0
17.5
15.0
12.5
10.0
2.2
2.0
1.8
1.6
1.4
5.0
4.0
3.0
2.0
900˚F 935˚F 970˚FReactor Temperature
FIGURE 2: Effect of DCR Riser Outlet Temperature on Yields of Straight Run Shale Oil
Grace Catalysts Technologies Catalagram® 9
For the three different reactor outlet temperatures, plots of catalyst
to oil ratio, dry gas, gasoline, LCO, bottoms and coke yields versus
conversion are shown in Figure 2. As expected, lowering reactor
temperature increases the amount of LCO produced. As seen in
the graphs, cracking straight run shale oil produces little coke and
bottoms. At the same conversion level, lowering reactor tempera-
ture results in slightly more gasoline yield (due to increased C/O),
which is consistent with prior Grace work29. Plots of gasoline
olefins, iso-paraffins and RON and MON estimated via G-Con®
software are shown in Figure 3. Cracking straight run Bakken
shale oil produces a low-quality gasoline with research octane less
than 80 and motor octane less than 70. At constant conversion, in-
creasing reactor temperature results in more gasoline olefins and
higher research octane number.
Diesel quality is of great interest to refiners. Syncrude produced in
the DCR runs was distilled to recover the 430°F to 650°F LCO frac-
tion. Aniline point and API gravity of the LCO were then measured
to allow calculation of the diesel index, a measure of LCO quality.
[Diesel Index = (aniline point x API Gravity) / 100] Figure 4 pres-
ents data for LCO yield and LCO quality as a function of conver-
sion. As seen in the data, increasing conversion lowers LCO
quality as a result of increased cracking of the LCO range paraffins
to lighter hydrocarbons. Similar to prior Grace work30, LCO quality
follows LCO yield and did not appear to be influenced by reactor
temperature at constant conversion. Diesel index values of the
LCO produced by cracking whole shale oil were significantly higher
than values obtained with typical VGO feeds.
As seen in the results from this study, widely varying ratios of prod-
ucts and product quality can be obtained by changing process con-
ditions. Information from pilot studies such as this one helps
refiners to determine the optimum processing setup to maximize
yields of desired products. The ability of the DCR to produce suffi-
cient liquid product for properties testing assisted greatly in the
measurement of LCO quality.
Total Surface Area, m2/g 196
Zeolite Surface Area, m2/g 110
Matrix Surface Area, m2/g 86
Unit Cell Size, Å 24.30
Rare earth, wt.% 2.1
Alumina, wt.% 52.1
78.0
77.0
76.0
75.0
74.0
17.0
16.0
15.0
14.0
13.0
70.0
69.0
68.0
67.0
66.0
26.0
25.5
25.0
24.5
24.0
75.0 77.5 80.0 82.5 85.0 75.0 77.5 80.0 82.5 85.0
G-Con® Software RON EST
G-Con® Software O, wt.% G-Con® Software I, wt.%
G-Con® Software MON EST
Conversion, wt.%
900˚F 935˚F 970˚FReactor Temperature
FIGURE 3: Effect of DCR Riser Outlet Temperature on Gasoline Properties of Cracked Straight Run Shale Oil
TABLE VI: Deactivated Catalyst Properties for WholeShale Oil Study
10 Issue No. 113 / 2013
Vegetable OilGovernment mandates on renewable biofuels have resulted in in-
terest in using vegetable oils and Fisher-Tropsch waxes obtained
from biomass. Vegetable oils could be co-fed with VGO to an FCC
unit31, or fed in their entirety32-34. While refiners would be highly un-
likely to ever process a 100% vegetable oil in a FCC unit, a 100%
soybean oil feed was chosen as a test case for pilot DCR work to
understand the impact this type of feed would have on yields and
operation. As a control case, a standard mid-continent VGO was
run. The catalyst was a low metals refinery equilibrium catalyst. A
riser outlet temperature of 970°F was used. Properties of the feed-
stocks are presented in Table VII. Note that the simulated distilla-
tion of the soybean oil is based on the carbon content and
molecular weight of the material and this can sometimes skew the
estimated boiling points. Biofeed sources typically have a true
boiling point that is much lower than that reported by simulated dis-
tillation equipment due to molecular weight interference. Proper-
ties of the equilibrium catalyst used in the testing are presented in
Table VIII. Figure 5 presents yield curves at constant coke. Figure
6 presents gasoline properties at constant coke. Table IX presents
yields of soybean oil and VGO at the same operating conditions.
On a constant coke basis, the soybean oil produced more LCO,
less gasoline, less C3’s, and less C4’s than the VGO. The gasoline
produced by cracking soybean oil was highly aromatic, consistent
with the results of References 33-35. Gas Chromatography-Atomic
Emission Detector (GC-AED) was performed in oxygen mode on
the liquid product in order to detect oxygen species, and only trace
amounts of oxygenates were found. While running soybean oil,
CO and CO2 were detected in the product gas, amounting to a total
of ~15% of the oxygen in the soybean oil. By difference, ~85% of
the oxygen in the soybean oil reacted to water. The DCR riser op-
erates in adiabatic mode. In typical endothermic gas oil cracking,
the riser bottom is ~70°F hotter than the riser top36. Interestingly
for the soybean oil cracking, the riser temperature profile was al-
most flat, with only a 10°F temperature difference between the riser
bottom and top. Figure 7 presents adiabatic riser temperature pro-
Soybean Oil Mid Continent VGO
°API 21.6 24.7
Sulfur, wt.% 0.00 0.35
Oxygen, wt.% 10.5 0.0
D2887 Distillation, °F
IBP 702 527
5% 1059 651
10% 1069 691
30% 1090 773
50% 1102 848
70% 1111 928
90% 1183 1045
95% 1232 1108
FBP 1301 1259
TABLE VII: Feedstock Properties for StudyComparing Vegetable Oil to a Mid-Continent VGO
22.0
20.0
18.0
16.0
12.0
14.0
10.0
30.0
25.0
35.0
40.0
45.0
75.0 77.5 80.0 82.5 85.0 75.0 77.5 80.0 82.5 85.0
LCO (430-650˚F), wt.% Diesel Index
Conversion, wt.%
900˚F 935˚F 970˚FReactor Temperature
FIGURE 4: Effect of Conversion Level on LCO Yield and Quality for Straight Run Shale Oil
Grace Catalysts Technologies Catalagram® 11
Total Surface Area, m2/g 171
Zeolite Surface Area, m2/g 134
Matrix Surface Area, m2/g 37
Unit Cell Size, Å 24.35
Rare earth, wt.% 3.2
Alumina, wt.% 44.2
Nickel, ppm 30
Vanadium, ppm 80
TABLE VIII: Equilibrium Catalyst Properties forSoybean Oil and Pyrolysis Oil Testing
Coke, wt.%
Soybean Oil VGO
C/O Ratio Total C3, wt.% Total C4, wt.%
Gasoline, wt.% LCO, wt.% Bottoms, wt.%
12.0
10.5
9.0
7.5
6.0
52.5
50.0
47.5
45.0
7.0
6.0
5.0
4.0
22.5
20.0
17.5
15.0
25.0
14.0
12.0
10.0
8.0
6.0
5.0
4.5
4.0
3.54.8 5.6 6.4 4.8 5.6 6.4 4.8 5.6 6.4
FIGURE 5: Yields at Constant Coke for 100% Soybean Oil and a Mid-Continent VGO with a 970°F Riser OutletTemperature
files for soybean oil and VGO at the same operating conditions
(250°F preheat, 1300°F catalyst temperature, 970°F Riser Outlet
Temperature.) Based on the temperature drop across the riser, the
heat of cracking of soybean oil is only about 15% of the heat of
cracking of standard vacuum gas oil, consistent with the exother-
mic formation of carbon monoxide, carbon dioxide and water from
oxygen present in the soybean oil. This heat behavior results in
the soybean oil running at a significantly lower catalyst to oil ratio
than VGO under the same conditions. The discovery of this very
interesting effect of running 100% soybean oil (which has implica-
tions for riser operation) shows the utility of the DCR in testing un-
conventional feedstocks and understanding their processing
implications.
Pine-based Pyrolysis OilDue to government renewable fuel credits and mandates, there is
considerable refiner interest in using bio-based feedstocks. Co-
processing bio-based pyrolysis oils with conventional vacuum gas
oil (VGO) has been proposed as one method of incorporating bio-
based feedstock into motor fuels37.
12 Issue No. 113 / 2013
VGO Feedstock Soybean Oil
Riser Outlet Temperature, ˚F 970 970
Feed Temperature, ˚F 1300 1300
Feed Temperature, ˚F 250 248
Pressure, psig 25.2 25.1
C/O Ratio 9.3 6.7
H2 Yield, wt.% 0.02 0.04
C1 + C2's, wt.% 2.1 1.9
Total C3’s, wt.% 6.7 4.3
C3, wt.% 1.1 0.6
C3=, wt.% 5.6 3.8
Total C4’s, wt.% 12.4 6.2
Total C4=, wt.% 6.8 4.3
C4=, wt.% 1.6 1.1
LPG Olefinicity 0.65 0.76
Gasoline (C5-430°F), wt.% 53.1 44.5
G-Con® software P, wt.% 3.6 3.5
G-Con® software I, wt.% 29.8 22.0
G-Con® software A, wt.% 33.9 39.0
G-Con® software N, wt.% 11.9 13.2
G-Con® software O, wt.% 20.8 22.4
G-Con® software RON EST 90.2 90.9
G-Con® software MON EST 79.5 79.0
LCO (430-700°F), wt.% 15.4 22.0
Bottoms (700°F+), wt.% 4.9 3.9
Coke, wt.% 5.2 4.6
Fuel Gas CO, wt.% 0.0 1.2
Fuel Gas CO2, wt.% 0.0 0.9
Fuel Gas H2O, wt.% (by difference) 0.0 10.3
TABLE IX: Yields at Same Operating Conditions for Base Case VGO and 100% Soybean Oil
Water content, wt.% 23.0
Carbon (as-is), wt.% 39.5
Hydrogen (as-is), wt.% 7.5
Oxygen (as-is), wt.% (by difference) 53.0
Carbon (dry basis), wt.% 55.5
Hydrogen (dry basis), wt.% 6.5
Oxygen (dry-basis), wt.% (by difference) 38.0
TABLE X: Properties of Pine-Derived Pyrolysis Oil used in VGO Co-Processing Experiments
Grace Catalysts Technologies Catalagram® 13
Many groups have published work on co-processing pyrolysis oil
and VGO where the testing was done in a batch fashion in ACE or
MAT units38-44. Continuous pilot operations can identify processing
issues that are not readily apparent in batch testing. Due their high
content of reactive oxygen containing compounds, pyrolysis oils
are not as stable as conventional petroleum feedstocks and have a
tendency to polymerize and form tars at elevated temperatures
(140°F-212°F)45,46. We are aware of two published reports of circu-
lating pilot plant work with blends of pyrolysis oil and petroleum
based feedstocks47,48. Lappas, et. al.47 describe pilot scale work in
the CPERI FCC circulating pilot plant. They attempted to co-
process the heavy fraction of thermally hydrotreated biomass flash
pyrolysis liquid (HBFPL) with VGO. This material had a 4.9 wt.%
oxygen content. They found that it was necessary to dilute the
HBFPL oil in light cycle oil to prevent plugging of the nozzle in their
pilot plant. Their final feed to the FCC pilot unit was 2.25 wt.%
HBFPL / 12.75 wt.% LCO / 85 wt.% VGO.
