Fluid Catalytic Cracking Handbook SECOND EDITION
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Fluid Catalytic
Cracking
Handbook
HandbookDesign,Operation and Troubleshooting of FCC Facilities
SECOND EDITION
GP Gulf Professional Publishing I'M
an imprint of Butterworth-Heinemann
uid atalytic racking andbook
gn, Operation, and bleshooting of Facilities
OND EDITION
yright © 2000 by Butterworth-Heinemann. All rights rved. Printed in the United States of America. This book, arts thereof, may not be reproduced in any form without mission of the publisher.
ginally published by Gulf Publishing Company, ston. TX.
information, please contact: nager of Special Sales erworth-Heinemann Wildwood Avenue bum,MA01801–2041 781-904-2500 :781-904-2620 information on all Butterworth-Heinemann publications lable, contact our World Wide Web home page at: ://www.bh.com 9
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ary of Congress Cataloging-in-Publication Data
ghbeigi, Reza. Fluid catalytic cracking handbook / Reza Sadeghbeigi.—2nd ed. p. cm. Includes bibliographical references and index. ISBN 0-88415-289-8 (alk. paper) 1. Catalytic cracking. 1. Title. P690.4.S23 2000 65.533 dc2l 00-035361
ted in the United States of America.
ted on acid-free paper (°°).
This book is dedicated to our respected clients who have contributed to the success of RMS Engineering, Inc. and to the content of this book
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Contents
cknowledgments
xi
reface to the Second Edition
xii
HAPTER 1
rocess Description
1
Feed Preheat, 6. Riser—Reactor—Stripper, 7, RegeneratorHeat/Catalyst Recovery, 13. Main Fractionator, 22. Gas Plant, 25. Treating Facilities, 31. Summary, 39. References, 39.
CC Feed Characterization
40
Hydrocarbon Classification, 41. Feedstock Physical Properties, 45. Impurities, 54. Empirical Correlations, 68. Benefits of Hydroprocessing, 81. Summary, 82. References, 82.
HAPTER 3
CC Catalysts
84
Catalyst Components, 84. Catalyst Manufacturing Techniques, 96. Fresh Catalyst Properties, 99. Equilibrium Catalyst Analysis, 102. Catalyst Management, 109. Catalyst Evaluation. 115. Additives, 117. Summary, 123. References, 124.
HAPTER 4
Chemistry of FCC Reactions Thermal Cracking, 126. Catalytic Cracking, 128. Thermodynamic Aspects, 136. Summary, 136. References, 138 References, 134.
125
HAPTER 5
Unit Monitoring and Control
_.._ 139
Material Balance, 140. Heat Balance, 158. Pressure Balance, 166. Process Control Instrumentation, 177. Summary, 180. References, 181.
HAPTER 6
Products and Economics
182
FCC Products, 182. FCC Economics, 202. Summary, 205. References, 205.
HAPTER 7
Project Management and Hardware Design
206
Project Management Aspects of an FCC Revamp, 206. Process and Mechanical Design Guidelines, 212. Summary, 232. References, 232.
HAPTER 8
roubleshooting
234
Guidelines for Effective Troubleshooting, 235. Catalyst Circulation, 236. Catalyst Losses, 244. Coking/Fouling, 248. Flow Reversal, 251. High Regenerator Temperature, 256. Increase in Afterburn, 259. Hydrogen Blistering, 260. Hot Gas Expanders, 263. Product Quantity and Quality, 264. Summary, 275.
HAPTER 9
Debottlenecking and Optimization Introduction, 276. Approach to Debottlenecking, 277. Reactor/Regenerator Structure, 281. Flue Gas System, 296. FCC Catalyst, 296. Instrumentation, 304. Utilities/Offsites, 305. Summary, 306.
276
Emerging Trends in Fluidized Catalytic Cracking _
307
Reformulated Fuels, 308. Residual Fluidized Catalytic Cracking (RFCC), 323. Reducing FCC Emissions, 327. Emerging Developments in Catalysts, Processes, and Hardware, 232. Summary, 335. References, 336.
PPENDIX 1
emperature Variation of Liquid Viscosity
338
PPENDIX 2
Correction to Volumetric Average Boiling Point
339
PPENDIX 3
OTAL Correlations
340
PPENDIX 4
-d-M Correlations _.._._ ._.._ __._
_
341
PPENDIX 5
Estimation of Molecular Weight of Petroleum Oils from Viscosity Measurements
342
PPENDIX 6
Kinematic Viscosity to aybolt Universal Viscosity _._
_
344
PPENDIX 7
API Correlations
345
PPENDIX 8
Definitions of Fluidization Terms
_._..._..._ _
Conversion of ASTM 50% Point o TBP 50% Point Temperature
_
347
350
PPENDIX 10
Determination of TBP Cut Points from ASTM D-86
351
PPENDIX 11
Nominal Pipe Sizes
Conversion Factors __
Glossary
ndex
__._ _
..._... ....._.
353 355 357
_._ _
About the Author
_
363 ..__ _
._
.__. 369
Acknowledgments
am grateful to the following individuals who played key roles in this ook's completion: Warren Letzsch of Stone & Webster Engineerng Corporation; Terry Reid of Akzo Nobel Chemicals, Inc.; Herb elidetzki of KBC Advanced Technologies, Inc.; and Jack Olesen of Grace/Davison provided valuable input. My colleagues at RMS ngineering, especially Shari Gauldin, Larry Gammon, and Walt Broad went the "extra mile" to ensure the book's accuracy and usefulness.
Preface to the
Second Edition
The first edition of this book was published nearly five years ago. he book was well received and the positive reviews were overwhelming. My main objective of writing this second edition is to rovide a practical "transfer of experience" to the readers of the nowledge that I have gained in more than 20 years of dealing with arious aspects of the cat cracking process. This second edition fulfills my goal of discussing issues related to he FCC process and provides practical and proven recommendations o improve the performance and reliability of the FCCU operations. he new chapter (Chapter 9) offers several "no-to-low" cost modificaons that, once implemented., will allow debottlenecking and optimizaon of the cat cracker. I am proud of this second edition. For one, I received input/feedback rom our valued clients, industry "FCC gurus," as well as my colleagues t RMS Engineering, Inc. Each chapter was reviewed carefully for ccuracy and completeness. In several areas, I have provided additional iscussions to cover different FCCU configurations and finally, both he metric and English units have been used to make it easier for eaders who use the metric system. Unfortunately, the future of developing new technologies for petroeum refining in general, and cat cracking in particular, is not promisng. The large, multinational oil companies have just about abandoned heir refining R&D programs. The refining industry is shrinking apidly. There is no "farm system" to replace the current crop of echnology experts. In cat cracking, we begin to see convergence and imilarity in the number of offered technologies. Even the FCC atalyst suppliers and technology licensers have been relatively quiet n developing "breakthrough" technologies since the introduction of
eolite in the late 1960s. More and more companies are outsourcing heir technical needs. In the next several years, refiners will be pending much of their capital to reduce sulfur in gasoline and diesel, n the area of cat cracking, the emphasis will be on improving the erformance and reliability of existing units, as well as "squeezing" more feed rate and/or conversion without capital expenditure. In light these developments, this book is needed more than ever. Reza Houston, Texas
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CHAPTER
Process Description
Fluid catalytic cracking (FCC) continues to play a key role in an ntegrated refinery as the primary conversion process. For many efiners, the cat cracker is the key to profitability in that the successful peration of the unit determines whether or not the refiner can remain ompetitive in today's market. Approximately 350 cat crackers are operating worldwide, with a otal processing capacity of over 12.7 million barrels per day [1]. Most f the existing FCC units have been designed or modified by six major chnology licensers: 1. ABB Lummus Global 2. Exxon Research and Engineering (ER&E) 3. Kellogg Brown & Root—KBR (formerly The M.W. Kellogg Company) 4. Shell Oil Company 5. Stone & Webster Engineering Corporation (SWEC)/IFP 6. UOP (Universal Oil Products)
Figures 1-1 through 1-3 contain sketches of typical unit configuraons offered by some licensers. Although the mechanical configuration f individual FCC units may differ, their common objective is to pgrade low-value feedstock to more valuable products. Worldwide, bout 45% of all gasoline comes from FCC and ancillary units, such s the alkylation unit. Since the start-up of the first commercial FCC unit in 1942, many mprovements have been made. These improvements have enhanced e unit's mechanical reliability and its ability to crack heavier, loweralue feedstocks. The FCC has a remarkable history of adapting to ontinual changes in market demands. Table 1-1 shows major developents in the history of the process. The FCC unit uses a microspheroidal catalyst, which behaves like liquid when properly aerated by gas. The main purpose of the unit
Fluid Catalytic Cracking Handbook Products Regen Flue Gas
Transfer Line Reactor
Air Blower
Figure 1-1.
Typical schematic of Exxon's flexicracker,
to convert high-boiling petroleum fractions called gas oil to highalue, high-octane gasoline and heating oil. Gas oil is the portion of rude oil that commonly boils in the 650+°F to 1,050+°F (330° to 50°C) range. Feedstock properties are discussed in Chapter 2. Before proceeding, it is helpful to examine how a typical cat cracker ts into the refinery process. A petroleum refinery is composed of everal processing units that convert raw crude oil into usable products uch as gasoline, diesel, and jet fuel (Figure 1-4). The crude unit is the first unit in the refining process. Here, the aw crude is distilled into several intermediate products: naphtha, erosene, diesel, and gas oil. The heaviest portion of the crude oil, (text continued on page 6)
Process Description Flue Gas
To Fractionator
Reactor
^Stripping Steam
Figure 1-2.
UOP FCC (courtesy of UOP).
Second stage regenerator Riser termination device
Combustion Air First stage regenerator
Combustion Air r
Feed Injection
Lift air
Figure 1-3. SWEC stacked FCC unit (courtesy of Stone & Webster Engieering Corporation),
Fluid Catalytic Cracking Handbook
915
936
938
942
943
947
948
950s 951 952 954 Mid-50s 956 961
964
972 974 975 981 983
985 994
996
Table 1-1 The Evolution of FCC
McAfee of Gulf Refining Co. discovered that a Friedel-Crafts aluminum chloride catalyst could catalytically crack heavy oil. Use of natural clays as catalyst greatly improved cracking efficiency. Catalyst Research Associates (CRA) was formed. The original CRA members were: Standard of New Jersey (Exxon), Standard of Indiana (Amoco), Anglo Iranian Oil Company (BP Oil), The Texas Company (Texaco), Royal Dutch Shell, Universal Oil Products (UOP), The M.W, Kellogg Company, and I.G. Farben (dropped in 1940). First commercial FCC unit (Model I upflow design) started up at Standard of New Jersey's Baton Rouge, Louisiana, refinery. First down-flow design FCC unit was brought on-line. First thermal catalytic cracking (TCC) brought on-line. First UOP stacked FCC unit was built. Kellogg introduced the Model III FCC unit. Davison Division of W.R. Grace & Co. developed microspheroidal FCC catalyst. Evolution of bed-cracking process designs. M.W. Kellogg introduced the Orthoflow design. Exxon introduced the Model IV. High alumina (A12 O2) catalysts were introduced. UOP introduces side-by-side design. Shell invented riser cracking. Kellogg and Phillips developed and put the first resid cracker onstream at Borger, Texas. Mobil Oil developed USY and ReY FCC catalyst. Last TCC unit completed. Amoco Oil invented high-temperature regeneration. Mobil Oil introduced CO promoter. Phillips Petroleum developed antimony for nickel passivation. TOTAL invented two-stage regeneration for processing residue, Mobil reported first commercial use of ZSM-5 octane/olefins additive in FCC Mobil started installing closed cyclone systems in its FCC units. Coastal Corporation conducted commercial test of ultrashort residence time, selective cracking. ABB Lummus Global acquired Texaco FCC technologies.
