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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


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)


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


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.%


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.%


FIGURE 4: Effect of Conversion Level on LCO Yield and Quality for Straight Run Shale Oil


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


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








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.


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



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







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


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

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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,

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6. D. Dharia, W. Letzsch, L. Chapin, “Advanced Catalytic Crack-

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Riser Outlet Temperature = 1050°F Riser Outlet Temperature = 1100°F



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








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


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.

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“Best Practices In Process Development Of New Technology,”

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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


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


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


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


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


FIGURE 5: Regenerator Temperatures


80

MAT

Act

ivity

,wt.%

Con

vers

ion

76

78

74

72

70

68


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.


120Ze

olite

Surf

ace

Are

a,m

2 /gm

110

115

105

100

95

90


FIGURE 7: Ecat Zeolite Surface Area

60

Mat

rixSu

rfac

eA

rea,

m2 /g

m

50

55

45

40

35



FIGURE 8: Ecat Matrix/Mesopore Surface Area


2.0

Cok

eFa

ctor

1.6

1.8

1.4

1.2

1.0

0.8


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


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


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


FIGURE 15: Delta Regenerator Temperatures


3.5

Gas

Fact

or 2.5

3.0

2.0

1.5

1.0


FIGURE 10: Ecat Gas Factor


14.0

Flue

Gas

Opa

city

,% 10.0

12.0

8.0

4.0

6.0

2.0

0.0


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


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.


2,000D

elta

Gas

olin

eR

ate,

BPS

D1,600

1,200

800

400

0

-400

-800


1,200

Del

taLC

OR

ate,

BPS

D

1,000

800

600

400

0

-200

200


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


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


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™


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


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


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


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


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.

When you can’t predictwhat’s in the pipeline

THINK MIDAS®

FCCCATALYST

TO CAPTURE THE VALUE OFOPPORTUNITY CRUDES

grace.com

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