G-Con® Software RON EST G-Con® Software MON EST G-Con® Software P, wt.%
G-Con® Software I, wt.% G-Con® Software A, wt.% G-Con® Software O, wt.%
4.8 5.6 6.4 4.8 5.6 6.4 4.8 5.6 6.4
90.8
90.6
90.4
90.2
30.0
27.5
25.0
22.5
20.0
91.0
79.75
79.50
79.25
79.00
80.00
38.0
36.0
34.0
40.0
21.6
20.4
19.2
18.0
22.8
3.6
3.5
3.7
18.0
Coke, wt.%
Soybean Oil VGO
FIGURE 6: Gasoline Properties Versus Coke for Soybean Oil and Mid-Continent VGO with a 970°F Riser OutletTemperature
Temperature, ˚F
960 970 980 990 1000 1010 1020 1030 1040 1050 1060
Soybean Oil VGO
Incr
easi
ng
Ris
erH
eig
ht
FIGURE 7: Adiabatic Riser Temperature Profiles for100% Soybean Oil and a Mid-Continent VGO Run atSame Catalyst Temperature and Same Feed Preheatwith Target Riser Outlet Temperature of 970°F
14 Issue No. 113 / 2013
Grace work in the DCR has also found that continuous processing
of pyrolysis oils can be difficult due to the high tendency of pyroly-
sis oil to form coke and plug the feed nozzle. Modifications to the
DCR feed delivery system were made that enabled co-processing
of pyrolysis oil with VGO in a continuous fashion. As a model
case, a blend of 3 wt.% pine-based pyrolysis oil was co-processed
with 97 wt.% mid-continent VGO using a low-metals commercial
equilibrium catalyst. The VGO properties are provided in Table VII
and the equilibrium catalyst properties are presented in Table VIII.
The properties of the pyrolysis oil feedstock are given in Table X.
The pyrolysis oil was not hydrotreated and contained 23 wt.%
water. The composition of the pyrolysis oil was 39.5 wt.% carbon,
7.5 wt.% hydrogen and 53 wt.% oxygen. 100% mid-continent VGO
was cracked as a control case. Riser outlet temperature was
970°F for both feeds. Yields at identical operating conditions are
presented in Table XI. Co-feeding pyrolysis oil resulted in more
coke, less gasoline, and production of CO and CO2 in the product
gas. These results are consistent with the observations of other
100% VGO 3 wt.% Pine-Based Pyrolysis Oil –97 wt.% VGO
Rx Exit Temperature, ˚F 970 970
Catalyst Temperature, ˚F 1300 1300
Pressure, psig 25 25
Conversion, wt.% (100-LCO-bottoms) 81.6 81.7
Kinetic Conversion 4.42 4.46
C/O Ratio 9.9 9.6
H2 Yield, wt.% 0.05 0.04
C1 + C2's, wt.% 3.15 2.97
Total C3’s, wt.% 8.51 8.05
C3, wt.% 2.57 2.55
C3=, wt.% 5.94 5.51
Total C4’s, wt.% 14.1 13.8
Total C4=, wt.% 5.9 5.5
Gasoline (C5-430°F), wt.% 49.1 47.5
G-Con® software P, wt.% 3.2 3.2
G-Con® software I, wt.% 24.3 24.5
G-Con® software A, wt.% 49.2 50.5
G-Con® software N, wt.% 9.3 9.3
G-Con® software O, wt.% 14.0 12.6
G-Con® software RON EST 92.5 92.1
G-Con® software MON EST 81.6 81.5
LCO (430-700°F), wt.% 14.1 14.2
Bottoms (700°F+), wt.% 4.4 4.2
Coke, wt.% 6.4 7.1
Fuel Gas CO, wt.% 0.0 0.48
Fuel Gas CO2, wt.% 0.0 0.11
Fuel Gas H2O, wt.% (by difference) 0.0 1.42
TABLE XI: Yields at Same Operating Conditions for Base Case Mid-Continent VGO and Blend of 3 wt.% Pine-BasedPyrolysis Oil and 97 wt.% VGO
researchers who processed high oxygen content pyrolysis oils47-49.
At the same feed preheat and catalyst temperature, the blend of
pyrolysis oil and VGO required ~0.3 less cat to oil to maintain a
970°F riser outlet temperature with the DCR operated in adiabatic
mode. We speculate that the exothermic reactions of the oxygen
in the pyrolysis oil reduce the heat requirements for co-processing
pyrolysis oil with VGO. Gas Chromatography-Atomic Emission De-
tector (GC-AED) was performed in oxygen mode on the liquid
product in order to detect oxygen species and only trace amounts
of oxygenates were found. While running pyrolysis oil, CO and CO2
were detected in the product gas, amounting to a total of ~22 per-
cent of the oxygen in the pyrolysis oil. By difference, ~78% of the
oxygen in the pyrolysis oil reacted to water. As seen by these re-
sults with pyrolysis oil, non-petroleum based feedstock compo-
nents can result in significant yield shifts, even at small addition
quantities. The DCR pilot plant has proven to be an invaluable tool
in understanding these yield shifts.
Grace Catalysts Technologies Catalagram® 15
Pilot Plant Work on UnconventionalProcessesAs mentioned in the introduction, the circulating fluidized bed tech-
nology of FCC is being applied to a wide range of processes in-
tended for a variety of conversions, including: heavy oil to olefins,
naphtha streams to olefins, paraffins to propylene, light alcohols to
olefins, and biomass to olefins and aromatics. Pilot plant work is
essential in reducing the risk of scaling up a new process. An ex-
ample of application of DCR technology to process development is
work done by Nippon Oil and King Fahd University in developing
their High Severity FCC process5. In their published work, they de-
scribe how they converted the DCR from a riser pilot plant to a
downer pilot plant. In comparing their pilot plant to their demon-
stration plant, they wrote: “the pilot plant and demonstration plant
performed similarly. It also confirmed that scaling up the process
was successful.”5
To show the versatility of FCC-type technology, three illustrative
examples of evaluating unconventional processes in the DCR pilot
plant are given below.
High Temperature Cracking forLight OlefinsThe high rate of growth in propylene demand has resulted in inter-
est in producing propylene from processes other than traditional
steam cracking. New designs for high temperature cracking to pro-
duce light olefins from heavy feed stocks have been developed,
such as High Severity Fluid Catalytic Cracking (HS-FCC)5, and
Deep Catalytic Cracking6. These processes typically operate at
higher temperatures and more severe conditions than typical FCC
operations. Pilot equipment such as the DCR can be used to eval-
uate the effect of different operating conditions on process yields.
Using data from the DCRTM pilot plant, Grace published an exten-
sive study on the effect of ZSM-5 additive concentration (0 to 8
wt.%) and reaction temperature (970°F to 1050°F) on olefins
yields50. Grace has also published DCR pilot plant results on the
effect of hydrocarbon partial pressure on propylene production51.
Presented below are three additional examples of work done in
Grace’s pilot plants using the DCR to gain insight into high temper-
ature cracking for light olefins.
To examine the effect of feedstock on light olefins production at
high temperature, cracking was done on a light VGO feed and a
resid feed using a blend of base catalyst and a ZSM-5 containing
additive at a riser outlet temperature of 1050°F. Feedstock proper-
ties are given in Table XII.
Interpolated yields at constant cat to oil ratio are presented in Table
XIII. Under these conditions high yields of propylene and butylene
were produced by both feeds. However, as expected, the heavier
feedstock did generate higher coke and lower light olefin yields at
the same catalyst to oil ratio.
VGO Feedstock Resid Feedstock
°API Gravity 23.9 20.6
K Factor 11.81 11.76
Refractive Index 1.5064 1.5222
Sulfur, wt.% 0.73 0.42
Basic Nitrogen, wt.% 0.04 0.07
Total Nitrogen, wt.% 0.10 0.18
Conradson Carbon, wt.% 0.33 5.10
ndm analysisArom Ring Carbons Ca, wt.% 19.6 25.4
Naphthenic Ring CarbonCn, wt.% 20.6 15.4
Paraffinic Carbons Cp, wt.% 59.8 59.2
Ni, ppm 0.5 6.6
V, ppm 0.2 16.5
Simulated Distillation, °F
IBP 464 455
10% 637 653
30% 730 793
50% 806 894
70% 883 1017
90% 977 1265
End Point 1152 1324
TABLE XII: Properties of Feedstocks for Study of Feedstock Effect on High Temperature Cracking for Light Olefins
16 Issue No. 113 / 2013
Determining the effect of added steam on yields is another exam-
ple of the insight that can be gained via pilot plant experimentation.
A mixture of equilibrium catalyst and lab deactivated ZSM-5 was
used to crack the vacuum gas oil described in Table XII at a riser
outlet temperature of 1050°F. Normally, the steam used for feed
atomization is about 3 wt.% of fresh feed. In this study, atomiza-
tion steam was varied between 3 wt.% and 18 wt.% of fresh feed to
understand the effect of increasing steam level on yield structure.
Higher steam rates are expected to reduce hydrocarbon partial
pressure, and reduce the residence time, favoring olefins maxi-
mization. Increasing the steam rate reduces the residence time,
resulting in lower conversion at the same cat to oil ratio. Table XIV
presents interpolated yields at constant conversion for three steam
levels. At constant conversion, increasing the steam level resulted
in the expected higher propylene and butylenes yields.
To examine the effect of going to very high temperatures, cracking
was done at riser outlet temperatures of 1050°F and 1100°F on the
VGO Feedstock Resid Feedstock
Conversion, wt.% 74.7 70.6
Kinetic Conversion 2.89 2.32
H2 Yield, wt.% 0.07 0.08
C1 + C2's, wt.% 7.2 6.6
Total C3’s, wt.% 16.7 14.8
C3=, wt.% 14.6 13.1
Total C4’s, wt.% 13.8 12.2
Total C4=, wt.% 11.2 10.3
Gasoline (C5-430°F), wt.% 34.3 30.5
G-Con® software ,P wt.% 3.2 3.2
G-Con® software I, wt.% 10.1 11.2
G-Con® software A, wt.% 51.5 51.7
G-Con® software N, wt.% 7.0 7.3
G-Con® software O, wt.% 28.1 26.7
G-Con® software RON EST 98.3 97.4
G-Con® software MON EST 83.7 83.8
LCO (430-700°F), wt.% 16.5 17.3
Bottoms (700°F+), wt.% 8.8 12.2
Coke, wt.% 2.0 5.6
TABLE XIII: Interpolated Yields at C/O = 11 for Two Feedstocks at a Riser Outlet Temperature of 1050°F
3 wt.% Added Steam 10 wt.% Added Steam 18 wt.% Added Steam
Cat/Oil Ratio 8.6 11.5 15.6
H2 Yield, wt.% 0.08 0.08 0.08
C1 + C2's, wt.% 8.4 8.0 8.2
C2=, wt.% 4.1 4.1 4.5
Total C3’s, wt.% 15.1 15.4 16.9
C3, wt.% 1.9 2.0 1.7
C3=, wt.% 13.2 13.4 15.2
Total C4’s, wt.% 11.3 11.7 12.1
Total C4=, wt.% 9.6 9.7 10.1
Gasoline (C5-430°F), wt.% 30.4 29.9 27.6
LCO (430-700°F), wt.% 18.8 18.7 18.4
Bottoms (700°F+), wt.% 14.2 14.3 14.6
Coke, wt.% 1.4 1.7 2.0
TABLE XIV: Interpolated Yields at 67 wt.% Conversion at Three Different Steam Levels on VGO Feed at a RiserOutlet temperature = 1050°F
Grace Catalysts Technologies Catalagram® 17
vacuum gas oil described in Table XII using a blend of base cata-
lyst and a large proportion of ZSM-5 based additive. Table XV
presents interpolated yields at constant cat to oil for the two riser
outlet temperatures. Increasing riser outlet temperature from
1050°F to 1100°F resulted in higher conversion and higher light
olefins yields. However, the increase in reactor temperature also
resulted in greater thermal cracking as seen in the higher dry gas
yields at 1100°F.