GASOLINE
f •XL
TAR
!j§
I
DELAYED COKER
HEATING OIL
DECANT OIL 3AS( gasoline to
REFORMER
Figure 1-4. A typical high conversion refinery.
NO. 6 OIL
Fluid Catalytic Cracking Handbook
ext continued from page 2)
which cannot be distilled in the atmospheric tower, is heated and sent o the vacuum tower where it is split into gas oil and tar. The tar from he vacuum tower is sent to be further processed in a delayed coker, easphalting unit, or visbreaker, or is sold as fuel oil or road asphalt. The gas oil feed for the conventional cat cracker comes primarily om the atmospheric column, the vacuum tower, and the delayed oker. In addition, a number of refiners blend some atmospheric or acuum resid into the feedstocks to be processed in the FCC unit. The FCC process is very complex. For clarity, the process descripon has been broken down into six separate sections: • • « • • «
Feed Preheat Riser—Reactor—Stripper Regenerator—Heat/Catalyst Recovery Main Fractionator Gas Plant Treating Facilities
EED PREHEAT
Most refineries produce sufficient gas oil to meet the cat crackers' emand. However, in those refineries in which the gas oil produced oes not meet the cat cracker capacity, it may be economical to upplement feed by purchasing FCC feedstocks or blending some esidue. The refinery-produced gas oil and any supplemental FCC eedstocks are generally combined and sent to a surge dram, which rovides a steady flow of feed to the charge pumps. This drum can lso separate any water or vapor that may be in the feedstocks. From the surge drum, the feed is normally heated to a temperature f 500°F to 700°F (260°C to 370°C). The main fractionator bottoms umparound and/or fired heaters are the usual sources of heat. The eed is first routed through heat exchangers using hot streams from he main fractionator. The main fractionator top pumparound, light ycle oil product, and bottoms pumparound are commonly used (Figre 1-5). Removing heat from the main fractionator is at least as mportant as preheating the feed. Most FCC units use fired heaters for FCC feed final preheat. The eed preheater provides control over the catalyst-to-oil ratio, a key ariable in the process. In units where the air blower is constrained.
Process Description Vent to Main Column
cSi—^
Slurry
I
Feed Preheater Figure 1-5. Typical feed preheat system.
ncreasing preheat temperature allows increased throughput. The effects f feed preheat are discussed in Chapter 6.
RISER—REACTOR—STRIPPER
The reactor-regenerator is the heart of the FCC process. In a modern at cracker, virtually all the reactions occur in 1.5 to 3.0 seconds before the catalyst and the products are separated in the reactor. From the preheater, the feed enters the riser near the base where it ontacts the regenerated catalyst (see Figure 1-6). The ratio of catalysto-oil is normally in the range of 4:1 to 9:1 by weight. The heat bsorbed by the catalyst in the regenerator provides the energy to heat he feed to its desired reactor temperature. The heat of the reaction ccurring in the riser is endothermic (i.e., it requires energy input). The irculating catalyst provides this energy. The typical regenerated catalyst emperature ranges between 1,250°F to 1,350°F (677°C to 732°C).
Fluid Catalytic Cracking Handbook To Reactor or Cyclone
Catalyst From Regenerator
(Typical for Multiple Nozzles)
Figure 1-6.
Typical riser "Y".
The catalytic reactions occur in the vapor phase. Cracking reactions egin as soon as the feed is vaporized. The expanding volume of the apors that are generated are the main driving force to carry the atalyst up the riser. Catalyst and products are quickly separated in the reactor. However, ome thermal and non-selective catalytic reactions continue. A number
Process Description
9
f refineries are modifying the riser termination devices to minimize hese reactions. The riser is a vertical pipe. It usually has s 4- to 5-inch (10 to 13 m) thick refractory lining for insulation and abrasion resistance. Typical risers are 2 to 6 feet (60 to 180 cm) in diameter and 75 to 120 feet (25 to 30 meters) long. The ideal riser simulates a plug flow eactor, where the catalyst and the vapor travel the length of the riser with minimum back mixing. Efficient contacting of the feed and catalyst is critical for achieving he desired cracking reactions. Steam is commonly used to atomize he feed. Smaller oil droplets increase the availability of feed at the eactive acid sites on the catalyst. With high-activity zeolite catalyst, irtually all of the cracking reactions take place in three seconds or less. Risers are normally designed for an outlet vapor velocity of 50 ft/sec o 75 ft/sec (15.2 to 22.8 m/sec). The average hydrocarbon residence me is about two seconds (based on outlet conditions). As a consequence f the cracking reactions, a hydrogen-deficient material called coke is eposited on the catalyst, reducing catalyst activity.
Catalyst Separation
After exiting the riser, catalyst enters the reactor vessel. In today's CC operations, the reactor serves as a housing for the cyclones. In he early application of FCC, the reactor vessel provided further bed racking, as well as being a device used for additional catalyst separation. Nearly every FCC unit employs some type of inertial separation evice connected on the end of the riser to separate the bulk of the atalyst from the vapors. A number of units use a deflector device to urn the catalyst direction downward. On some units, the riser is irectly attached to a set of cyclones. The term "rough cut" cyclones enerally refers to this type of arrangement. These schemes separate pproximately 75% to 99% of the catalyst from product vapors. Most FCC units employ either single or two-stage cyclones (Figure -7) to separate the remaining catalyst particles from the cracked apors. The cyclones collect and return the catalyst to the stripper hrough the diplegs and flapper/trickle valves (See Figure 1-8). The roduct vapors exit the cyclones and flow to the main fractionator or recovery. The efficiency of a typical two-stage cyclone system s 99.995+%.
0
Fluid Catalytic Cracking Handbook
igure 1-7. A two-stage cyclone system. (Courtesy of Bill Dougherty, BP Oil efinery, Marcus Hook, Pa.)
It is important to separate catalyst and vapors as soon as they enter he reactor. Otherwise, the extended contact time of the vapors with he catalyst in the reactor housing will allow for non-selective catalytic ecracking of some of the desirable products. The extended residence me also promotes thermal cracking of the desirable products.
Process Description
11
Pivot Cyclone Dipleg
Restraint PLAN
Cyclone Dipleg*
Pivot
Restraint ELEVATION Figure 1-8. Typical trickle valve (courtesy of Emtrol Corporation),
tripping Section
As the spent catalyst falls into the stripper, hydrocarbons are adsorbed n the catalyst surface, hydrocarbon vapors fill the catalyst pores, and he vapors entrained with the catalyst also fall into the stripper. tripping steam, at a rate of 2 to 5 Ibs per 1,000 lbs (2 kg to 5 kg er 1,000 kg,) is primarily used to remove the entrained hydrocarbons etween catalyst particles. Stripping steam does not address hydroarbon desorption and hydrocarbons filling the catalyst pores. Howver, reactions continue to occur in the stripper. These reactions are
2
Fluid Catalytic Cracking Handbook
riven by the reactor temperature and the catalyst residence time in he stripper. The higher temperature and longer residence time allow onversion of adsorbed hydrocarbons into "clean lighter" products. oth baffled and unbaffled stripper designs (Figure 1-9) are in commercial use. An efficient stripper design generates intimate contact etween the catalyst and steam. Reactor strippers are commonly esigned for a steam superficial velocity of 0.75 ft/sec (0.23 m/sec) nd a catalyst flux rate of 500 to 700 lbs per minute per square foot .4 kg to 3.4 kg per minute per square meter). At too high a flux,
UPPER STEAM DISTRIBUTOR
LOWER STEAM DISTRIBUTOR
Figure 1-9.
An example of a two-stage stripper.
Process Description
13
he falling catalyst tends to entrain steam, thus reducing the effective– ess of stripping steam. It is important to minimize the amount of hydrocarbon vapors arried through to the regenerator, but not all the hydrocarbon vapors an be displaced from the catalyst pores in the stripper. A fraction of hem are carried with the spent catalyst into the regenerator. These ydrocarbon vapors/liquid have a higher hydrogen-to-carbon ratio than he coke on the catalyst. The drawbacks of allowing these hydrogench hydrocarbons to enter the regenerator are as follows: * Loss of liquid product. Instead of the hydrocarbons burning in the regenerator, they could be recovered as liquid products. « Loss of throughput. The combustion of hydrogen to water produces 3.7 times more heat than the combustion of carbon to carbon dioxide. The increase in the regenerator temperature caused by excess hydrocarbons could exceed the temperature limit of the regenerator internals and force the unit to a reduced feed rate mode of operation. * Loss of catalyst activity. The higher regenerator temperature combined with the formation of steam in the regenerator reduces catalyst activity by destroying the catalyst's crystalline structure.
The flow of spent catalyst to the regenerator is typically controlled y a valve that slides back and forth. This slide valve is controlled y the catalyst level in the stripper. The catalyst height in the stripper rovides the pressure head, which allows the catalyst to flow into the egenerator. The exposed surface of the slide valve is usually lined with refractory to withstand erosion. In a number of earlier FCC esigns, lift air is used to transport the spent catalyst into the regenertor (Figure 1-10).
REGENERATOR–HEAT/CATALYST RECOVERY
The regenerator has two main functions: it restores catalyst activity nd supplies heat to crack the feed. The spent catalyst entering the egenerator contains between 0.4 wt% and 2.5 wt% coke, depending n the quality of the feedstock. Components of coke are carbon, ydrogen, and trace amounts of sulfur and nitrogen. These burn ccording to the following reactions.
4
Fluid Catalytic Cracking Handbook
Products Reactor
Regen Catalyst Standpipe
LiftAir
Air Blower
Figure 1-10. A typical Model II using lift air to transfer spent catalyst.
+ 1/2 CX, O + 1/2 02 + O2 H2,+ 1/2 02 + xO + xO
_> —> —> -> —» ->
CO CO2
CO2 H2O
sox
NO y
K Cal/kg of C, H2, or S
BTU/lb of C, H2, or S
2,200 5,600 7,820 28,900 2,209
3,968 10,100 14,100 52,125 3,983
(1-1) (1-2) (1-3) (1-4) (1-5) (1-6)
Process Description
15
Air provides oxygen for the combustion of coke and is supplied by ne or more air blowers. The air blower provides sufficient air velocity nd pressure to maintain the catalyst bed in a fluid state. The air enters he regenerator through an air distributor (Figure 1-11) located near he bottom of the vessel. The design of an air distributor is important n achieving efficient and reliable catalyst regeneration. Air distributors re typically designed for a 1.0 psi to 2.0 psi (7 to 15 Kpa) pressure rop to ensure positive air flow through all nozzles. There are two regions in the regenerator: the dense phase and the ilute phase. At velocities common in the regenerator, 2 ft/sec to 4 ft/sec 0.6 to 1.2 m/sec), the bulk of catalyst particles are in the dense bed mmediately above the air distributor. The dilute phase is the region bove the dense phase up to the cyclone inlet, and has a substantially ower catalyst concentration.