The preceding three examples show how a flexible pilot plant can
be used to quickly conduct studies to provide insight into the ef-
fects of operating variables like feedstock, steam level and temper-
ature for processes designed for high temperature production of
olefins.
Processing Naphtha FeedsDemand for ethylene, propylene and other chemical feedstocks
has resulted in refiners examining naphtha cracking, and in the de-
velopment of FCC-type processes such as ExxonMobil PCCSM7
and KBR Superflex™8 to crack naphtha range feedstocks prefer-
entially to light olefins. In evaluating new processes, it is important
to understand the effects of critical variables like feedstock,
process conditions and catalyst. Instituto Colombiano de Petróleo
(a DCR licensee), published a study where they used their DCR
pilot plant to evaluate the potential yields of four different naphtha
feedstocks52. These feedstocks ranged in API from 53° to 60° and
included straight run naphthas and naphthas from FCC operations.
Table XVI provides a summary of some of their findings. The re-
searchers at Instituto Colombiano de Petróleo (ICP) found that
feedstock had an important effect on product yields. Compared to
straight run naphtha, FCC naphtha produced less propane, less
butane and iso-butane and more toluene and xylenes. While the
researchers at ICP used the DCR pilot plant to focus on the effect
of the naphtha feedstock type at typical FCC process conditions, a
DCR pilot plant could be readily used to evaluate the effect of tem-
peratures and severities greater than typical FCC conditions on
converting naphtha to olefins.
Alcohols to OlefinsEthanol and methanol have both been proposed as petrochemical
feedstocks. Ethanol can be produced via fermentation and then
reacted via dehydrogenation to produce ethylene. Methanol can
be produced from coal or from natural gas. Catalytic processes
such as methanol to olefins can then be used to convert the
methanol into valuable products like ethylene and propylene. Sev-
eral reactor designs for MTO have been proposed, including fixed
bed reactors, fluidized bed reactors and riser reactors15,53. While
methanol is much lighter than conventional FCC feeds, modifica-
tions to the DCR feed system enabled the processing of methanol.
As a model case, a blend of 50 wt.% methanol and 50 wt.% water
was reacted over a SAPO-34 based catalyst in the DCR at a series
of increasing riser temperatures. Figure 8 presents ethylene and
propylene yield as a function of riser temperature. Consistent with
other published work54, olefins yield increased with temperature
over this operating range. The exothermic nature of the methanol
to olefins reaction was clearly apparent in the temperature profile
of the adiabatic riser. In typical endothermic gas oil cracking, the
riser bottom is hotter than the riser top. In the case of methanol to
olefins, the riser bottom was ~30°F cooler than the riser top, even
with the addition of 50% water in the feed as a heat sink. This ex-
ample shows that the DCR can be used to examine unconven-
tional processes beyond the traditional feeds and process
conditions associated with fluid catalytic cracking.
ConclusionsThe DCR pilot unit is an excellent tool for simulating commercial
FCC units. When run at the same operating conditions with the
same feedstock and catalyst, the DCR produces yields nearly
identical to commercial FCC units. The DCR can also be used to
test unconventional feedstocks to determine their suitability as
feeds for commercial FCC units. The ability of the DCR to produce
sufficient quantity of liquid product for properties testing greatly en-
hances the measurement of LCO quality. The adiabatic reactor
operating system can provide insight into the temperature control
behavior of non-petroleum feedstocks. The flexibility of the DCR
allows for evaluation of process conditions and modes of operation
outside of typical FCC conditions. Feedstocks and process de-
signs will continue to change and evolve and pilot plant testing is a
key step in evaluating these changes. Pilot plant testing reduces
risk and uncertainty by identifying the optimum feedstocks and
process conditions on the lab scale so that fuel and petrochemical
manufacturers can “make their profits on a large scale.”
Riser Outlet Temperature, ˚F
700 750 800 850 900 950 1000
C2
+C
3O
lefi
ns
Yie
ld
FIGURE 8: Olefins Yield as a Function of Riser OutletTemperature for Reacting a 50 wt.% Methanol/50wt.% Water Blend Over SAPO-34 Based Catalyst
18 Issue No. 113 / 2013
AcknowledgementsThe hard work and dedication of the technicians and operators as-
sociated with Grace’s DCR pilot plants is gratefully acknowledged.
References1. R.P. Fletcher, “The History of Fluidized Catalytic Cracking: A
History of Innovation: 1942-2008,” In Innovations in Industrial and
Engineering Chemistry; Flank, W., et al.; ACS Symposium Series;
American Chemical Society: Washington, DC, 2008, pp. 189-249.
2. “Renewable Fuel Standard: Potential Economic and Environ-
mental Effects of U.S. Biofuel Policy,” Committee on Economic and
Environmental Impacts of Increasing Biofuels Production; National
Research Council, National Academies Press, 2011.
3. M.M. Bomgardner, “Transforming Trash,” Chemical and Engi-
neering News, November 5, 2012, pp. 19-21.
4. “Review of Emerging Resources: U.S. Shale Gas and Shale
Oil Plays,” U.S. Energy Information Agency, July 2011.
5. M.H. Al-Tayyar, A.B. Fox, C.F. Dean, Y. Fujiyama, T. Okuhara,
A.M. Aitani, M.R. Saeed, “Development of a Novel Refinery
Process: From Laboratory Experiments to Commercial Applica-
tions,” Saudi Aramco Journal of Technology, Spring 2008, pp. 2-9.
6. D. Dharia, W. Letzsch, L. Chapin, “Advanced Catalytic Crack-
ing Technologies for Production of Light Olefins from Low Cost Re-
finery Based Feedstocks,” Catalagram® Number 94 (2004), pp.
37-41.
7. M.W. Bedell, P.A. Ruziska, T.R. Steffens, “On-Purpose Propy-
lene from Olefinic Streams,” Catalagram® Number 94 (2004), pp.
33-36.
8. C. Eng, R. Orriss, M. Tallman, “Meeting Propylene Demands
with SUPERFLEX Technology,” Catalagram® Number 94 (2004),
pp. 27-30.
9. KBR Press Release titled “KBR Awarded EPC Contract by
KiOR, Inc. for Biomass-to-Renewable Crude Project in the United
States,” Houston, Texas, April 18, 2011.
10. K. Bourzac, “From Biomass to Chemicals in One Step,” MITTechnology Review, March 29, 2010.
Riser Outlet Temperature = 1050°F Riser Outlet Temperature = 1100°F
Conversion, wt.% 69.0 75.3
Kinetic Conversion 2.25 3.06
H2 Yield, wt.% 0.09 0.14
C1 + C2's, wt.% 8.8 13.2
CH4 Yield, wt.% 1.8 3.4
C2, wt.% 1.5 2.7
C2=, wt.% 5.5 7.1
Total C3’s, wt.% 16.0 18.5
C3=, wt.% 12.8 15.0
Total C4’s, wt.% 12.3 11.9
Total C4=’s, wt.% 9.5 9.7
Gasoline (C5-430°F), wt.% 30.4 29.9
G-Con® software P, wt.% 3.0 1.8
G-Con® software I, wt.% 9.2 7.1
G-Con® software A, wt.% 62.0 74.6
G-Con® software N, wt.% 6.2 3.8
G-Con® software O, wt.% 19.7 12.7
G-Con® software RON EST 99.4 101.8
G-Con® software MON EST 85.9 88.4
LCO (430-700°F), wt.% 18.3 15.3
Bottoms (700°F+), wt.% 12.7 9.4
Coke, wt.% 1.4 1.5
TABLE XV: Interpolated Yields at C/O = 13 for VGO Feedstock at Two Riser Outlet Temperatures
Grace Catalysts Technologies Catalagram® 19
FeedstockNaphthenic Straight-
Run NaphthaParaffinic Straight-
Run NaphthaLight Naphtha fromModel IV FCC Unit
Total Naphtha fromUOP II FCC Unit
Feedstock API° 53.2 50.1 60.0 58.6
Feedstock PIANO
Paraffins, wt.% 42.2 56.7 42.0 28.4
Iso-paraffins, wt.% 31.7 34.2 35.0 22.7
Olefins, wt.% 7.5 1.3 20.8 33.7
Aromatics, wt.% 15.7 13.2 26.1 29.7
Naphthenes, wt.% 34.6 28.8 11.1 8.2
Product Yields
H2 0.1 0.1 0.2 0.1
Total Dry Gas 4.2 3.7 4.5 4.4
Total LPG 29.5 28.9 22.4 22.4
C2 0.8 0.7 0.9 0.8
C2= 2.0 1.9 2.0 2.2
C3 6.4 5.6 3.5 2.6
C3= 6.6 7.8 7.5 8.4
nC4 2.9 2.7 1.5 1.5
iC4 10.3 8.4 5.7 4.8
Naphtha (C5-430°F) 60.9 63.1 64.3 65.0
Benzene 1.5 1.7 1.6 1.9
Toluene 6.5 5.9 6.8 8.9
Xylenes 9.8 6.5 9.2 11.0
LCO (430-650°F) 1.8 1.5 3.5 4.1
Slurry (>650°F) 0.7 0.6 1.4 1.2
Coke 2.8 2.1 3.7 2.8
TABLE XVI: Effect of Naphtha Feedstock Properties on Product Yields from DCR Pilot Plant (C/O = 15, 1000°FReaction Temperature, with 4% ZSM-5 Additive). Adapted from Reference 52.
11. “Biochemical Startup Announces Para-Xylene Breakthrough,”
Chemical Week, December 5, 2012.
12. D. Sanfilippo, F. Buonomo, G. Fusco, I. Miracca, “Paraffins
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17. ASTM D3907 - 03(2008) Standard Test Method for Testing
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22 Issue No. 113 / 2013
Rive Molecular HighwayTM CatalystDelivers Over $2.50/bbl Upliftat Alon’s Big Spring, Texas Refinery
Gautham Krishnaiah1
Director of TechnicalService
Dr. Barry SperonelloResearch Fellow
Allen HansenProcess ModelingEngineer
Rive Technology, Inc.Monmouth Junction, NJUSA
Jimmy CrosbyVice President,Refining
Alon USABig Spring, TX, USA
AbstractRefiners are continuously striving to expand margins by optimizing their operations. Flexible FCC catalyst
technologies that offer better coke selectivity and bottoms upgrading can help refiners overcome operating
constraints, make a more valuable product slate, and optimize profitability. FCC catalyst innovation the last
decade has largely focused on achieving these results through improved matrices, binders and additives,
rather than through improvements to the zeolite component of the catalyst.