Standpipe/Slide Valve
During regeneration, the coke level on the catalyst is typically educed to 0.05%. From the regenerator, the catalyst flows down a ransfer line commonly referred to as a standpipe. The standpipe rovides the necessary pressure to circulate the catalyst around the nit. Some standpipes extend into the regenerator, and the top section s often called a catalyst hopper. The hopper, internal to the regenertor, is usually an inverted cone design. In units with "long" catalyst tandpipes, external withdrawal hoppers are often used to feed the tandpipes. The hopper provides sufficient time for the regenerated atalyst to be "de-bubbled" before entering the standpipe. Standpipes are typically sized for a flux rate in the range of 100 to 00 lb/sec/ft2 (500 to 1,500 kg/sec/m2) of circulating catalyst. In most ases, sufficient flue gas is carried down with the regenerated catalyst o keep it fluidized. However, longer standpipes may require external eration to ensure that the catalyst remains fluidized. A gas medium, uch as air, steam, nitrogen, or fuel gas, is injected along the length f the standpipe. The catalyst density in a well-designed standpipe is n the range of 35 to 45 lb/ft3 (560 to 720 kg/m3). The flow rate of the regenerated catalyst to the riser is commonly egulated by either a slide or plug valve. The operation of a slide valve s similar to that of a variable orifice. Slide valve operation is often ontrolled by the reactor temperature. Its main function is to supply
6
Fluid Catalytic Cracking Handbook
igure 1-11. Examples of air distributors. (Top: courtesy of Enpro Systems, nc., Channelview, Texas; bottom: courtesy of VAL-VAMP, Incorporated, ouston, Texas.) Note: These distributors are upside down for fabrication.
Process Description
1?
nough catalyst to heat the feed and achieve the desired reactor emperature. In Exxon Model IV and flexicracker designs (see Figure 1-1), the regenerated catalyst flow is mainly controlled by adjusting he pressure differential between the reactor and regenerator.
Catalyst Separation
As flue gas leaves the dense phase of the regenerator, it entrains atalyst particles. The amount of entrainment largely depends on the lue gas superficial velocity. The larger catalyst particles, 50}i-90p, fall ack into the dense bed. The smaller particles, 0|J,-50ji, are suspended n the dilute phase and carried into the cyclones. Most FCC unit regenerators employ 4 to 16 parallel sets of primary nd secondary cyclones. The cyclones are designed to recover catalyst articles greater than 20 microns diameter. The recovered catalyst articles are returned to the regenerator via the diplegs. The distance above the catalyst bed at which the flue gas velocity as stabilized is referred to as the transport disengaging height (TDH). At this height, the catalyst concentration in the flue gas stays constant; one will fall back into the bed. The centerline of the first-stage cyclone nlets should be at TDH or higher; otherwise, excessive catalyst entrainment will cause extreme catalyst losses.
Flue Gas Heat Recovery Schemes
The flue gas exits the cyclones to a plenum chamber in the top of he regenerator. The hot flue gas holds an appreciable amount of nergy. Various heat recovery schemes are used to recover this energy. n some units, the flue gas is sent to a CO boiler where both the ensible and combustible heat are used to generate high-pressure team. In other units, the flue gas is exchanged with boiler feed water o produce steam via the use of a shell/tube or box heat exchanger. In most units, about two-thirds of the flue gas pressure is let down via n orifice chamber or across an orifice chamber. The orifice chamber is vessel containing a series of perforated plates designed to maintain a iven backpressure upstream of the regenerator pressure control valve. In some larger units, a turbo expander is used to recover this ressure energy. To protect the expander blades from being eroded by atalyst, flue gas is first sent to a third-stage separator to remove the
8
Fluid Catalytic Cracking Handbook
ines. The third-stage separator, which is external to the regenerator, ontains a large number of swirl tubes designed to separate 70% to 5% of the incoming particles from the flue gas. A power recovery train (Figure 1-12) employing a turbo expander sually consists of four parts: the expander, a motor/generator, an air blower, and a steam turbine. The steam turbine is primarily used for tart-up and, often, to supplement the expander to generate electricity. The motor/generator works as a speed controller and flywheel; it can roduce or consume power. In some FCC units, the expander horsepower xceeds the power needed to drive the air blower and the excess power s output to the refinery electrical system. If the expander generates less ower than what is required by the blower, the motor/generator provides he power to hold the power train at the desired speed. From the expander, the flue gas goes through a steam generator to ecover thermal energy. Depending on local environmental regulations, n electrostatic precipitator (ESP) or a wet gas scrubber may be placed ownstream of the waste heat generator prior to release of the flue as to the atmosphere. Some units use an ESP to remove catalyst fines n the range of 5|i-20ji from the flue gas. Some units employ a wet as scrubber to remove both catalyst fines and sulfur compounds from he flue gas stream.
Partial versus Complete Combustion
Catalyst can be regenerated over a range of temperatures and flue as composition with inherent limitations. Two distinctly different modes of regeneration are practiced: partial combustion and complete ombustion. Complete combustion generates more energy when coke ield is increased; partial combustion generates less energy when the oke yield is increased. In complete combustion, the excess reaction omponent is oxygen, so more carbon generates more combustion. In artial combustion, the excess reaction component is carbon, all the xygen is consumed, and an increase in coke yield means a shift from CO2 to CO. FCC regeneration can be further subdivided into low, intermediate, nd high temperature regeneration. In low temperature regeneration about 1,190°F or 640°C), complete combustion is impossible. One f the characteristics of low temperature regeneration is that at 1,190°F, ll three components (O2, CO, and CO2) are present in the flue gas at
UE GAS FROM REGENERATOR
REGENERATOR
Figure 1-12, A typical flue gas power recovery scheme.
ELECTRO PRECIPIT
C F
0
Fluid Catalytic Cracking Handbook
ignificant levels. Low temperature regeneration was the mode of peration that was used in the early implementation of the catalytic racking process. In the early 1970s, high temperature regeneration was developed. High temperature regeneration meant increasing the temperature until ll the oxygen was burned. The main result was low carbon on the egenerated catalyst. This mode of regeneration required maintaining n the flue gas, either a small amount of excess oxygen and no CO, r no excess oxygen and a variable quantity of CO. If there was excess xygen, the operation was in a full burn. If there was excess CO, the peration was in partial burn. With the advent of combustion promoter, the regeneration temerature could be reduced and still maintain full burn. Thus, intermediate emperature regeneration was developed. Intermediate regeneration is ot necessarily stable unless combustion promoter is used to assist in he combustion of CO in the dense phase. Table 1 -2 contains a 2 x 3 matrix summarizing various aspects of regeneration. The following matrix of regeneration temperatures and operating modes shows the inherent limitations of operating regions. Regeneraon is either partial or complete, at low, intermediate, or high ternTable 1-2 A Matrix of Regeneration Characteristics
Operating Region Regenerator Combustion
Partial Combustion Mode
Full Combustion Mode
ow temperature (nominally ,190°F/640°C)
Stable (small afterburning) O2, CO, and CO2 in the flue gas
Not achievable
ntermediate temperature nominally 1,275 °F/690°C)
Stable (with combustion promoter); tends to have high carbon on regenerated catalyst
Stable with combustion promoter
igh temperature (nominally ,350°F/730°C)
Stable operation
Stable operation
Process Description
21
eratures. At low temperatures, regeneration is always partial, carbon n regenerated catalyst is high, and increasing combustion air results n afterburn. At intermediate temperatures, carbon on regenerated atalyst is reduced. The three normal "operating regions" are indicated n Table 1-2. There are some advantages and disadvantages associated with full nd partial combustion:
dvantages of full combustion • Energy efficient • Heat-balances at low coke yield • Minimum hardware (no CO boiler) • Better yields from clean feed
Disadvantages of full combustion • Narrow range of coke yields unless some heat removal system is incorporated • Greater afterburn, particularly with an uneven air or spent catalyst distribution system • Low cat/oil ratio
The choice of partial versus full combustion is dictated by FCC feed uality. With "clean feed," full combustion is the choice. With low uality feed or resid, partial combustion, possibly with heat removal, s the choice.
Catalyst Handling Facilities
Even with proper operation of the reactor and regenerator cyclones, atalyst particles smaller than 20 microns still escape from both of hese vessels. The catalyst fines from the reactor collect in the fraconator bottoms slurry product storage tank. The recoverable catalyst nes exiting the regenerator are removed by the electrostatic preipitator or lost to the environment. Catalyst losses are related to: « « • • •
The design of the cyclones Hydrocarbon vapor and flue gas velocities The catalyst's physical properties High jet velocity Catalyst attrition due to the collision of catalyst particles with the vessel internals and other catalyst particles
2
Fluid Catalytic Cracking Handbook
The activity of catalyst degrades with time. The loss of activity is rimarily due to impurities in the FCC feed, such as nickel, vanadium, nd sodium, and to thermal and hydrothermal deactivation mechanisms. o maintain the desired activity, fresh catalyst is continually added to he unit, Fresh catalyst is stored in a fresh catalyst hopper and, in most nits, is added automatically to the regenerator via a catalyst loader. The circulating catalyst in the FCC unit is called equilibrium atalyst, or simply E-cat. Periodically, quantities of equilibrium catalyst re withdrawn and stored in the E-cat hopper for future disposal. A efinery that processes residue feedstocks can use good-quality E-cat om a refinery that processes light sweet feed. Residue feedstocks ontain large quantities of impurities, such as metals and requires high ates of fresh catalyst. The use of a good-quality E-cat in conjuncon with fresh catalyst can be cost-effective in maintaining low atalyst costs.
MAIN FRACTIONATOR
The purpose of the main fractionator, or main column (Figure 1-13), to desuperheat and recover liquid products from the reactor vapors. he hot product vapors from the reactor flow into the main fractionator ear the base. Fractionation is accomplished by condensing and evaporizing hydrocarbon components as the vapor flows upward hrough trays in the tower. The operation of the main column is similar to a crude tower, but with two differences. First, the reactor effluent vapors must be cooled efore any fractionation begins. Second, large quantities of gases will avel overhead with the unstabilized gasoline for further separation. The bottom section of the main column provides a heat transfer one. Shed decks, disk/doughnut trays, and grid packing are among ome of the contacting devices used to promote vapor/liquid contact. he overhead reactor vapor is desuperheated and cooled by a pumpround stream. The cooled pumparound also serves as a scrubbing medium to wash down catalyst fines entrained in the vapors. Pool uench can be used to maintain the fractionator bottoms temperature elow coking temperature, usually at about 700°F (370°C). The recovered heat from the main column bottoms is commonly sed to preheat the fresh feed, generate steam, serve as a heating medium or the gas plant reboilers, or some combination of these services.
Process Description
Figure 1-13.
23
A typical FCC main fractionator circuit.
The heaviest bottoms product from the main column is commonly alled slurry or decant oil. (In this book, these terms are used interhangeably.) The decant oil is often used as a "cutter stock" with acuum bottoms to make No. 6 fuel oil. High-quality decant oil (low ulfur, low metals, low ash) can be used for carbon black feedstocks. Early FCC units had soft catalyst and inefficient cyclones with ubstantial carryover of catalyst to the main column where it was bsorbed in the bottoms. Those FCC units controlled catalyst losses wo ways. First, they used high recycle rates to return slurry to the eactor. Second, the slurry product was routed through slurry settlers.