Rive has focused its research on the zeolite component of the FCC catalyst and has developed Molecular
Highway™ mesoporous zeolite technology for improved mass transfer into and within zeolite crystals. The
enhanced porosity of the zeolite, when incorporated into FCC catalysts, allows FCC feed molecules to more
readily access the interior of the zeolite, undergo the desired reactions, and then quickly exit, leading to im-
proved selectivity – namely, better coke selectivity, lower bottoms yields and higher yields of products such
as gasoline, diesel, and light olefins, depending on the refiner’s objectives.
Last year Rive successfully trialed the first generation of Molecular Highway™ technology on a paraffinic
VGO feed in the CountryMark refinery FCCU. This year Rive has successfully demonstrated the second
generation of its technology, on a resid feed, at the Alon USA FCCU in Big Spring, TX.
This paper will discuss Molecular Highway technology and the results from the trial at Alon’s Big Spring, TX
refinery, where a W. R. Grace & Co.-manufactured catalyst containing Rive zeolite was used successfully to
achieve over $2.50/bbl value uplift in the FCCU.
IntroductionIn 2011, Rive successfully trialed its first generation mesoporous zeolite, called Molecular Highway™ tech-
nology, at the CountryMark Refining FCC unit in Mount Vernon, Indiana.
1 Currently with KBR Corp., Houston, TX, USA
Grace Catalysts Technologies Catalagram® 23
The catalyst demonstrated good:
• hydrothermal stability
• activity maintenance
• attrition resistance and fluidization
• coke selectivity and bottoms cracking, with an overall increase
in transportation fuels.
At the AFPM’s 2012 Annual Meeting, Rive reported the development
and manufacturing scale-up of a second generation (Gen II) version
of the Rive zeolite. Since then, Grace was granted a permit for the
commercial production of Gen II Rive zeolite, and 125 tons were
produced for a second refinery trial.
Using this commercially-produced Rive® zeolite and Grace® matrix
technology, Rive/Grace developed catalyst formulations for ACE
testing with Alon’s feedstock. The ACE unit is an industry accepted
tool for evaluating FCC catalysts in the laboratory and then predict-
ing commercial performance. Our results were used to build an eco-
nomic model for the Alon, Big Spring refinery that predicted a
$2.00/FCC bbl value increase over the incumbent catalyst, substan-
tial enough to easily justify a commercial trial.
Formulation DevelopmentBased on learnings from Rive’s first successful trial at CountryMark
Refining, a second generation of mesoporous zeolite was developed
(Gen II) for production at Grace’s Valleyfield, Canada catalyst plant.
Nominally 125 tons of Rive zeolite were produced, and this commer-
cial zeolite was used for all formulation test work. Testing conditions
were designed to simulate Alon’s Big Spring operation. All catalysts
were impregnated with nickel and vanadium and cyclic propylene
steam (CPS) deactivated for testing in an ACE unit.
Table I contains the results of ACE testing comparing Alon’s incum-
bent catalyst (which incorporated both state-of–the-art matrix metals
trapping and matrix bottoms upgrading technologies) with Rive’s
proposed formulation (referred to as the Rive MH-1 catalyst) on
Alon’s FCC unit feed (22° API gravity, 1.6% CCR, and 2% sulfur).
They show a substantial improvement in coke selectivity and a sub-
stantial increase in gasoline yield with the Rive catalyst. The antici-
pated increase in bottoms cracking at constant conversion was
smaller than typical, but was much larger when converted to a con-
stant coke basis (as was evident in actual unit operation).
The ACE test results were used in modeling (Profimatics™) the im-
pact of Rive’s MH-1 catalyst on the Alon FCC unit operation. The
yields predicted3 with a 100% change-out of the catalyst to Rive’s
formulation would provide a value uplift of $2.00/FCC feed bbl using
the refinery’s constraints and product pricing. It should be noted
that the MH-1 catalyst’s improved coke selectivity also predicted that
optimized operation would occur at both lower reactor and lower re-
generator temperatures.
Grace manufactured 328 tons of MH-1 catalyst for Alon. Rive MH-1
had an average zeolite surface area of 227 m2/g, a matrix4 surface
area of 100 m2/g, and a Grace Attrition Index (DI) of 6. The incumbent
catalyst had a similar fresh zeolite surface area and a slightly lower
matrix surface area. Catalyst quantity was sufficient for a 109-day trial
(at 3 tons/day) and was projected to yield an 80% change-out.
Unit DescriptionThe Alon, Big Spring, TX FCC unit is a UOP stacked design, re-
vamped to include an external vertical riser with state-of-the-art feed
injection nozzles. The riser terminates into a pair of primary cy-
clones that discharge the spent catalyst via dip legs in the stripper.
The spent catalyst is subjected to steam for stripping absorbed
product vapors before it flows down the spent catalyst standpipe into
the regenerator. The riser product vapors exiting the primary cy-
clone gas tubes are quenched by LCO sprays. The quenched prod-
uct vapors and stripping steam exit the reactor via a pair of
secondary cyclones which further separate out entrained catalyst.
Coke on catalyst is burned off in the regenerator, which is operated
in a partial combustion mode. Combustion air is supplied by three
air blowers working in parallel. The flue gas flows through four pairs
of cyclones which recover and return entrained catalyst to the re-
generator bed. After exchanging heat in a flue gas steam generator
cooler, the CO-rich flue gas is incinerated to CO2 in a CO boiler.
Cooled flue gas from the CO boiler flows through an Electro-Static
Precipitator (ESP) before it is discharged to the atmosphere.
Catalyst2 IncumbentCatalyst
Rive Catalyst(MH-1)
C/O Ratio 6.2 6.4
Conversion 75.0 75.0
Yields, wt.%Dry Gas, wt.% 3.37 3.19
LPG, wt.% 15.64 15.94
Propane, wt.% 0.83 0.83
Propylene, wt.% 4.67 4.77
Butanes, wt.% 3.85 3.95
Butenes, wt.% 6.29 6.39
Gasoline, wt.% 50.39 51.33
LCO, wt.% 19.09 19.14
Bottoms, wt.% 5.91 5.86
Coke, wt.% 5.60 4.54
TABLE I: Rive MH-1 Catalyst Showed Lower Coke andIncreased Gasoline in ACE Testing
2 Excludes Grace gasoline sulfur reduction additive D-PRISM™
3 Attachment 1 - Yields and operation predicted for Alon (Big Spring, TX) with a 100%
catalyst change-out to MH-1
4 Proxy for mesoporous surface area
24 Issue No. 113 / 2013
The reactor vapors are fractionated into various products in the main
fractionator and gas plant. The main fractionator operates with three
side-draws (HCN, LCO and HCO5) and a bottom draw (CSO6) – to
provide improved column operation and fractionation. The LCO
product (HCN + LCO) is routed to the diesel hydrotreater. Catalyst
entrained with the CSO is recovered in a Gulftronic® Separator, and
the recovered catalyst is returned to the reactor riser. The overhead
from the fractionator is routed to a wet gas compressor (WGC) and
gas plant and fractionated into gasoline, LPG and fuel gas. The
gasoline product is hydro-desulfurized to produce ultra-low sulfur
gasoline. LPG is separated into C3’s and C4’s, and the C4 olefins.
Trial DescriptionThe feed to the FCC unit is a mix of vacuum gas oils and mildly hy-
drotreated PDA7 oil.
Prior to the Rive trial, the unit typically operated with a riser tempera-
ture at or above 1000°F, to maximize LCO minus conversion. Reac-
tor product vapors were quenched with LCO to minimize dry gas
rate. In addition, the regenerator bed temperature was controlled to
maintain a carbon on regenerated catalyst (CRC) below 0.3 wt.%.
The main constraints on the FCC unit are the air blower, wet gas
compressor, catalyst circulation and main fractionator (in that order).
Combined feed temperature was maintained to minimize coke yield.
During summer (May – September), the feed rate to the FCC unit
cycles daily, varying by between 1,000 BPSD and 1,500 BPSD - in-
versely with ambient temperature due to the air blower limit.
Rive® MH-1 catalyst reached the refinery on June 28th, and the trial
began on June 30th. Catalyst addition rate was 3 tons/day – the
same as the incumbent catalyst. One feed sample and three ecat
samples were collected per week. All collected samples were
shipped to Grace for analysis.
As the Rive® catalyst changed-out in the unit, coke selectivity im-
proved and feed rate to the FCC unit steadily increased despite an
increase in the ambient temperatures. Slurry yields decreased with
an increase in slurry density due to improved bottoms upgrading.
Based on the improvement in coke selectivity and improved bottoms
upgrading, the refinery progressively increased feed rate and pro-
gressively lowered the riser and regenerator temperatures by a total
of 20 degrees over a period of several weeks while maintaining less
than 0.3% CRC. Towards the end of the trial, fresh feed to the unit
had been increased by 2,000 BPSD, albeit part of the increase was
due to a lowering ambient temperature. However as the data analy-
sis shows, the improved coke selectivity of the Rive catalyst allowed
the refiner to process an additional 700 BPSD over the prior opera-
tion at similar ambient temperatures.
Feed PropertiesThe FCC feed properties of API gravity and Conradson Carbon
residue (CCR) are shown in Figures 1 and 2, respectively. The
gravity of the feed processed during the trial ranged within the
norms for this unit. However the CCR during the trial was higher
than typical.
Equilibrium Catalyst AnalysesEcat samples were analyzed by Grace to monitor the catalyst activ-
ity, zeolite and matrix surface areas, coke and gas factors, and
physical and chemical properties.
The catalyst change-out shown in Figure 3, is calculated on the basis
of catalyst chemical composition. It rises normally with time until
there is a “dip” after a unit shutdown towards the end of the trial when
Ecat from early in the trial was added to the unit. Following unit
startup and resuming normal catalyst additions, the catalyst change-
Rive CatalystIncumbent Catalyst
22.5
24.0
25.0
Feed
API
Gra
vity
22.0
21.0
21.5
23.5
24.5
23.0
20.5
20.003Nov1105Aug11 01Feb12 01May12 30Jul12 28Oct12
2.0
Feed
CC
R,w
t.%
1.6
1.8
1.4
1.2
1.0
0.803Nov1105Aug11 01Feb12 01May12 30Jul12 28Oct12
Rive CatalystIncumbent Catalyst
FIGURE 1: FCC Unit Feed API
FIGURE 2: FCC Unit Feed CCR
5 HCN = Heavy Cat Naphtha; LCO = Light Cycle Oil; HCO = Heavy Cycle Oil
6 CSO = Clarified Slurry Oil
7 PDA = Propane de-asphalted oil
Grace Catalysts Technologies Catalagram® 25
face areas. The stability of the Rive® catalyst zeolite and mesoporous
(matrix) surface areas is shown in Figures 7 and 8.
The Rive® catalyst also demonstrated a reduction in the coke8 and
gas9 factors (Figures 9 and 10).