4
Fluid Catalytic Cracking Handbook
ither gravity or centrifugal, to remove catalyst fines. A slipstream f FCC feed was used as a carrier to return the collected fines from he separator to the riser. Since then, improvements in the physical roperties of FCC catalyst and in the reactor cyclones have lowered atalyst carry-over. Most units today operate without separators. The ecant oil is sent directly to the storage tank. Catalyst fines accumulate n the tank, which is cleaned periodically. Some units continue to use ome form of slurry settler to minimize the ash content of decanted oil. Above the bottoms product, the main column is often designed for hree possible sidecuts: * Heavy cycle oil (HCO)—used as a pumparound stream, sometimes as recycle to the riser, but rarely as a product * Light cycle oil (LCO)—used as a pumparound stream, sometimes as absorption oil in the gas plant, and stripped as a product for diesel blending; and * Heavy naphtha—used as a pumparound stream, sometimes as absorption oil in the gas plant, and possible blending in the gasoline pool
In many units, the light cycle oil (LCO) is the only sidecut that eaves the unit as a product. LCO is withdrawn from the main column nd routed to a side stripper for flash control. LCO is sometimes eated for sulfur removal prior to being blended into the heating oil ool. In some units, a slipstream of LCO, either stripped or unstripped, sent to the sponge oil absorber in the gas plant. In other units, ponge oil is the cooled, unstripped LCO. Heavy cycle oil, heavy naphtha, and other circulating side pumpround reflux streams are used to remove heat from the fractionator. hey supply reboil heat to the gas plant and generate steam. The mount of heat removed at any pumparound point is set to distribute apor and liquid loads evenly throughout the column and to provide he necessary internal reflux. Unstabilized gasoline and light gases pass up through the main olumn and leave as vapor. The overhead vapor is cooled and partially ondensed in the fractionator overhead condensers. The stream flows o an overhead receiver, typically operating at <15 psig (<1 bar). ydrocarbon vapor, hydrocarbon liquid, and water are separated in he drum.
Process Description
25
The hydrocarbon vapors flow to the wet gas compressor. This gas ream contains not only ethane and lighter gases, but about 95% of he C3 and C4 and about 10% of the naphtha. The phrase "wet gas" efers to condensable components of the gas stream. The hydrocarbon liquid is split. Some is pumped back to the main olumn as reflux and some is pumped forward to the gas plant, ondensed water is also split. Some is pumped back as wash to the verhead condensers and some is pumped away to treating. Some might be used as wash to the wet gas compressor discharge coolers,
GAS PLANT
The FCC gas plant (Figure 1-14) separates the unstabilized gasoline nd light gases into the following: • Fuel gas • C3 and C4 compounds • Gasoline
C3's and C4's include propane, propylene, normal butane, isobutane, nd butylene. Propylene and butylene are used to make ethers and kylate, which are blended to produce high-octane gasoline. Most gas lants also include treating facilities to remove sulfur from these products. The gas plant starts at the wet gas compressor. A two-stage centrifugal ompressor is typically employed. This type of compressor generally ncorporates an electric motor or a multistage turbine that is driven ypically by high-pressure steam. The steam is exhausted to a surface ondenser operating under vacuum. It should be noted that there are CC units in which single-stage wet gas compressors are used. In a two-stage system, the vapors from the compressor's first stage ischarge are partially condensed and flashed in an interstage drum. he liquid hydrocarbon is pumped forward to the gas plant, either to he high pressure separator (HPS) or directly to the stripper. The vapor from the interstage drum flows to the second-stage ompressor. The second-stage compressor discharges through a cooler o the high pressure separator. Gases and light streams from other finery units are often included for recovery of LPG. Recycle streams om the stripper and the primary absorber also go to the high pressure eparator. Wash water is injected to dilute contaminants, such as
Overhead Vapor
Figure 1-14.
Heavy Gasoline
Debut Naphtha
A typical FCC gas plant.
Le
Process Description
27
mmonium salts, that can cause equipment fouling. This mixture Is artially condensed and flashed in the HPS. The vapor from the HPS flows to the primary absorber and the iquid is pumped to the stripper. The HPS is essentially a separation tage with an external cooler located between the primary stripper and bsorber. In some units, they are a single tower,
Primary Absorber
The HPS overhead vapor contains appreciable amounts of C3's and eavier components. The primary absorber recovers these components. The HPS vapor enters below the bottom tray and proceeds up the ower contacting absorption oil. Heavy components are absorbed in he oil. Two sources of absorption oil are normally utilized in this tower. The first is the hydrocarbon liquid from the main fractionator overhead eceiver. This stream, often called "wild," or unstabilized, naphtha, nters the absorber a few trays below the top tray. The second absorbent s cooled debutanized gasoline, which generally enters on the top tray. t has a lower vapor pressure and can be considered a trim absorbent. The expression "lean oil" generally refers to the debutanized gasoline lus the unstabilized naphtha from the overhead receiver. The absorption process is exothermic. To improve C3+ recovery, iquid from one or more of the middle trays is pumped through an ntercooler and returned to the tray below. In some FCC units, the lean il feed is chilled. To enhance C3+ recovery, some units have installed presaturator rums that function as an additional absorption stage. In this operation, he cooled debutanized gasoline is mixed (presaturated) with the bsorber overhead gas. The mixture is cooled and flashed in the resaturator drum. The liquid from this drum is then pumped to the op of the primary absorber.
Sponge OH Absorber
The vapor from the primary absorber or the presaturator contains a mall quantity of gasoline. The sponge oil absorber recovers this asoline. "Sponge oil" is stripped or unstripped light cycle oil. It is sed for final absorption of the dry gas stream. Instead of LCO, a
8
Fluid Catalytic Cracking Handbook
ew FCC units use cooled heavy naphtha from the main column as ponge oil. The lean sponge oil enters the absorber on the top tray. The gas om the presaturator or from the primary absorber enters below the ottom tray. The rich sponge oil from the bottom is then returned to he main fractionator. The lean gas leaves the top of the absorber o an amine unit for H2S removal prior to entering the refinery fuel as system.
tripper or De-ethanizer
The HPS liquid consists mostly of C3's and heavier hydrocarbons; owever, it also contains small fractions of C2's, H2S, and entrained water. The stripper removes these light ends. The liquid enters the tripper on the top tray. The heat for stripping is provided by an xternal reboiler, using steam or debutanizer bottoms as the heat medium. The vapor from the reboiler rises through the tower and strips he lighter fractions from the descending liquid. The rich overhead apor flows to the HPS via the condenser and is fed to the primary bsorber. The stripped naphtha leaves the tower bottoms and goes to he debutanizer. Usually, at least one draw is installed in the tower to emove the entrained water.
Debutanizer
The stripper bottoms contain C3's, C4's, and gasoline; the debutaizer separates the C3's and C4's from the gasoline. In some units, he hot stripper bottoms can be further preheated before entering the ebutanizer. In a number of units, the stripper bottoms is sent directly o the debutanizer. The feed enters about midway in the tower. Debutaizer feed is always partially vaporized because the debutanizer perates at a lower pressure than the stripper. A control valve that egulates stripper bottoms level is the means of this pressure drop. As result of this drop, part of the feed is vaporized across the valve. The debutanizer separates the feed into two products. The overhead roduct contains a mixture of C3's and C4's. The bottoms product is he stabilized gasoline. Heat for separating these products comes from n external reboiler. The heating source is usually the main fractionator eavy cycle oil or slurry. Steam can also be used.
Process Description
29
The overhead product is totally liquefied in the overhead condensers. A portion of the overhead liquid is pumped and returned to the tower s reflux. The remainder is sent to a treating unit to remove H2S and ther sulfur compounds. The mixed C3's and C4's stream can then be ed to an ether or an alkylation unit. It can be fed to a depropanizer ower where the C3's are separated from C4's. The C3's are processed or petrochemical feedstock and the C4's are alkylated. The debutanized gasoline is cooled, first by supplying heat to the tripper reboiler or preheating the debutanizer feed. This is followed y a set of air or water coolers. A portion of the debutanizer bottoms s pumped back to the presaturator or to the primary absorber as lean il. The balance is treated for sulfur and blended into the refinery asoline pool.
Gasoline Splitter
A number of refiners split the debutanized gasoline into "light" and heavy" gasoline. This optimizes the refinery gasoline pool when lending is constrained by sulfur and aromatics. In a few gasoline plitters, a third "heart cut" is withdrawn. This intermediate cut is low n octane and it is processed in another unit for further upgrading.
Water Wash System
The cat cracker feedstock contains low concentrations of organic ulfur and nitrogen compounds. Cracking of organic nitrogen comounds liberates hydrogen cyanide (HCN), ammonia (NH3), and other itrogen compounds. Cracking of organic sulfur compounds produces ydrogen sulfide (H2S) and other sulfur compounds. A wet environment exists in the FCC gas plant. Water comes from he condensation of process steam in the main fractionator overhead ondensers. In the presence of H2S, NH3, and HCN, this environment s conducive to corrosion attacks. The corrosion attack can be any or ll of the following types [2]: * General corrosion from ammonium bisulfide * Hydrogen blistering and/or embrittlement * Pitting corrosion under fouling deposits
Ammonium bisulfide is produced by the reaction of ammonia and ydrogen sulfide [2]:
0
Fluid Catalytic Cracking Handbook
NH3 + H,S -> (NH4) HS MW=17 " MW=34 Weight ratio: NH3/H2S = 0.5
(1-7)
Ammonia bisulfide is extremely corrosive to steel. The corrosion roduct is hydrogen gas and iron sulfide. The reaction is normally selferminating because iron sulfide coats the metal surface with a rotective film that inhibits further corrosion. However, if cyanide is resent, the iron sulfide is removed and bisulfide corrosion is no longer elf-terminating. Hydrogen cyanide is formed in the riser from the eaction of ammonia and CO. Ammonia cyanide is formed from the eaction of hydrogen cyanide and ammonia. The ammonia cyanide will issolve in a wet environment and ionize into cyanide and ammonium ons. The cyanide ion reacts with the insoluble iron sulfide to form a oluble ferrocyonide complex. This destroys the iron sulfide protective lm and exposes fresh metal to further attack. As this corrosion roceeds, it produces hydrogen atoms that penetrate into the metal urfaces causing hydrogen blistering. This leads to stress corrosion racking (SCC). The chemical reactions are: 1. Generation of hydrogen cyanide CO + NH3 -> HCN + H2O
(1-8)
2. Formation of ammonium cyanide HCN + NH3 -»NH 4 CN
(1-9)
3. lonization in water (1-10)
4. Cyanide Corrosion FeS + cyanide —> ferrocyanide + ammonium sulfide
(1-11)
Ammonia can also react with hydrogen sulfide to form ammoium sulfide: 2NH 3 + H2S -> (NH4)2S MW~= 34, MW = 34 " Weight ratio NH3/H2S = 1.0
(1-12)
Process Description
31
Ammonia sulfide is not corrosive, but it can precipitate. Undereposit corrosion and pitting can occur. Typically, sour water from the FCC contains a mixture of ammo– ium sulfide and ammonium bisulfide with an ammonia-to-hydrogen ulfide ratio between 0.5 and 1.0 Most refiners employ continuous water wash as the principal method f controlling corrosion and hydrogen blistering. The best source of water is either steam condensate or well-stripped water from a sour water stripper. A number of refiners use ammonium polysulfate to eutralize hydrogen cyanide and to control hydrogen stress cracking. In the gas plant, corrosive agents (H2S, HCN, and NH3) are most oncentrated at high-pressure points. Water is usually injected into the rst and second-stage compressor discharges. The water contacts the ot gas and scrubs these agents. There are two common injection methods: forward cascading and reverse cascading. In forward cascading (Figure 1-14A), the water is normally injected nto the discharge of the first-stage compressor and condenses in the nterstage cooler. From the interstage drum, the water is pumped to he second-stage discharge, condenses in the cooler, and collects in he EPS, From the high pressure separator, the water is then pressured o the sour water stripper. In reverse cascading (Figure 1-14B), fresh water is injected into the econd-stage discharge. The water containing corrosive agents is ressured to the first-stage discharge and then back to the main actionator overhead. From the overhead receiver, the water is then umped to the sour water stripper. Reverse cascading requires one less ump, but a portion of cyanide captured in the second stage is released n the interstage, forming a cyanide recycle. Consequently, forward ascading is more effective in minimizing cyanide attack.