Flue gas opacity (an indicator of catalyst losses) was steady as the
Rive catalyst replaced the prior incumbent catalyst during the trial
(Figure 11). The opacity meter was reset (calibrated) early in the
trial and that prevented comparisons with prior opacity measure-
ments. Spent catalyst withdrawals and disposal were normal during
the trial leading to the conclusion that the Rive® catalyst’s in-unit re-
tention and attrition resistance were similar to that of the prior incum-
bent catalyst.
The in-unit catalyst fluidization characteristic, as measured by the
ratio of the minimum bubbling velocity to the minimum fluidization
velocity (Umb/Umf) was constant during the transition from the in-
cumbent to the Rive® catalyst and during the trial (see Figure 12).
The unit did not have any circulation issues, but issues with the re-
out curve was re-established, albeit with a “break”. A regression fit to
this data prior to the ‘disturbance’ (dashed red line) projects a catalyst
change-out of 78% in the absence of the ecat addition.
The change in metals (Ni and V) on Ecat during the trial is shown in
Figure 4. Nickel on Ecat steadily increased from 1200 to 1600 ppm.
Vanadium increased from about 1900 ppm to 2500 ppm over the
same period.
Figure 5 shows the change in regenerator temperature during the
trial and it shows the reduction which occurred during operation on
the Rive MH-1 catalyst.
During the trial, the Rive® catalyst addition rate was maintained at the
same level as the prior catalyst. Despite a greater than 30% increase
in Ecat contaminant metals (Ni & V) and periods of high temperature
operation, the Rive® catalyst demonstrated good hydrothermal stabil-
ity and maintained activity (Figure 6). With no significant differences
in fresh catalyst zeolite and matrix surface areas between the incum-
bent and Rive® catalyst, no changes were expected in the in-unit sur-
90
Cat
alys
tCha
nge-
Out
,%80
70
60
50
40
30
20
10
026Jun1227May12 26Jul12 25Aug12 24Oct1224Sep12
FIGURE 3: Catalyst Change-Out
28Mar1218Jan12 06Jun12 15Aug12 24Oct12
3,000
Ecat
Met
als,
pmw
2,600
2,800
2,400
2,000
2,200
1,800
1,400
1,600
1,200
1,000
Vanadium
Nickel
Rive CatalystIncumbent Catalyst
FIGURE 4: Ecat Metals (Nickel and Vanadium)
Dilute Phase
29Aug1231May12 30Jun12 28Sep12 28Oct1230Jul12
80
Del
taR
egen
erat
orTe
mpe
ratu
re,˚
F
40
60
20
-20
0
-40
Dense Phase
Rive CatalystIncumbent Catalyst
FIGURE 5: Regenerator Temperatures
28Mar1218Jan12 06Jun12 15Aug12 24Oct12
80
MAT
Act
ivity
,wt.%
Con
vers
ion
76
78
74
72
70
68
Rive CatalystIncumbent Catalyst
FIGURE 6: Ecat Activity
26 Issue No. 113 / 2013
generator flue gas analysis caused calculated circulation rate to ap-
pear constant when, in fact, it rose during the trial in response to im-
proved coke selectivity (manifesting as constant conversion at
increased feed rate and lower reactor temperature).
YieldsAhallmark of Rive’s mesoporous Molecular Highway zeolite is that it
provides improved coke selectivity over current FCC catalyst technolo-
gies employing conventional zeolites. Refiners with FCC units con-
strained by air blower rate10 take advantage of improved coke selectivity
in a variety of ways (singly or in combination) through the ability to:
• increase FCC unit feed rate
• raise conversion severity
• lower regenerator and riser temperatures
• process heavier feedstocks
The FCC unit at Alon, Big Spring is constrained by air blower capac-
ity, particularly in summer. The constraint is severe enough that dur-
ing summer, the FCC feed rate cycles daily in sync with the ambient
temperature, as air density impacts the air blower rate. Conse-
quently the refinery builds up an inventory of unprocessed FCC
feedstock over the summer.
Alon Operations was able to consistently and continuously increase
feed throughput, even at the height of summer. Figure 13 shows
that the unit processed a higher feed rate while on the Rive® catalyst
at a given ambient temperature. On average while on the Rive® cat-
alyst, the FCCU was processing over 700 BPSD of additional feed.
At the start of the trial with the higher riser temperature used with the
incumbent catalyst, CSO (bottoms) yield and gravity decreased to near
the gravity limit. Alon Operations was able to take advantage of the
improved bottoms cracking of the Rive® catalyst and lower the riser
temperature without exceeding pre-trial bottoms yields. Simultane-
ously, regenerator temperature fell without affecting the carbon on re-
generated catalyst. As a result of these changes CSO gravity
remained within limits while bottoms selectivity remained low and the
split between gasoline and LPG was shifted to favor more valuable
gasoline.
Eventually, riser temperature was lowered by 20°F while the regen-
erator bed temperature decreased 30°F in partial burn at <0.3%
CRC (Figures 14 and 15). Typically a 1°F drop in riser temperature
lowers regenerator bed temperature 0.8-1.0°F. During the trial, the
regenerator bed temperature clearly dropped by more than this typi-
cal value, indicating that it was due to the improved coke selectivity
of the Rive® catalyst.
28Mar1218Jan12 06Jun12 15Aug12 24Oct12
120Ze
olite
Surf
ace
Are
a,m
2 /gm
110
115
105
100
95
90
Rive CatalystIncumbent Catalyst
FIGURE 7: Ecat Zeolite Surface Area
60
Mat
rixSu
rfac
eA
rea,
m2 /g
m
50
55
45
40
35
3028Mar1218Jan12 06Jun12 15Aug12 24Oct12
Rive CatalystIncumbent Catalyst
FIGURE 8: Ecat Matrix/Mesopore Surface Area
28Mar1218Jan12 06Jun12 15Aug12 24Oct12
2.0
Cok
eFa
ctor
1.6
1.8
1.4
1.2
1.0
0.8
Rive CatalystIncumbent Catalyst
FIGURE 9: Ecat Coke Factor
8 Coke factor is defined as coke yield per unit of kinetic conversion
9 Gas factor is the ratio of Hydrogen to Methane in the Dry Gas yield
10 An informal refining survey suggests that about 60% of FCC units are constrained
by air blower rate
Grace Catalysts Technologies Catalagram® 27
Post 50% Catalyst Change-Out
Incumbent Catalyst
Mean Ambient Temperature, ˚F
2,500
Del
taFe
edR
ate,
BPS
D
2,000
1,500
1,000
500
0
-50040 50 60 70 80 90 100
Rive Catalyst
FIGURE 13: FCCU Feed Rate Increased by 700 BPSD
5
Del
taR
iser
Top
Tem
pera
ture
,˚F
-5
0
-10
-20
-15
-2530Jun1231May12 30Jul12 29Aug12 28Oct1228Sep12
Rive CatalystIncumbent Catalyst
FIGURE 14: Riser Top Temperatures
29Aug1231May12 30Jun12 28Sep12 28Oct1230Jul12
20
Del
taR
egen
erat
orTe
mpe
ratu
re,˚
F
0
10
-20
-30
-40
-10
Dilute Phase
Dense Phase
Rive CatalystIncumbent Catalyst
FIGURE 15: Delta Regenerator Temperatures
28Mar1218Jan12 06Jun12 15Aug12 24Oct12
3.5
Gas
Fact
or 2.5
3.0
2.0
1.5
1.0
Rive CatalystIncumbent Catalyst
FIGURE 10: Ecat Gas Factor
30Jun1231May12 30Jul12 29Aug12 28Oct1228Sep12
14.0
Flue
Gas
Opa
city
,% 10.0
12.0
8.0
4.0
6.0
2.0
0.0
Rive CatalystIncumbent Catalyst
FIGURE 11: Flue Gas Opacity Remained Steady
19
Cat
alys
tCirc
ulat
ion,
TPM
18
17
16
15
14
13
12
11
1016Jul1218Jan12 17Apr12 14Oct12
3.3
Flui
diza
tion
Para
met
er,U
mb/
Um
f
3.1
3.2
2.9
2.7
2.8
2.6
3.0
Rive CatalystIncumbent Catalyst
FIGURE 12: Catalyst Fluidization Parameter andCirculation
28 Issue No. 113 / 2013
As the trial progressed, the combined impact of increased feed rate
from better coke selectivity and improved bottoms cracking, and de-
spite the lowering of riser temperature, the refiner observed a signifi-
cant increase (~ +1,000 BPSD) in gasoline production, and an
increase (~ +400 BPSD) in LCO production (Figures 16a and 16b).
The LPG/gasoline split shifted towards more gasoline (~ -400 BPSD
LPG), and CSO production was constant (~ +50 BPSD).
Gasoline OctaneWith the goal of developing data for an optimized operation on the
Rive catalyst, the riser and regenerator temperatures were lowered
and unit responses observed during the trial. A key response from
lowering the riser temperature in a FCC unit is a drop in gasoline oc-
tane. During the Rive® catalyst trial, the riser temperature was re-
duced by 20°F, and Figure 17 shows the impact of riser temperature
on gasoline octane during normal operation in the Alon, Big Spring
FCC unit.
Octane shows a linear increase with riser temperature. However the
magnitude of the change is smaller than typical. One would expect
a 0.4 (R+M)/2 octane response for a 10°F change in riser tempera-
ture. The observed response with Rive MH-1 catalyst is only 0.15
(R+M)/2; less than half the typical response. One possible explana-
tion is that Rive catalyst produces more olefinic gasoline which par-
tially offsets the effect of a reduction in riser temperature. More
olefinic gasoline has been observed in ACE testing with Rive® zeo-
lite, but Alon FCC gasoline was not tested for this change.
Figure 18 examines the impact of sulfur levels of gasoline feed to
the ultra-low sulfur gasoline unit on the low sulfur gasoline product.
It shows that catalyst type had no discernible impact on the low sul-
fur gasoline product octane.
30Jun1231May12 30Jul12 29Aug12 28Oct1228Sep12
2,000D
elta
Gas
olin
eR
ate,
BPS
D1,600
1,200
800
400
0
-400
-800
Rive CatalystIncumbent Catalyst
1,200
Del
taLC
OR
ate,
BPS
D
1,000
800
600
400
0
-200
200
30Jun1231May12 30Jul12 29Aug12 28Oct1228Sep12
FIGURE 16a & 16b: Rising Gasoline and LCO Rates with Rive® MH-1 Catalyst
Incumbent Catalyst - June ‘12
Rive Catalyst
89.5
FCC
Gas
olin
eO
ctan
e(R
+M)/2 89.0
88.5
88.0
87.5
87.0
86.5
86.0
Delta Riser Top Temperature, ˚F
-5 0 5 10 15 20 25 30
Incumbent Catalyst - July-Oct ‘11
Octane = 0.015 x (Riser Top Temp) + 87.20
FIGURE 17: Impact of Riser Temperature on GasolineOctane
Incumbent Catalyst - June ‘12
Rive CatalystIncumbent Catalyst - July-Oct ‘11
88.0
ULS
GO
ctan
e(R
+M)/2
87.5
87.0
86.5
86.0
85.5
85.0
FCC Gasoline Sulfur, wt.%
0.15 0.17 0.19 0.21 0.23 0.25 0.27 0.29 0.31 0.33
FIGURE 18: Gasoline Sulfur and Low Sulfur GasolineOctane
Grace Catalysts Technologies Catalagram® 29
Butylene YieldsLPG yields and composition are another key response to riser tem-
perature – an increase in riser temperature increases the LPG yield
and olefinicity of the LPG. While the propane/propylene market
value in the US is depressed due to the current excess of natural
gas11 in North America; butylenes are required as alkylation unit
feed. Since alkylate is a prime high-octane gasoline blend compo-
nent it is important to maintain adequate butylene feed to the alkyla-
tion unit from the FCC plant.