TREATING FACILITIES
The gas plant products, namely fuel gas, C3's, C4's, and gasoline, ontain sulfur compounds that require treatment. Impurities in the gas lant products are acidic in nature. Examples include hydrogen sulfide H2S), carbon dioxide (CO2), mercaptan (R-SH), phenol (ArOH), and aphthenic acids (R-COOH). Carbonyl and elemental sulfur may also e present in the above streams. These compounds are acidic. (text continued on page 34)
^f
^^ i-2
-^
\
^
)
~~~ ^""> r
•-£3"
L
^
1
:noi aage]
Interstage
1
r«ia
-£T
Main Column Receiver I j ^
_
>
o-o
i
x^-x
^— i
1
1
Figure 1-14A. A typical forward cascading scheme for water wash.
our Water to SWS
f_r
t 1V
Ste ge^
^
5^
1!
oo
N
f F
d Drum i it i j
i
ij -^
Sour W toSW
L J
HPS
1
oo
- -
Main Column Receiver
/*—
( \
Sour Water to SWS
Water
H
OO
±
Figure 1-14B. A typical reverse cascading scheme for water wash.
4
Fluid Catalytic Cracking Handbook
ext continued from page 31)
Amine and caustic solutions are used to remove these impurities, The amine solvents known as alkanolomines remove both H2S and CO2. Hydrogen sulfide is poisonous and toxic. For refinery furnaces nd boilers, the maximum H2S concentration is normally about 160 ppm. Amines remove the bulk of the H2S; primary amines also remove he CO2. Amine treating is not effective for removal of mercaptan. In ddition, it cannot remove enough H2S to meet the copper strip orrosion test. For this reason, caustic treating is the final polishing tep downstream of the amine units. Table 1-3 illustrates the chemistry f some of the important caustic reactions.
Sour Gas Absorber
An amine absorber (Figure 1-15) removes the bulk of H2S from the our gas. The sour gas leaving the sponge oil absorber usually flows nto a separator that removes and liquefies hydrocarbon from vapors, The gas from the separator flows to the bottom of the H2S contactor where it contacts a countercurrent flow of the cooled lean amine from he regenerator. The treated fuel gas leaves the top of the H2S absorber, oes to a settler drum for the removal of entrained solvent, and then ows to the fuel system. Rich amine from the bottom of the H2S contactor goes to a flash eparator to remove dissolved hydrocarbons from the amine solution. he rich amine is pumped from the separator to the amine regenerator. Table 1-3 Acid/Base Reactions Encountered Most Frequently by Oil Industry Caustic Treaters
Carbon Dioxide CO2 + 2 NaOH Hydrogen Sulfide H2S + 2 NaOH Mercaptan Sulfur RSH + NaOH Naphthenic Acid RCOOH + NaOH
—»
Na2CO3 + H2O
—>
Na 2 S + 2 H2O
—»
RSNa + H2O
—>
RCOONa + H2O
(
SWEETGAS
Figure 1-15,
A typical amine treating system.
6
Fluid Catalytic Cracking Handbook
In the amine regenerator, the rich amine solution is heated to reverse he acid-base reaction that takes place in the contactor. The heat is upplied by a steam reboiler. The hot, lean amine is pumped from the ottom of the regenerator and exchanges heat with the rich amine in he lean-rich exchanger and a cooler before returning to the contactor. A portion of the rich amine flows through a particle filter and a arbon bed filter. The particle filters remove dirt, rust, and iron sulfide. he carbon filter, located downstream of the particle filters, removes esidual hydrocarbons from the amine solution. The sour gas, containing small amounts of amine, leaves the top of he regenerator and flows through a condenser to the accumulator. The our gas is sent to the sulfur unit, while the condensed liquid is efluxed to the regenerator. For many years, nearly all the amine units were using monoethanolaiine (MEA) or diethanolamine (DEA). However, in recent years the use f tertiary amines such as methyl diethanolamine (MDEA) has increased. hese solvents are generally less corrosive and require less energy to egenerate. They can be formulated for specific gas recovery requirements.
LPG Treating
The LPG stream containing a mixture of C3's and C4's must be eated to remove hydrogen sulfide and mercaptan. This produces a oncorrosive, less odorous, and less hazardous product. The C3's and 4's from the debutanizer accumulator flow to the bottom of the H2S ontactor. The operation of this contactor is similar to that of the fuel as absorber, except that this is a liquid-liquid contactor. In the LPG contactor the amine is normally the continuous phase with the amine-hydrocarbon interface at the top of the contactor. This nterface level controls the amine flow out of the contactor. (Some quid/liquid contactors are operated with the hydrocarbon as the ontinuous phase. In this case, the interface is controlled at the bottom f the contactor.) The treated C3/C4 stream leaves the top of the contactor. final coalescer is often installed to recover the carry-over amine.
Caustic Treating
Mercaptans are organic sulfur compounds having the general formula f R-S-H. As stated earlier, amine treating is not effective for the
Process Description
37
emoval of mercaptan. There are two options for treating mercaptans. n each option, the mercaptans are first oxidized to disulphides. One option, extraction, dissolves the disulfides in caustic and removes them. The other option, sweetening, leaves the converted disulfides in the product. Extraction removes sulfur, while sweetening just removes the mercaptan odor. Extraction is used for light products (up to light naphtha) and sweetening for heavy products (gasoline through diesel). Sweetening of the FCC gasoline is usually sufficient to meet its ulfur specifications. However, in areas where "reformulated" gasoline s marketed, sulfur specifications in the gasoline may require more reatment. The mercaptans in the LPG need to be extracted to protect he downstream processes, such as alkylation. Sulfur increases acid consumption and produces undesirable by-products. Both sweetening and extraction processes (Figure 1-16) commonly use caustic and catalyst. If the LPG and the gasoline contain high evels of H2S, a caustic prewash is needed to protect the catalyst. The sweetening process utilizes a caustic solution, catalyst, and air, Mercaptans are converted to disulfides in a mixing vessel or fiber film contactor. The reactions take place according to the following equations:
H2O + catalyst -> RSSR + 2NaOH
(1-14)
The mixture of caustic and disulfides is transferred to a settler. From he settler, the treated gasoline flows to a coalescer, sand filter, or wash water tower, before going to storage. The caustic solution is recirculated to the mixing vessel/fiber film contactor. In the extraction process, the LPG from the prewash tower enters he bottom of an extractor column. The extractor is a liquid/liquid contactor in which the LPG is counter-currently contacted by a caustic olution. Another option is the use of a fiber film contacting device. The mercaptans dissolve in the caustic (Equation 1-14). The treated LPG eaves the top of the extractor and goes on to a settler, where entrained austic is separated. From the bottom of the extractor, the caustic solution, containing odium mercaptide, enters the regenerator. Plant air supplies oxygen o react with the sodium mercaptide to form disulfide oil (Equation 1-11), which is insoluble in caustic. The oxidizer overhead stream
CAUSTIC IN (BATCH)
(
•I 1 ] J
CAUSTIC OUT RSNa + NaOH
SECOND STAGE CONTACTOR
CAUSTIC IN
HYDROCARBON STREAM^ w/o H2S, CONTAINS R-SH
f
\
§
1
X^
CATALYST X_ J
4,
AIR
TREATED PRODUCT
1
>
*- AIR
SOLVENT RECYCLE
J ">
O
^
CONTACT
iNER
SOLV
OF
Caustic sweetening and extraction process, (Adapted from Merichem Company, Houston,
EAM
Process Description
39
lows to a disulfide separator. A hydrocarbon solvent, such as naphtha, washes the disulfide oils out of the regenerated caustic. The regenrated caustic is returned to the extractor and the solvent containing isulfide oil is disposed in other units.
SUMMARY
Fluid catalytic cracking is one of the most important conversion rocesses in a petroleum refinery. The process incorporates most hases of chemical engineering fundamentals, such as fluidizalon, heat/mass transfer, and distillation. The heart of the process is he reactor-regenerator, where most of the innovations have occurred ince 1942. The FCC unit converts low-value, high-boiling feedstocks into valuable products such as gasoline and diesel. The FCC is extremely fficient with only about 5% of the feed used as fuel in the process. Coke is deposited on the catalyst during the reaction and burned off n the regenerator, supplying all the heat for the reaction. Products from the reactor are recovered in the main fractionator and he gas plant. The main fractionator recovers the heaviest products, uch as light cycle and decanted oil, from the gasoline and lighter roducts. The gas plant separates the main fractionator overhead vapors nto gasoline, C3's, C4's and fuel gas. The products contain sulfur ompounds and need to be treated prior to being used. A combination f amine and caustic solutions are employed to sweeten these products.
REFERENCES
. Rader, Mari Lyn "Worldwide Refining," Oil & Gas Journal, December 23, 1996, p. 52. , Go, Tony, Baker Petrolite, Houston, TX, personal correspondence 1997.
CHAPTER 2
FCC Feed Characterization
Refiners process many different types of crude oil. As market onditions and crude quality fluctuate, so does cat cracking feedstock. Often the only constant in FCC operations is the continual change in eedstock quality. Feed characterization is the process of determining the physical and hemical properties of the feed. Two feeds with similar boiling point anges may exhibit dramatic differences in cracking performance and roduct yields. FCC feed characterization is one of the most important activities n monitoring cat cracking operation. Understanding feed properties nd knowing their impact on unit performance are essential. Troublehooting, catalyst selection, unit optimization, and subsequent process valuation all depend on the feedstock. Feed characterization relates product yields and qualities to feed uality. Knowing the effects of a feedstock on unit yields, a refiner an purchase the feedstock that maximizes profitability. It is not ncommon for refiners to purchase raw crude oils or FCC feedstocks without knowing their impact on unit operations. This lack of knowldge can be expensive. Sophisticated analytical techniques, such as mass spectrometry, are ot practical for determining complete composition of FCC feedstocks n a routine basis. Simpler empirical correlations are more often used. hey require only routine tests commonly performed by the refinery aboratory. They are excellent alternatives, but they have their limitations: • They are usually intended for an olefin-free feed. • They cannot distinguish among different paraffinic molecules. • They cannot segregate an aromatic compound that may also contain a paraffinic and naphthenic structure group.
40
FCC Feed Characterization
41
Nevertheless, these correlations are very practical tools for tracking nit performance and for troubleshooting. They are also important in process design and catalyst research. The two primary factors that affect feed quality are: * Hydrocarbon Classification * Impurities
HYDROCARBON CLASSIFICATION The hydrocarbon types in the FCC feed are broadly classified as araffins, olefins, naphthenes, and aromatics (PONA),
Paraffins
Paraffins are straight or branched chain hydrocarbons having the hemical formula CnH2n+2. The name of each member ends with -ane; xamples are propane, isopentane, and normal heptane (Figure 2-1). In general, FCC feeds are predominately paraffinic. The paraffinic arbon content is typically between 50 wt% and 65 wt% of the total eed. Paraffinic stocks are easy to crack and normally yield the greatest mount of total liquid products. They make the most gasoline and the east fuel gas, but also the lowest octane number.
H
7
?
H
H
?
H—
H—C — H
PROPANE (C3H8)
7
f
f
C — C — C—- H H
I H H—T— H j H
ISOPENTANE (CSH12)
HHHT H
H
H
H
H
H
H
NORMAL HEPTANE (C7H16) Figure 2-1.
Examples of paraffins.
2
Fluid Catalytic Cracking Handbook
Olefins
Olefins are unsaturated compounds with a formula of CnH2n. The ame of these compounds ends with ~ene, such as ethene (ethylene) nd propene (propylene). Figure 2-2 shows typical examples of olefins. Compared to paraffins, olefins are unstable and can react with themelves or with other compounds such as oxygen and bromine solution. Olefins do not occur naturally; they show up in the FCC feed as a esult of preprocessing the feeds elsewhere. These processes include hermal cracking and other catalytic cracking operations. Olefins are not the preferred feedstocks to an FCC unit. This is not ecause olefins are inherently bad, but because olefins in the FCC feed ndicate thermally produced oil. They often polymerize to form ndesirable products, such as slurry and coke. The typical olefin ontent of FCC feed is less than 5 wt%, unless unhydrotreated coker as oils are being charged.
Napfathenes
Naptithenes (CirH2n) have the same formula as olefins, but their haracteristics are significantly different. Unlike olefins that are traight-chain compounds, naphthenes are paraffins that have been bent" into a ring or a cyclic shape. Naphthenes, like paraffins, are aturated compounds. Examples of naphthenes are cyclopentane, yclohexane, and methylcyclohexane (Figure 2-3). H
H
H H
H—C = C—H
H—C—C =C—H H
ETHYLENE
H
PROPYLENE
m?
c = c—c — H
H
H
BUTENE-2
Figure 2-2,
Examples of olefins.
FCC Feed Characterization
43
CH3
H2C H2C /
\
H2C
H2C
\
/
CH
/ H2C
\ H2C
/ H2C
1 H2C
1 H2C
1i C* H 2<-
H2C— H2C
\
Cyclopentane (C5H!0)
/
H2C
\ H2C 1i
1 "Ut Jtlif" 2C
\
Cyclohexane (C6H12)
H2C
/
Methy-Cyclohexane (C7H14)
Figure 2-3. Examples of naphthenes.
Naphthenes are desirable FCC feedstocks because they produce high-octane gasoline. The gasoline derived from the cracking of naphthenes has more aromatics and is heavier than the gasoline produced from the cracking of paraffins.
Aromatics
Aromatics (CnH2n_6) are similar to naphthenes, but they contain a esonance stabilized unsaturated ring core. Aromatics (Figure 2-4) are compounds that contain at least one benzene ring. The benzene ring s very stable and does not crack to smaller components. Aromatics
(:— H
H— C)""
H
CH,
H
e v»^
CD— H
V y H BENZENE (C6H6)
H
r \_
C-H
H-C
i_i P-i
H-C
C
H-^<
^c^ o H
TOLUENE (C7H8)
Figure 2-4. Examples of aromatics.
H "C I H
ANILINE (C6H5NH2}
4
Fluid Catalytic Cracking Handbook
re not a preferred feedstock because few of the molecules will crack, he cracking of aromatics mainly involves breaking off the side chains esulting in excess fuel gas yield. In addition, some of the aromatic ompounds contain several rings (polynuclear aromatics) than can compact" to form what is commonly called "chicken wire." Figure -5 illustrates three examples of a polynuclear aromatic compound. ome of these compacted aromatics will end up on the catalyst as arbon residue (coke), and some will become slurry product. In
Anthracene
Naphthalene
Fluorene Figure 2-5. Examples of a polynuclear aromatic molecules.
FCC Feed Characterization
45
omparison with cracking paraffins, cracking aromatic stocks results n lower conversion, lower gasoline yield, and less liquid volume gain, but with higher gasoline octane.
Characterizing an FCC feedstock involves determining both its hemical and physical properties. Because sophisticated analytical echniques, such as mass spectrometry, are not practical on a daily basis, physical properties are used. They provide qualitative measurement of the feed's composition. The refinery laboratory is usually quipped to carry out these physical property tests on a routine basis. The most widely used properties are: * » * * * * *
°API Gravity Distillation Aniline Point Refractive Index (RI) Bromine Number (BN) and Bromine Index (BI) Viscosity Conradson, Ramsbottom, Microcarbon, and Heptane Insoluble
°API Gravity
The °API gravity measures the density of a hydrocarbon liquid. Specific gravity (SG) is another common measurement of density. The iquid SG is the relative weight of a volume of sample to the weight f the same volume of water at 60°F (15.5°C). Compared with specific gravity, °API gravity magnifies small hanges in the feed density. For example, going from 24°API to 26°API changes the specific gravity by 0.011 and the density by 0.72 b/ft3 (0.0115 gram/cm3). Neither is very significant, but a two-number hift in °API gravity can have significant effects on yields. The SG relates to °API gravity by the following equations: ^J(@60°F)
3
141.5 131.5+ °
API Gravitv =
SG(at 60°F)
1 3 1.5
§
Fluid Catalytic Cracking Handbook
Since °API gravity is inversely proportional to specific gravity, the igher the °API gravity, the lighter the liquid sample. In petroleum efining, °API gravity is routinely measured for every feed and product tream. The ASTM D-287 is a hydrometer test typically performed by lab technician or unit operator. The method involves inserting a glass ydrometer into a cylinder containing the sample and reading the °AP1 ravity and the fluid temperature on the hydrometer scale. Standard ables similar to Table 2-1 convert the °API at any temperature back o 60°R The °API gravity is always reported at 60°F (15.5°C). For a highly paraffinic (waxy) feed, the sample should be heated o about 120°F (49°C) before immersing the hydrometer for testing. Heating ensures that the wax is melted, eliminating erroneous readings. Daily monitoring of °API gravity provides the operator with a tool o predict changes in unit operation. For the same distillation range, he 26°API feed cracks more easily than the 24°API feed because the 6°API feed has more long-chain paraffinic molecules. In contact with he 1,300°F (704°C) catalyst, these molecules are easier to rupture into aluable products. Long straight-chain paraffins are important to the economics of an CC unit. They crack easily to gasoline and LPG, with minimal roduction of slurry and fuel gas.
Table 2-1 API Gravity at Observed Temperature
Observed emperature °F 70 75 80 85 90 95 100 105 110 115
18
19
20
21
22
23
24
Corresponding API Gravity at 60°F
17.4 17.2 16.9 16.6 16.4 16.1 15.8 15.6 15.3 15.0
18.4 18.2 17.9 17.6 17.3 17.1 16.8 16.5 16.3 16.0
19.4 19.1 18.9 18.6 18.3 18.0 17.8 17.5 17.2 16.9
20.4 20.1 19.8 19.6 19.3 19.0 18.7 18.4 18.2 17.9
21.4 21.1 20.8 20.5 20.3 20.0 19.7 19.4 19.1 18.9
22.4 22.1 21.8 21.5 21.2 20.9 20.7 20.4 20.1 19.8
23.4 23.1 22.8 22.5 22.2 21.9 21.6 21.3 21.1 20.8
FCC Feed Characterization
47
The simple °API gravity test provides valuable information about he quality of a feed. But the shift in °API usually signals changes in other feed properties, such as carbon residue and aniline point. Addiional tests are needed to fully characterize the feed.
Distillation
Boiling point distillation data also provides information about the quality and composition of a feed. The significance is discussed later n this chapter. Distillation indicates molecular weight and carbon number. It indicates whether the feed contains any "clean" products hat could be sold "as is." Before discussing the data, the different esting methods and their limitations should be reviewed. The feed to the cat cracker in a typical refinery is a blend of gas oils rom such operating units as the crude, vacuum, solvent deasphalting, and coker. Some refiners purchase outside FCC feedstocks to keep the FCC feed rate maximized. Other refiners process atmospheric or vacuum residue in their cat crackers. In recent years, the trend has been toward heavier gas oils and residue. Residue is most commonly defined as the fraction of feed that boils above 1,050°F (565°C). Each FCC feed stream has different distillation characteristics. The frequency and method of testing feed streams varies from one efiner to another. Some refiners analyze daily, others two or three imes a week, and some once a week. The frequency depends on how he distillation results are applied, the variation in crude slates, and he availability of lab personnel. The fractional distillation test conducted in the laboratory involves measuring the temperature of the distilled vapor at the initial boiling point (IBP), as volume percent fractions 5, 10, 20, 30, 40, 50, 60, 70, 80, 90, and 95 are collected, and at the end point (EP). Three ASTM methods are currently used to measure boiling points: D-86, D-1160, and D-2887. D-86 is the most common method used in refineries. The distillation s done at atmospheric pressure. It is used for samples with an EP less han 750°F (400°C). Above this temperature, the sample begins to rack. Thermal cracking is identified by a drop in the temperature of distilled vapor, the presence of brown smoke, and a rise in the system pressure. Above 750°F liquid temperature, the distilling flask begins o deform. All of today's FCC feeds are too heavy to use the D-86
8
Fluid Catalytic Cracking Handbook
method, but it is used for light products. Therefore, the boiling points re obtained using either ASTM D-1160 or ASTM D-2887. D-1160 is run under vacuum (one millimeter of mercury). The esults are converted to atmospheric pressure, using standard corelations. Some newer apparatuses have built-in software that performs he conversion automatically. D-1160 is limited to a maximum EP emperature of about 1,000°F (538°C) at atmospheric pressure. Above his temperature, the sample begins to crack thermally. D-2887 is a low-temperature simulated distillation (SIMDIS) method hat measures the vol% of true boiling point (TBP) cuts using gas hromatography (GC). Like D-1160, its use is limited to a maximum nd point of about 1,000°F (538°C). However, new GC~based systems n the market can measure boiling point temperatures as high as ,35Q°F (750°C). This development will be useful in determining oiling points of feedstocks that contain residue and in the charcterizing of raw crades. Compared with D-1160, SIMDIS is less labor ntensive, more reproducible, and generally more accurate at the IBP nd at the 5% and 10% points. Distillation data provides information about the light fraction of feed oiling at less than 650°F (343°C). Light virgin feed, the fraction that oils under 650°F (343°C), often results in a greater LCO yield and ower unit conversion. Sources of these fractions come from atmosheric gas oil, light vacuum gas oil, and light coker gas oil. Lower onversion of light virgin feed is caused by: 1. Lower molecular weight means the oil is more difficult to crack, 2. Light coker stocks are very aromatic. 3. Light aromatics have fewer crackable side chains.