Butylenes production rates during the trial were similar to the pre-
trial rates despite operation at a lower riser temperature. Figure 19
provides the impact of riser temperature on butylenes production for
the Rive catalyst and the incumbent catalyst. The data suggests
that at a given riser temperature, the Rive catalyst improved
butylenes production compared with the incumbent catalyst.
Value UpliftThe unit operating data and production rates were heat and mass
balanced with Profimatics. From the yields and selectivities dis-
cussed in the sections above, it is clear that the product rate
changes are a combined result of both increased feed rates and
yield selectivity improvements.
The incremental FCC unit value uplift on a daily basis due to the
change to Rive catalyst is shown in Figure 20. The value uplift was
calculated by applying the April 2012 product pricing to the mass
balanced product rates from the Profimatics model and subtracting
the average FCC value uplift on the incumbent catalyst during the
month prior to the trial. The FCC value uplift progressively in-
creased as the Rive catalyst changed out in the unit.
Trending the value uplift, it is estimated that at the end of the trial the
incremental uplift due to the Rive catalyst was over $2.50/FCC bbl.
Conclusions1. FCC catalyst combining Rive® mesoporous Molecular High
way zeolite and Grace® matrix technology delivered over
$2.50/FCC bbl of additional value in mild resid operation
2. Performance benefits resulting from Rive’s unique zeolite
were substantially improved coke selectivity, improved
bottoms upgrading, and increased olefinicity of the cracked
products
3. No capital investment was needed to realize this value, and
operational changes were within the normal range of
practice.
AcknowledgementsThe authors wish to acknowledge the work and support of the Alon
(Big Spring, TX) refinery staff in making the trial a success. In par-
ticular, Gordon Leaman, Ted Tarbet, Clarence Palmer, Jeff Brorman,
Eric Selden, Manoj Katak and the FCC Operators all made valuable
contributions to the trial.
The author also wishes to acknowledge with sincere appreciation
the entire team from Grace Catalysts Technologies for their collabo-
ration, expertise, and support of the research and development de-
scribed in this paper, and Allen Hansen, Steve McGovern, Ken
Peccatiello, and the other advisors to Rive Technology for their con-
tributions.
$4.00
Incr
emen
talV
alue
Upl
ift,$
/bbl $3.00
$2.00
$1.00
$0.00
-$1.00
-$2.0030Jun1231May12 30Jul12 29Aug12 28Oct1228Sep12
$2.74/bbl
Rive CatalystIncumbent Catalyst
FIGURE 20: Rive Value UpliftIncumbent Catalyst - June ‘12
Rive CatalystIncumbent Catalyst - July-Oct ‘11
2,000
Del
taC
4O
lefin
s,B
PSD
1,500
1,000
500
0
-500
-1,000
Delta Riser Temperature, ˚F-5 0 5 10 15 20 25 30
FIGURE 19: Riser Temperature versus ButylenesProduction
11 Excess due to Shale Gas production
30 Issue No. 113 / 2013
Catalyst13 Incumbent Formulation (MT-4) Rive Formulation (MH-1)
Key Op. Conditions
Riser Temp, °F 1002 9.80
Feed Temp, °F 625 606
C/O, wt/wt 6.31 7.11
Regen Bed Temp, °F 1308 1250
Carbon on Regenerated Catalyst, wt.% 0.17 0.19
Commercial Yields
Dry Gas, wt.% 5.5 4.7
Total C3’s+C4’s, wt.% 14.6 14.6
C3+C4 non-olefins, wt.% 4.3 4.4
C3+C4 olefins, wt.% 10.3 10.2
Cat Gasoline, wt.% 47.9 49.6
LCO, wt.% 18.1 17.9
Btms, wt.% 8.4 7.9
Coke, wt.% 5.4 5.2
Commercial Yields, vol.%
Total C3’s+C4’s, vol.% 24.9 24.9
C3+C4 non-olefins, vol.% 7.4 7.5
C3+C4 Olefins, vol.% 17.5 17.4
Cat Gasoline, vol.% 58.4 60.6
LCO, vol.% 17.1 17.0
Bottoms, vol.% 7.1 6.7
Total Liquid, vol.% Yield: 107.5 109.1
Conversion, vol.% 75.7 76.4
LCO Conversion, vol.% 92.9 93.3
Rive Uplift, $/bbl fcc feed NA $2.00
Product Properties:
Cat Gasoline:
SG 0.749 0.748
RON 92.4 91.2
MON 81.3 80.7
(RON+MON)/2 86.9 85.9
Sulfur, wt.% 0.41 0.40
LCO:
SG 0.966 0.967
Sulfur, wt.% 2.32 2.36
Bottoms:
SG 1.080 1.082
Sulfur, wt.% 3.30 3.36
ATTACHMENT 1: Yields and Operation Predicted12 for Alon (Big Spring, TX)
12 Based on a 100% catalyst change-out
13 Excludes Grace gasoline sulfur reduction additive D-PRISM™
Grace Catalysts Technologies Catalagram® 31
People on the Move
Bob Gatte will assume a new role as VicePresident, Marketing & Technology for Refining
Technologies. Wu-Cheng Cheng, Director, Re-
fining Technologies R&D, will continue to lead
the R&D organization and will report to Bob.
Also reporting directly to Bob will be Tom
Habib, Sudhakar Jale, Rosann Schiller and
Phyl Strawley. Bob will continue reporting directly to Shawn Abrams
Dennis Kowalczyk has been named GeneralManager, Americas for Refining Technologies.
In his new role, Dennis will have overall re-
sponsibility for Americas Sales and Technical
Service. Reporting directly to Dennis will be
Ruben Cruz, Jeff Koebel and Shahab Parva.
Dennis will report directly to Shawn Abrams
Angela Jones has accepted the position ofSenior FCC Technical Sales Representative,
reporting to Dennis Kowalczyk, based in
Houston, TX.
Olivia Topete has joined RT North AmericaMarketing, reporting to Dennis Kowalczyk.
Olivia, is based in Houston, TX.
Mattias Scherer has been named DirectorSales, Northern Europe, RT EMEA with focus-
ing on securing and growing the Northern Eu-
ropean FCC market segment.
Emmanuel Smaragdis will assume theposition of Regional Technical Sales Manager
for Germany, Austria, Greece and the
OMV/Petrom business in Romania, reporting
to Mattias Scherer.
Joanne Deady will assume a new position asVice President, Marketing for Advanced Refin-
ing Technologies. In this role, Joanne will have
global responsibility for ART Marketing, with a
focus on long-term strategy, product strategy,
competitive intelligence and new product de-
velopment. Reporting to Joanne will be
Charles Wear and Ingrid Du. Joanne will report directly to Scott
Purnell.
Eboni Adams has joined ART as Sales Opera-tions Manager, ART. In this role, Eboni will be
responsible for demand forecasting, driving
stronger linkage between sales, customer
service and S&OP in the order fulfillment
process, and leading process improvement op-
portunities within the business.
Grace Catalysts Technologies has announced that it has completed
its acquisition of the assets of Noblestar Catalysts Co., Ltd, a Qing-
dao, China-based manufacturer of fluid catalytic cracking (FCC) cat-
alysts, catalyst intermediates and related products used in the
petroleum refining industry.
Qingdao Bureau of Commerce Vice Director General, Cong Yan,
welcomed Grace’s investment during the ribbon cutting ceremony
and said, “Qingdao is a leading economic center in China. We wel-
come foreign investment, especially from companies like Grace,
which has world-class, leading technologies that can help develop
our fast-growing petrochemical industry while also acknowledging
environmental and safety concerns.”
“The successful acquisition of Noblestar’s assets in Qingdao is an-
other milestone in Grace’s long relationship with China. And it is an
important step in our strategy to provide world-class products and
support to the petroleum refining industry,” said Grace’s Chairman
and CEO Fred Festa. “Our goal is for customers to look to Grace for
innovative technology and industry-leading technical service, as well
as a globally integrated manufacturing network that aligns with the
world’s demand.” Grace expects to make additional investments at
the Qingdao site for environmental, safety and manufacturing up-
grades.
Chao Cui, CEO and President of Noblestar Catalysts, said, “We
have been happy and proud to be a business partner of Grace’s re-
fining technologies business for years and we are excited to con-
tinue a business relationship with Grace in the future.”
Grace first established a presence in China when it founded Grace
China Ltd. in 1986 as the first Wholly Foreign-Owned Company to
do business in the People’s Republic of China through its can
sealants plant in Shanghai.
Currently, Grace operates five manufacturing facilities,three sales
offices and two technical service centers in mainland China, includ-
ing its Asia Pacific regional headquarters in Shanghai.
Grace Acquires Assets in Chinato Serve Refiners in Region
New employees of Grace Catalysts Qingdao atsigning ceremony, November 29, 2012
Grace, Noblestar, and Qingdao officials celebrate Grace’s acquisition of Noblestar
32 Issue No. 113 / 2013
Grace Catalysts Technologies Catalagram® 33
Custom Catalyst Systems forHigher Yields of Diesel
DIESEL
Brian WatkinsManager,Hydrotreating PilotPlant and TechnicalService Engineer
Charles OlsenDirector, Distillate R&Dand Technical Service
AdvancedRefiningTechnologiesChicago, IL, USA
In those areas of the world which are experiencing low costs of natural gas, there has been a decrease in
the cost of hydrogen, and this, combined with the growth in global demand for middle distillates, has
prompted refiners to look to improve profitability by increasing middle distillate yields. Options under consid-
eration have included operating an FCC (Fluid Catalytic Cracker) pretreater in a mild hydrocracking mode,
switching to maximum LCO1 (Light Cycle Oil) mode or extending the endpoint of feed to a ULSD (Ultra Low
Sulfur Diesel) unit and converting the heavy fraction into diesel range material. The use of opportunity feed-
stocks and synthetic type feedstocks can also be considered2. These approaches require specialized cata-
lyst systems capable of providing some cracking conversion or changes to traditional unit operation, and
careful attention must be given to minimizing production of excess gas and naphtha while maximizing
diesel. Another seemingly simple option is to maximize the product volume swell from an existing ULSD
unit through a change in catalyst and understanding the demand on operating conditions. This approach to
increasing diesel yields requires a detailed understanding of feed and operating conditions such that the hy-
drotreater can be operated at the maximum product volume swell for the majority of the unit cycle. In this
case, the benefits of increased diesel yield need to be balanced against the potential costs of increased hy-
drogen consumption and decreased cycle length.
A critical element in all the approaches to increasing diesel yield is the proper design and selection of a cat-
alyst system for the hydrotreater. This paper summarizes some of these various catalytic options and the
operating conditions that can be implemented to increase yields of middle distillate using existing assets
with minimal investment.