Economics and unit configuration dictate whether to include 650°F material in the FCC feed. As a general rule, this fraction should be minimized. Minor improvements in the operation of the upstream istillation columns can substantially reduce the amount of light gas il in the FCC feed. However, including light gas oil in FCC feed educes the amount of coke laid on the catalyst. Less coke means a ower regenerator temperature. Light gas oil can be used as a "quench" o decrease the regenerator temperature and to increase the catalysto-oil ratio. The distillation also provides information about the fractions that oil over 900°F (482°C). These fractions provide an indication of the
FCC Feed Characterization
41
coke-making tendency of a given feed. Associated with this 900°F+ raction is a higher level of contaminants such as metals and nitrogen. As discussed later in this chapter, these contaminants deactivate the catalyst and produce less liquid product and more coke and gas. Distillation data is the backbone of FCC feed analyses. As will be shown, published correlations use distillation data to determine the chemical composition of FCC feed.
Aniline Point
Aniline is an aromatic amine (C6H5NH2). When used as a solvent, t is selective to aromatic molecules at low temperatures, and paraffins and naphthenes at higher temperature. Aniline is used to determine aromaticity of oil products, including FCC feedstocks. Aniline point AP) is the minimum temperature for complete solubility of an oil sample in aniline. ASTM D-611 involves heating a 50/50 mixture of the feed sample and aniline until there is only one phase. The mixture is then cooled and the temperature at which the mixture becomes suddenly cloudy s the aniline point. The test senses solubility via a light source that penetrates through the sample. The aniline point increases with paraffinicity and decreases with aromaticity. It also increases with molecular weight. Naphthenes and olefins show values that lie between those for paraffins and aromatics. Typically, an aniline point higher than 200°F (93°C) indicates paraffinicity and an aniline point lower than 150°F (65°C) indicates aromaticity. Aniline point is used in some correlations to estimate the aromaticity of gas oil and light stocks. TOTAL'S [1] correlation uses aniline point and refractive index. Others, such as n-d-M [2], employ refractive ndex to characterize FCC feed.
Refractive Index
Similar to aniline point, refractive index (RI) shows how refractive or aromatic a sample is. The higher the RI, the more the aromatics and the less crackable the sample. A feed having an RI of 1.5105 is more difficult to crack than a feed with an RI of 1.4990. The RI can be measured in a lab (ASTM D-1218) or predicted using correlations uch as the one published by TOTAL.
0
Fluid Catalytic Cracking Handbook
In the laboratory, RI is measured using a refractometer. The instrument has two prisms and a light source. The technician spreads a small mount of sample on the faces of both prisms in the refractometer. he light is then directed at the sample and the scale is read. The bserved scale is then converted to a refractive index with tables upplied with the instrument and corrected for the sample temperature. Both refractive index and aniline point tests qualitatively measure he aromaticity of a stock. With dark and viscous samples, both methods have their limitations. For darker samples, the aniline point est is slightly more accurate because of its larger scale over the ame range of aromatics. The industry does not agree as to which method is more accurate. The three published correlations that will e discussed later, use the refractive index at 68°F (20°C) for alculating feed composition. However, at 68°F most FCC feeds are olid and their refractive indexes cannot be determined accurately. oth the TOTAL and API [31 correlations predict Rl values using eed properties such as specific gravity, molecular weight, and verage boiling point.
Bromine Number and Bromine Index
Bromine number (ASTM D-1159) and bromine index (ASTM D710) are qualitative methods to measure the reactive sites of a sample. romine reacts not only with olefin bonds, but also with basic nitrogen molecules and with some aromatic sulfur derivatives. Nevertheless, lefins are the most common reactive sites and the bromine number used to indicate olefinicity of the feed. The bromine number is the number of grams of bromine that will eact with 100 grams of the sample. Typical bromine numbers are: » Less than 5 for hydrotreated feeds • 10 for heavy vacuum gas oil * 50 for coker gas oil
A general rule-of-thumb is that the olefin fraction of the sample is ne-half of its bromine number. Alternatively, the bromine index is the number of milligrams of romine that will react with 100 grams of the sample, and is most ften used by the chemical industry for stocks that have very low lefin contents.
FCC Feed Characterization
51
Viscosity
Viscosity indicates the chemical composition of an oil sample. As he viscosity of a sample increases, paraffins increase, hydrogen ontent increases, and the aromatic fraction decreases. Viscosity is normally measured at two different temperatures: typically OO°F (38°C) and 210°F (99°C). For many FCC feeds, the sample is oo thick to flow at 100°F and the sample is heated to about 130°F, The viscosity data at two temperatures are plotted on a viscosityemperature chart (see Appendix 1), which shows viscosity over a wide temperature range [4]. Viscosity is not a linear function of temerature and the scales on these charts are adjusted to make the elationship linear. Viscosity is a measurement of resistance to flow. Although the unit f absolute viscosity is poise, its measurement is difficult. Instead, kinematic (flowing) viscosity is determined by measuring the time for given flow through a capillary tube of specific diameter and length. The unit of kinematic viscosity is the stoke. However, in general ractice, centistoke is used. Poise is related to stoke by the equation: „ . . Centipoise Centistokes = Density
Saybolt viscosity (ASTM D-88) is the most popular method of measuring the viscosity of oils such as FCC feed. This method covers wo procedures: * Saybolt Universal Viscometer (SUV) for light oils • Saybolt Furol Viscometer (SFV) for heavy oils
Both procedures measure the time for a fixed volume of the sample o flow though a calibrated tube at a controlled temperature. The difference between the two instruments is the inside diameter ID) of the outlet flow tube. The SUV uses a 0.176 centimeter ID and he SFV uses a 0.315 centimeter ID. The SFV is used for samples that ave a flowing time greater than 600 seconds. For most conventional as oils, the flowing time is short enough that the Universal Viscometer is frequently used. The tube dimensions in the two procedures
2
Fluid Catalytic Cracking Handbook
re such that the Furol viscosity of oil is numerically one-tenth (-^-) f the Universal viscosity at the same temperature.
Conradson, Rams bottom, Microcarbon, and Heptane Insolubles
One area of cat cracking not fully understood is the proper determiation of carbon residue of the feed and how it affects the unit's coke make. Carbon residue is defined as the carbonaceous residue formed fter thermal destruction of a sample. Cat crackers are generally imited in coke burn capacity, therefore, the inclusion of residue in he feed produces more coke and forces a reduction in FCC throughput. Conventional gas oil feeds generally have a carbon residue less than .5 wt%; for feeds containing resid, the number can be as high as 15 wt%. Four popular tests are presently used to measure carbon residue or oncarbon of FCC feedstocks: » « • *
Conradson Ramsbottom Micro-method Heptane insolubles
The object is to indicate the relative coke forming tendency of eedstocks. Each test has advantages and disadvantages, but none of hem provide a rigorous definition of carbon residue or asphaltenes. The Conradson test (ASTM D-189) measures carbon residue by vaporative and destructive distillation. The sample is placed in a reweighed sample dish. The sample is heated, using a gas burner, ntil vapor ceases to burn and no blue smoke is observed. After ooling, the sample dish is reweighed to calculate the percent carbon esidue. The test, though popular, is not a good measure of the cokeorming tendency of FCC feed because it indicates thermal, rather than atalytic, coke. In addition, the test is labor intensive and is usually ot reproducible, and the procedure tends to be subjective. The Ramsbottom test (ASTM D-524) is also used to measure carbon esidue. The test calls for introducing 4 grams of sample into a preweighed glass bulb, then inserting the bulb in a heated bath for 20 minutes. The bath temperature is maintained at 1,027°F (553°C). After 0 minutes, the sample bulb is cooled and reweighed. Compared with he Conradson test, Ramsbottom is more precise and reproducible,
FCC Feed Characterization
53
Both tests produce similar results and often are interchangeable (see Figure 2-6). The Micro-method uses an analytical instrument to measure Conradson arbon in a small automated set. The Micro-method (ASTM D4530) ives test results that are equivalent to the Conradson carbon residue est (D189). The purpose of this test is to provide some indication of he relative coke forming tendency of such material. The heptane insoluble (ASTM D-3279) method is commonly used o measure the asphaltene content of the feed. Asphaltenes are clusters f poly nuclear aromatic sheets, but no one has a clear understanding f their molecular structure. They are insoluble in C3 to C7 paraffins. The amount of asphaltenes that precipitate varies from one solvent to nother, so it is important that the reported asphaltene values be dentified with the appropriate solvent. Both normal heptane and i \J\J
q _j
1 10 5
, f: - ' Jj
c**
o
8 $
/'
/
~ \
/^
1
E
_ . -I
- :\
o
?
' "1
Ht! '
Q
o | 1
_-!
0.1
X^
••'^
T- =
U _ \
0 01
OJ31
0.1
1
10
100
Conradson Carbon, wt%
igure 2-6. Ramsbottom Carbon Residue versus Conradson Carbon Residue, Copyright ASTM D-524. Reprinted with permission.)
4
Fluid Catalytic Cracking Handbook
entane insolubles are widely used for measuring asphaltenes. Although hey do not provide rigorous definitions of asphaltenes, they provide ractical ways of assessing coke precursors in FCC feedstocks. It hould be noted that the traditional definition of asphaltenes is that hey are heptane insoluble. Pentane insoluble minus heptane insoluble s the definition of resins. Resins are molecules larger than aromatics nd smaller than asphaltenes,
MPURITIES
In recent years, refiners have been processing heavier crudes. Heavier crudes provide a slim economic margin to the refiner, but slim s better than none. The cat cracker, as the main conversion unit, is esigned to handle a variety of feedstocks. Today's FCC feedstocks re generally heavier and contain higher levels of impurities like itrogen, sulfur, and metals. These impurities have negative effects on nit performance. Understanding the nature and effects of these ontaminants is essential in feed and catalyst selection as well as roubleshooting the unit. Most of the impurities in the FCC feed exist as components of large rganic molecules. The most common contaminants are: * Nitrogen * Sulfur » Nickel « Vanadium * Sodium
Except for sulfur, all these contaminants poison the FCC catalyst, ausing it to lose its ability to produce valuable products. Sulfur in he feed increases operating costs because additional feed and product reatment facilities are required to meet product specifications and omply with environmental regulations.
Nitrogen
Nitrogen in the FCC feed refers to organic nitrogen compounds. The itrogen content of FCC feed is often reported as basic and total itrogen. Total nitrogen is the sum of basic and nonbasic nitrogen. Basic nitrogen is about one fourth to one half of total nitrogen.
FCC Feed Characterization
55
The word "basic" denotes molecules that react with acids. Basic itrogen compounds will neutralize acid sites on the catalyst. This auses a temporary loss of catalyst activity and a drop in unit conersion (Figure 2-7). However, nitrogen is a temporary poison. The urning of nitrogen in the regenerator restores the activity of he catalyst. In the regenerator, about 70% to 90% of the nitrogen in he coke is converted to elemental nitrogen. The remaining nitrogen is onverted to nitrogen oxides (NOX). The NOx leaves the unit with the lue gas. Catalyst poisoning from the presence of basic nitrogen in the FCC eedstock is significant and, unfortunately, very little attention is often iven to the deleterious effects of basic nitrogen. Virtually all the basic nitrogen ends up in coke. As shown in Figure 2-7, about 1 vol% of he FCC gasoline is lost for each 100 ppm of basic nitrogen in the eedstock. To compensate for nitrogen poisoning, the reactor temperature s raised. A catalyst with high zeolite content and an active matrix Is lso recommended. For some refiners, hydrotreating the feed may be an appropriate conomical approach. Except for most of the California crudes and a
r
82 —
«
80 78
C
o
1 C
o o
76 74 72 70 5C)0
1000
1500
2000
Total Nitrogen, ppm Figure 2-7.