As a first step, it is useful to understand the chemistry involved in hydrotreating and, in particular, the chem-
istry required for maximizing product volume swell. Table I lists several different classes of hydrocarbon
compounds that can be found in diesel range feeds. The data shows that as hydrogen is added to a mole-
cule, the density of the compound decreases. This indicates that even some simple reactions involved in
hydrotreating result in a decrease in density of the product or put another way, result in an increase in prod-
uct volume. This is especially apparent with aromatics species.
34 Issue No. 113 / 2013
Table II lists several different aromatic and fully saturated com-
pounds which occur in diesel range feedstock along with some se-
lected properties. It is apparent that dramatic shifts in boiling point
and density can be realized by hydrogenating aromatic compounds.
The density decreases by 20-25% with boiling points shifts any-
where from 50-150°F upon saturation of the aromatic rings.
This suggests that in order to achieve a high degree of product vol-
ume swell in ULSD, a detailed understanding of aromatic and
polynuclear aromatic (PNA) hydrogenation is required. It is well un-
derstood that hydrogenation of aromatic compounds is a reversible
reaction, and that the equilibrium conversion is less than 100%
under typical conditions. The equilibrium conversion is highly de-
pendent on temperature and hydrogen partial pressure. Figure 1
shows how the saturation of aromatics in diesel changes with H2
partial pressure at a typical temperature for ULSD. The base pres-
sure is around 500 psi, so the data cover the range of H2 pressures
typically encountered in ULSD. The total aromatics conversion
nearly doubles with a 2.5 times increase in H2 partial pressure.
Figure 2 shows how the aromatics conversion changes with temper-
ature in a typical ULSD unit. The figure compares the conversion
observed for both a NiMo and a CoMo catalyst. The data clearly in-
dicates that the NiMo catalyst has the greater aromatic saturation
activity of the two catalysts shown. The product aromatics concen-
tration is over 4% (absolute) lower for the NiMo catalyst compared to
the CoMo catalyst. This difference in aromatics conversion ac-
counts for the higher H2 consumption typically seen for a NiMo com-
pared to a CoMo catalyst. The chart also shows the influence of
equilibrium on aromatics conversion. As the temperature increases
beyond about 670-680°F the conversion actually begins to decrease
as the rate of the dehydrogenation reactions has increased enough
Class Compound Formula Density, g/cc ˚API H/C Ratio
Iso Paraffin 2,3-dimethyl-octane C10H22 0.738 60.3 2.2
Paraffin n-decane C10H22 0.730 62.3 2.2
Olefin 1-decene C10H20 0.741 59.5 2.0
Naphthene Decalin C10H18 0.897 26.3 1.8
Mono Aromatic Tetralin C10H12 0.970 14.3 1.2
Poly Aromatic Naphthalene C10H8 0.738 -7.4 0.8
TABLE I: Selected Compounds Boiling in the Diesel Range
Aromatics Saturates
Rings Compound Formula Density, g/cc Boiling Point, ˚F Compound Formula Density, g/cc Boiling Point, ˚F
2 Naphthalyne C10H8 1.140 424 Decaline C10H8 0.897 374
3 Fluorene C13H10 1.202 563 Perhydro Fluorene C13H22 0.920 487
3 Phenanthrene C14H10 1.180 630 PerhydroPhenanthrene C14H24 0.944 518
4 Pyrene C16H10 1.271 759 Perhydro Pyrene C16H26 0.962 604
TABLE II: Aromatic Compounds Found in Diesel Range Feed
80.0
Aro
mat
ics
Con
vers
ion,
wt.% 70.0
75.0
65.0
60.0
55.0
50.0
45.0
40.0
35.0
30.0
Constant Temperature
1.00.5 1.5 2.5 3.02.0
H2 Pressure/Base H2 Pressure
FIGURE 1: Effect of Pressure on AromaticsHydrogenation
29.0
Tota
lAro
mat
ics,
vol.%
25.0
27.0
23.0
21.0
19.0
17.0
15.0620600 640 680 700 720 740660
Reactor Outlet Temperature, ˚F
CoMo NiMo
FIGURE 2: Aromatic Reduction in ULSD
Grace Catalysts Technologies Catalagram® 35
to compete with saturation reactions. At high enough temperatures
both catalysts give the same conversion since they are operating in
an equilibrium-controlled regime.
One significant consequence of achieving a high level of saturation
of multi-ring and mono-aromatic ring compounds is higher hydrogen
consumption. However, not all aromatic species are created equal
when it comes to hydrogen consumption. Figure 3 shows a simple
schematic of the reaction pathway for saturating a 4-ring poly aro-
matic compound. The hydrogenation occurs in a stepwise fashion
where one aromatic ring at a time is being saturated, with each step
along the pathway being subject to equilibrium constraints. The rate
limiting step to the fully saturated species is hydrogenation of the
last aromatic ring (the mono aromatic), and this step consumes the
most hydrogen of the reactions shown in the reaction pathway.
Three moles of hydrogen are required to hydrogenate the mono-
ringed compound compared to two moles of hydrogen to hydro-
genate the rings in the poly aromatic compounds.
A number of poly aromatic species have been studied over the
years leading to a good understanding of the chemistry involved in
PNA saturation.3 In the case of naphthalene, the reaction begins
with the hydrogenation of one of the aromatic rings to form tetralin, a
mono-ring aromatic. The next reaction is hydrogenation of the re-
maining aromatic ring to produce decalin, the fully saturated
species. The reactions occur sequentially with the rate of hydro-
genation of the final aromatic ring an order of magnitude lower than
saturation of the first aromatic ring. The reactions can be modeled
as a series of first order reversible reactions. Figure 4 shows the
species concentration profiles as a function of residence time for a
hydrogenation reaction sequence such as that for naphthalene just
discussed. The rate of the first hydrogenation reaction in the series
is an order of magnitude faster than the rate of the second hydro-
genation reaction. There is a rapid decrease in the concentration of
the 2-ringed aromatic species at short residence times and a corre-
sponding increase in the mono-ringed species. As contact time in-
creases however, the mono-ring aromatic concentration begins to
decrease and the fully saturated species begin to build up. This
type of concentration profile suggests that there is a range of resi-
dence times in the unit corresponding to a maximum in the mono-
ringed aromatic concentration.
A variety of substituted naphthalene’s have also been shown to fol-
low a similar reaction network with the rate of hydrogenation of the
first aromatic ring approximately equal to that observed for naphtha-
lene. The hydrogenation of biphenyl occurs in a stepwise fashion as
well, with the rate of hydrogenation of the first aromatic ring about an
order of magnitude faster than that of the mono ring compound. An
interesting difference is that the rate of the first hydrogenation reac-
tion in naphthalene is approximately an order of magnitude faster
+3 H2
~+2 H2
~+2 H2
~+2 H2
“Naphthene”
Mono-Aromatic
Tri-AromaticPoly-Nuclear Aromatic
Di-Aromatic
FIGURE 3: Stepwise Saturation of a Poly AromaticCompound
k1-10*k21.00
Con
cent
ratio
n
0.80
0.90
0.70
0.60
0.50
0.40
0.30
0.20
0.10
0.00Residence Time
Monorings Polyrings Sat’d
FIGURE 4: 1st Order Reversible Reactions in SeriesConcentration Profile
1
1
20
16
FIGURE 5: Relative Rate Constants for SaturatingAromatics
36 Issue No. 113 / 2013
than the rate of hydrogenation of the first ring in biphenyl. Figure 54
compares relative reactions rates for selected aromatics species.
Figure 6 summarizes pilot plant data demonstrating how the aro-
matic species change in ULSD product as a function of the resi-
dence time (i.e. 1/LHSV) (Liquid Hourly Space Velocity) in the
reactor. Notice how the curves look very similar to the simple exam-
ple discussed in Figure 4. For PNA saturation, the 2-ringed aro-
matic going to the mono ring aromatic, there is a fairly steep decline
in concentration as a function of residence time below about 0.5 hr.
Above that point, which represents space velocities of 2 hr-1 or less,
there is very little change due to equilibrium constraints. For mono-
ringed aromatic saturation there is a steady increase in conversion
as the residence time is increased, and eventually the mono-ringed
concentration begins to decrease indicating that mono-ring satura-
tion gets a lot more favorable as the LHSV is decreased. These
data show that PNA saturation occurs fairly readily under typical hy-
drotreating conditions, but saturation of mono rings aromatics is
much more difficult and is aided by lower LHSV.
Hydrotreaters with very short residence time (high LHSV) will have
difficulty achieving higher volume swells due to the much slower rate
of saturating the final aromatic ring. These units will require a higher
temperature in order to drive the kinetic saturation portion of the re-
action. This can have some negative effects on catalyst perform-
ance by decreasing the expected cycle time due to the higher start
of run temperature and the increased fouling rate associated with it.
ART (Advanced Refining Technologies) was interested in exploring
aromatics saturation and the impact of product volume further, and
completed some pilot plant work for a refiner. The feedstock used
for this case study contained 50% cracked material, and the operat-
ing conditions included 850 psi hydrogen pressure and a H2/Oil ratio
over four times the hydrogen consumption.
Figure 7 summarizes the HDS and aromatics conversion observed
for the CoMo catalyst in that test. A temperature of 665°F was re-
quired to achieve 10 ppm sulfur in this case. At that temperature
about 36% aromatics hydrogenation was achieved which is less
than the maximum possible aromatic saturation for these conditions.
The maximum aromatic saturation in this case is about 42% at just
under 700°F as shown on the chart.
Figure 8 shows results for a NiMo catalyst on the same feed and
conditions. In this case just over 640°F is required to achieve 10
ppm product sulfur and, at that temperature, about 38% aromatics
conversion is achieved. Comparing with the data in Figure 7 it is ap-
parent that the NiMo catalyst is significantly more active for HDS
than the CoMo catalyst, and it achieves slightly higher aromatics sat-
uration when running to make 10 ppm product sulfur despite running
at a lower temperature.
Comparing the catalysts in maximum aromatic saturation mode re-
veals significantly larger differences between catalysts. Maximum
aromatics conversion occurs at 685°F for the NiMo catalyst and, at
that temperature, the aromatics conversion is 52%. The NiMo cata-
lyst is achieving over 10 numbers higher aromatics conversion than
the CoMo catalyst.
Figures 9 and 10 summarize the same data, but now show the im-
pact on product volume. The yields for the CoMo catalyst system
are shown in Figure 9 along with the product sulfur. In this case the
difference in yields from operating in ULSD mode versus a maxi-
mum volume swell mode is very low. The difference in aromatics
Aro
mat
icC
onte
nt40
35
30
25
20
15
10
5
00 0.2 0.4 0.6 0.8 1
1/LHSV, hrs
Mono Rings Poly Rings Saturates
FIGURE 6: Aromatic Concentrations in ULSD
Sulfur Total18
Prod
uctS
ulfu
r,w
ppm
1416
121086420640 680 700 720 740660
Temperature, ˚F
55%
Aro
mat
icC
onve
rsio
n,vo
l.%
45%
50%
40%
35%
30%
25%
FIGURE 7: Comparison of HDS and AromaticSaturation Using CoMo
Sulfur Total18
Prod
uctS
ulfu
r,w
ppm
1416
121086420
640 680 700 720 740660
Temperature, ˚F
620
55%
Aro
mat
icC
onve
rsio
n,vo
l.%
45%
50%
40%
35%
30%
25%
FIGURE 8: Comparison of HDS and AromaticSaturation Using NiMo
Grace Catalysts Technologies Catalagram® 37
conversion for ULSD and maximum aromatics is not large enough to
result in any significant change in product volume. There is no eco-
nomic incentive to run for maximum volume swell with this system.
The situation is different for the NiMo catalyst as shown in Figure 10.
Operating in ULSD mode results in estimated distillate yields that
are about 1% higher compared to the CoMo catalyst. Of course this
comes at the cost of additional hydrogen consumption with the NiMo
catalyst and assumes that the extra hydrogen required for stable op-
eration is readily available at a reasonable cost. The figure also
highlights the yields for running to maximum aromatic saturation.
Running the unit for maximum volume swell requires an increase in
temperature to around 670°F. At this temperature there is over 1.0%
additional volume gain which also results in 40-60 SCFB (standard
cubic feet per barrel) additional hydrogen consumption.
When estimating the benefits of operating a hydrotreater in maxi-
mum saturation mode vs. simply maintaining ULSD it is also impor-
tant to realize that the entire cycle is not expected to produce the
additional volume swell. Figure 11 shows the results of modeling the
differences in ULSD temperature profiles during the cycle for the two
operating strategies. In ULSD mode the reactor temperature is in-
creased to maintain a constant product sulfur of 10 wppm. The end
of run (EOR) is typically determined by a maximum outlet tempera-
ture and often this is the point when the product color is out of speci-
fication. In this case, EOR is reached in about 59 months. In
Sulfur Yield, vol.%
640 680 700 720 740660
Temperature, ˚F
18
Prod
uctS
ulfu
r,w
ppm
1416
121086420
102.8
Yiel
ds,v
ol.%
102.4
102.6
102.2
102.0
101.8
101.6
101.4
101.2
FIGURE 9: Distillate Yields Using a CoMo CatalystSystem
18
Prod
uctS
ulfu
r,w
ppm
1416
1210
86420
105.0
Yiel
ds,v
ol.%104.0
104.5
103.5
103.0
102.5
102.0640 680 700 720 740660
Temperature, ˚F
620
FIGURE 10: Distillate Yields Using a NiMo CatalystSystem
EOR for ULSD740
Tem
pera
ture
,˚F 700
720
680
660
640
620
600100 20 40 50 60 7030
Cycle Length, Months
HDS - NiMoHDPNA - NiMo
EndofA
SATM
ode
FIGURE 11: Maximum Saturation Versus ULSD ModeComparison
switching to maximum saturation mode the reactor temperature is
ramped up to the conditions resulting in maximum PNA/HDA aro-
matic conversion. The temperature is then adjusted to maintain a
constant saturation level. PNA/HDA saturation activity deactivates
at a slower rate relative to HDS activity, so the rate of temperature
increase in PNAmode is much slower than for the HDS mode. The
EOR for the PNAmode of operation is determined by the required
sulfur level of 10 wppm, at which point the ULSD unit switches to
maintaining the 10 wppm product sulfur until the EOR temperatures
are met.
The fact that saturation activity deactivates at a slower rate than
HDS activity is validated by comparing commercial operating data.
API upgrade is often used as a simple measure of aromatics satura-
tion and can be tracked through the cycle. Figure 12 summarizes
commercial data for API upgrade from eight different units ranging in
operating pressure from 615-1900 Psig and 0.77-3.7 LHSV. The
data indicate that the API upgrade is maintained throughout the
cycle in these cases.
ART next examined the value of operating in a maximum volume
swell mode vs. ULSD mode. For this model it was assumed that the
unit processes 50,000 barrels per day, and the data from Figure 10
is used to estimate the yields. Data, like that from Figure 12, is used
to determine the cycle length expected for operating in either the
HDS (ULSD mode) or maximum yield mode. Based on the under-
standing from Figure 11 that the expected run length will be the
same for either mode, costs such as turnaround costs and operating
costs will be equal and will not need to be applied in determining the
financial impact of operating in ASAT mode. For this example the
catalyst systems are also identical, so that cost of the catalyst would
not need to be included in the financial evaluation, but when consid-
ering catalyst changes this cost would be included. Therefore, the
only difference between these two modes of operation are the addi-
tional barrels of product produced by operating in ASAT mode, and
38 Issue No. 113 / 2013
the incremental hydrogen required to do so. For this financial analy-
sis the cost of the feed to the ULSD unit is assumed to be $5 lower
than the cost of the product being sold.
With a detailed look at only the first 50 months of the cycle where
the two modes of operation have different yield structures and hy-
drogen usages, Figure 13 shows that the HDS mode produces
nearly 79 million barrels of product, while the ASAT mode produces
almost 80 million barrels of product. The barrels produced in the last
9 months of the run are not included in this total, as both operating
modes will need to finish in HDS operation in order to simply meet
the ULSD product targets and maximum run length.
As was stated earlier, the difference between the two modes of oper-
ation is the use of hydrogen in order to produce the additional prod-
Del
taA
PI
14
12
10
8
6
4
2
00
Days on Stream200 400 600 800 1000 1200 1400
Refiner A Refiner B Refiner C Refiner DRefiner E Refiner F Refiner G Refiner H
FIGURE 11: Commercial Delta API over a ULSD Cycle
HDS Mode
80.0
Mill
ion
Bar
rels 79.6
79.8
79.479.279.078.878.678.4
ASAT Mode
FIGURE 12: Commercial Delta API Over a ULSD
HDS Mode
45,000
MM
SCF
Hyd
roge
n 43,00044,000
42,00041,00040,00039,00038,00037,00036,00035,000
ASAT Mode
FIGURE 13: Barrels Produced in HDS or ASAT Mode
uct barrels. There is almost 80 SCFB standard cubic feet per barrel
in additional hydrogen consumption to produce those barrels which,
over the 50-month cycle, amounts to over 5 million cubic feet of in-
cremental hydrogen consumed as shown in Figure 14. Using a hy-
drogen value of $3.00 per 1000 scf, the incremental hydrogen
consumed amounts to a cost of just over $16 million dollars, or
about $0.22 per barrel more than operating in HDS mode.
However, the revenue from sale of the additional product barrels
produced in ASAT mode is more than sufficient to cover the cost of
incremental hydrogen consumed. The net impact for this mode of
operation is a $1.20 per barrel premium for operating in ASAT mode
versus HDS mode for the first 50 months of the cycle.
ART has the ability to conduct detailed customer specific pilot plant
testing to provide the refiner the confidence and understanding of
Grace Catalysts Technologies Catalagram® 39
the various options available when considering a catalyst change.
Numerous refiners have chosen to place ART catalyst into their
ULSD hydrotreater in order to achieve the optimization between
ULSD and maximum yield ULSD.
Both the hydrotreating catalyst system and the operating strategy for
the ULSD unit are critical to providing the highest quality products.
Driving the hydrotreater to remove sulfur and PNA's improves prod-
uct value, but this needs to be balanced against the increased costs
of higher hydrogen consumption. Use of tailored catalyst systems
can optimize the ULSD hydrotreater in order to produce higher vol-
umes of high quality products while balancing the refiners available
hydrogen.
The complex relationship between hydrotreater operation and cata-
lyst kinetics underscores the importance of working with a catalyst
technology supplier that can tailor product offerings for each refiner’s
unique operating conditions. This knowledge enables ART to meet
the refiner’s objectives and maximize revenue.
References1. Watkins, B., Olsen, C., Hunt, D., NPRAAnnual Meeting, Paper
AM 11-21, Balancing the Need for Low Sulfur FCC Products and In-
creasing FCC LCO Yields by Applying Advanced Technology for Cat
Feed Hydrotreating
2. Olsen, C., Watkins, B., 2009 NPRAAnnual Meeting, Paper AM
09-78, Distillate Pool Maximization by Exploiting the use of Opportu-
nity Feedstock’s Such as LCO and Synthetic Crudes.
3. Olsen, C., D’Angelo, G., 2006 NPRAAnnual Meeting, Paper
AM-60-06, No Need to Trade ULSD Catalyst Performance for Hy-
drogen Limits: SmART Approaches
4. Girgis, M.J., Gates, B.C., Ind. Eng. Chem. Res., 30, 1991, p
2021
40 Issue No. 113 / 201340 Issue No. 113 / 2013
ART Announces CLGHydrocracking CatalystsSales Agreement
COLUMBIA, Md. - February 28, 2013Advanced Refining Technologies LLC (ART) announced that it has
signed an agreement with Chevron Lummus Global (CLG) regarding
hydrocracking and lubes hydroprocessing catalysts. Under this
agreement, ART will have the exclusive right to sell CLG's hydroc-
racking and lubes hydroprocessing catalysts to CLG's licensees and
other petroleum refiners for unit refills. The agreement will stream-
line hydroprocessing catalyst supply and improve technical service
for refining customers by establishing ART as the single point of con-
tact for all their hydroprocessing catalyst needs.
ART is a joint venture between subsidiaries of W. R. Grace & Co.
and Chevron Corporation. CLG is a joint venture between a sub-
sidiary of Chevron and CB&I’s Lummus Technology group.
ART is a leading supplier of hydroprocessing catalysts, with a portfo-
lio of distillate hydrotreating, fixed bed resid hydrotreating, and ebul-
lated bed resid hydrocracking catalysts. CLG is a world leader in
hydroprocessing technology development and commercialization,
with licensing, engineering, and petroleum refining expertise. Its
portfolio includes hydrocracking (ISOCRACKING®), lubes hydropro-
cessing (ISODEWAXING® and ISOFINISHING®), ebullating bed
resid hydrocracking (LC-FINING®), and hydrotreating (ISOTREAT-
ING®) technologies.
Scott Purnell, managing director of ART, commented, "We are
pleased to add hydrocracking and lubes hydroprocessing catalysts
to our current product portfolio. CLG's ISOCRACKING®, ISOTREAT-
ING®, ISODEWAXING®, and ISOFINISHING® catalysts are proven
products that will help our refining customers improve quality and
yield. With this new agreement, all of our customers’ hydroprocess-
ing catalyst needs can be provided through a single point of con-
tact."
Leon de Bruyn, managing director of CLG, added, "We continually
invest to provide our licensees with world-class process technology,
catalysts and support services. This agreement represents a unique
combination of ART's well-established portfolio of hydrotreating cata-
lysts, extensive sales network and manufacturing expertise, together
with our hydrocracking and lubes hydroprocessing catalyst technolo-
gies, and engineering and technical know-how. It will allow our cus-
tomers to receive broader service and more advanced catalyst
materials, and will improve the competitiveness and profitability of
their refineries.”
Under the agreement, ART will be the worldwide provider for hydroc-
racking and lubes hydroprocessing catalysts. CLG will continue to
focus on its world-class technology development, licensing, design,
and revamp of hydrocracking, lubricant base oil, resid hydrotreating,
and resid hydrocracking plants globally. Both ART and CLG cus-
tomers will continue to have access to the broad depth of Chevron
technical service and hydroprocessing operating expertise.
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