Effect of FCC feed nitrogen on unit conversion.
6
Fluid Catalytic Cracking Handbook
ew others, feeds with high nitrogen also have other impurities. Therefore, it is difficult to evaluate deleterious effects of nitrogen lone. Hydrotreating the feed reduces not only the nitrogen content ut also most other contaminants. Aside from catalyst poisoning, nitrogen is detrimental to the unit peration in several other areas. In the riser, some of the nitrogen is onverted to ammonia and cyanide (H-CN). Cyanide accelerates the orrosion rate of the FCC gas plant equipment; it removes the protecive sulfide scale and exposes bare metal to further corrosion. This orrosion generates atomic hydrogen that ultimately results in hydrogen listering. Cyanide formation tends to increase with cracking severity. In addition, some of the nitrogen compounds end up in light cycle il (LCO) as pyrolles and pyridines [5]. These compounds are easily xidized and will affect color stability. The amount of nitrogen in the LCO depends on the conversion. An increase in conversion decreases he percentage of nitrogen in the LCO and increases the percentage n the catalyst. The source and gravity range of raw crude greatly influence the mount of nitrogen in the FCC feed (Table 2-2). Generally speaking, Table 2-2 API Gravity, Residue, and Nitrogen Content of Typical Crudes
Crude Source
Maya Alaska North Slope (ANS) Arabian Medium orcados Cabinda Arabian Light onny Light rent West Texas Intermediate Gushing (WTIC) orties
°API Gravity
Vacuum Bottoms, vol%
Total Nitrogen* of Heavy Vacuum Gas Oil, PPM
21.6 28.4
33.5 20.4
2498 1845
28.7 29.5 32.5 32.7 35.1 38.4 38.7
23.4 7.6 23.1 17.2 5.3 11.4 10.6
829 1746 1504 1047 1964 1450 951
39.0
10.1
1407
Nitrogen level varies with crude source and residue content.
FCC Feed Characterization
57
eavier crudes contain more nitrogen than the lighter crudes. In ddition, nitrogen tends to concentrate in the residue portion of the rude. Figure 2-8 shows examples of nitrogen compounds found in rude oil.
A.
B.
C.
Neutral N - Compounds
N-H Carbazole
Basic N-Compounds
"N Acridine
'N' Quinoline
o
COO
Phenanthridine
Weakly Basic N-Compounds
IN
N--OH Hydroxiquinoline
Deriviates with R = H, alkyl, phenyl-, naphthyl-
Nitrogen Distribution in Several Middle Eastern Oils (6)
Content:
ype:
*OH
Hydroxipyridine
20-25% of nitrogen in 225-540°C gas oil fraction. 75-80% of nitrogen in 540°C plus vacuum resid fraction. 225-540°C gas oil fraction: 50% of nitrogen as neutral nitrogen compounds; 33% as basic, 17% as weakly basic. 540°C plus vacuum resid fraction: 20% of nitrogen in asphaltenes, 33% as neutral, 20% as basic, 27% as weakly baste.
Figure 2-8. Types of nitrogen compounds in crude oil [12].
8
Fluid Catalytic Cracking Handbook
UOP Test Method 313 is commonly employed to determine the asic nitrogen content of FCC feed. The feed sample is first mixed 0/50 with acetic acid. The mixture is then titrated with perchloric acid. ASTM Method D-3228 (or chemiluminescent nitrogen detector) is ften employed to measure the total nitrogen and involves converting ll the nitrogen in the feed to ammonia and then titrating it with tandard sulfuric acid. Another method (ASTM D-4629) is used to measure the total nitrogen through oxidative combustion and chemiluminescent detection.
Sulfur
FCC feedstocks contain sulfur in the form of organic-sulfur comounds such as mercaptan, sulfide, and thiophenes. Frequently, as the esidue content of crade oil increases, so does the sulfur content (Table -5). Total sulfur in FCC feed is determined by the wavelength ispersive x-ray fluorescence spectrometry method (ASTM D-2622). The results are expressed as elemental sulfur. Although desulfurization is not the goal of cat cracking operations, pproximately 50% of sulfur in the feed is converted to H2S. In ddition, the remaining sulfur compounds in the FCC products are ighter and can be desulfurized by low-pressure hydrodesulfurizaon processing. In the FCC, H2S is formed principally by the catalytic decomposion of non-thiophenic (non-ring) sulfur compounds. Table 2-3 shows he effects of feedstock sulfur compounds on H2S production. As with H2S, the distribution of sulfur among the other FCC proucts depends on several factors, which include feed, catalyst type, onversion, and operating conditions. Feed type and residence time are he most significant variables. Sulfur distribution in FCC products of everal feedstocks is shown in Table 2-4. Figure 2-9 illustrates the ulfur distribution as a function of the unit conversion. For nonhydrotreated feeds at 78 vol% conversion, about 50 wt% of he sulfur in the feed is converted to hydrogen sulfide (H2S). The emaining 50% of the sulfur is distributed approximately as follows: » « * •
6 wt% in gasoline 23 wt% in light cycle oil 15 wt% in decanted oil 6 wt% in coke
FCC Feed Characterization,
59
Table 2-3 Effects of Feedstock Sulfur Compounds on H2S Production Cracking Conditions: 7 Cat/Oil Ratio, 950°F, Zeolite Catalyst
Feed Source
Mid Continent West Texas Coker Gas Oil Hydrotreated West Texas Heavy Cycle Oil
Conversion Vol%
% of Feed Sulfur Which Is Mercapcan or Sulfide and Not Aromatic in Nature
Vol% of Sulfur Converted* to H2S
72 69 56 77
38 33 30 12
47 4! 35 26
50
6
16
The % sulfur converted to H-,S depends largely on the type of sulfur in the feed and the residence time of the hydrocarbons in the riser. ource: Wollaston [6]
Adding residue to the feed increases the sulfur content of coke roportional to the incremental sulfur in the feed (Table 2-6). Thiophenic ring-type) sulfur compounds crack more slowly, and the uncracked hiophenes end up in gasoline, light cycle oil, and decanted oil. Hydrotreating reduces the sulfur content of all the products. With ydrotreated feeds, more of the feed sulfur goes to coke and heavy iquid products. The same sulfur atoms that were converted to H2S in he FCC process are also being removed first in the hydrotreating rocess. The remaining sulfur compounds are harder to remove. The eavier and more aromatic the feedstock, the greater the level of sulfur n the coke (Table 2-7). Although hydrotreating increases the percentage of sulfur in coke nd slurry, the actual amount of sulfur is substantially less than in the ontreated feeds. Sulfur still plays a minor role in unit conversion and ields. Its affect on processing is minimal. Some aromatic sulfur ompounds do not convert, but this is no different from other aromatic ompounds. They become predominately cycle oil and slurry. This ends to lower conversion and reduce maximum yields. (text continued on page 62)
0
Fluid Catalytic Cracking Handbook Table 2-4 Sulfur Distribution in FCC Products
Feedstock Sources Kuwait
Feedstock
DAO & Gas
W. Texas Virgin Gas Oil
W. Texas Virgin Gas Oil (HOT)
California
Oil Blend
Gas Oil
(HOT)
1.75
0.21
1.15
3.14
ulfur Content, wt% Conversion, vol%
78.7
77.8
77.8
80.1
Sulfur Distribution, Wt% of Feed Sulfur
H2S Light Gasoline Heavy Gasoline LCO Decanted Oil Coke
60.2 1.6 7.9 20.7 6.8 2.8
19.2 0.9 1.9 34.6 34.7 8.7
42.9 0.2 3.3 28.0 20.5 5.1
50.0 1.9 5.0 17.3 15.3 10.3
ource: Huling [7]
D
Table 2-5 API Gravity, Residue, and Sulfur Content of Some Typical Crudes
Crude Source
Maya Alaska North Slope (ANS) Arabian Medium orcados Cabinda Arabian Light Bonny Light Brent West Texas Intermediate Gushing (WTIC) orties
Sulfur Content of Vacuum Gas Oil,
°API Gravity
Vacuum Bottoms, vol%
21.6 28.4 28.7 29.5 32.5 32.7 35.1 38.4 38.7
33.5 20.4 23.4 7,6 23.1 17.2 5.3 11.4 10.6
3.35 1.45 3.19 0.30 0.16 2.75 0.25 0.63 0.63
39.0
10.1
0.61
Sulfur level varies with crude source and residue content.
Wt%*
Figure 2-9,
Sulfur distribution of the FCC products as a function of unit conversio
2
Fluid Catalytic Cracking Handbook Table 2-6 Sulfur Content of Coke vs. Quantity of Residue in FCC Feed* Pilot Plant Data, Riser Cracking for Maximum Liquid Recovery Feedstock Type
Gas Oil Gas Oil + 10% of West Texas Sour VTB Gas Oil + 10% of West Texas Sour VTB
Feed Sulfur, Wt%
Sulfur in Coke Wt% of Feed
0.7 1.0
3,5 13,8
1.32
18.6
As the residue content of the feed is increased, there is a marked increase in the coke's sulfur due to higher coke yield and a higher sulfur content of the coke precursors, ource: Campagna [8]
Table 2-7 Sulfur Content of Coke vs. Hydrotreated* FCC Feed Quality Pilot Plant Data, Riser Cracking for Maximum Liquid Recovery Feedstock Source
Feedstock Sulfur, Wt%
Hydrocarbon Type % Tri-aromatics*
Sulfur in Coke, Wt% of Feed
Light Arabian HDS Heavy Arabian HDS Maya HDS
0.21 0.37 0.70
7.3 17.6 5.0
28. 1 48. 2 43.7
In a hydrotreated feed, the more polyaromatic type sulfur compounds, the more sulfur ends up in coke. ource: Campagna [8]
text continued from page 59)
Metals
Metals, such as nickel, vanadium, and sodium, are present in crude oil. These metals are concentrated in the heavy boiling range of tmospheric bottoms or vacuum residue, unless they are carried over with the gas oil by entrainment. These metals are catalysts themselves and promote undesirable eactions, such as dehydrogenation and condensation. Dehydrogenation means the removal of hydrogen; and condensation means polymerization,
FCC Feed Characterization
S3
which is the formation of "chicken wire" aromatic molecules. Hydrogen nd coke yields are increased and gasoline yields are reduced. Metals educe the catalyst's ability to produce the desired products. These metals permanently poison the FCC catalyst by lowering the atalyst activity, thereby reducing its ability to produce the desired roducts. Virtually all the rnetals in the FCC feed are deposited on he cracking catalyst. Paraffinic feeds tend to contain more nickel than anadium. Each metal has negative effects.
Nickel (Ni) As discussed in Chapter 3, an FCC catalyst has two parts: « The non-framework structure called matrix • The crystalline structure called zeolite
In contact with the catalyst, nickel deposits on the matrix. Nickel romotes dehydrogenation reactions, removing hydrogen from stable ompounds and making unstable olefins, which can polymerize to eavy hydrocarbons. These reactions result in high hydrogen and coke ields. The higher coke causes higher regenerator temperatures. This owers the catalyst-to-oil ratio and lowers conversion. High nickel levels are normally encountered when processing heavy eed. Neither excess hydrogen nor excess regenerator temperature is esirable. Excess hydrogen lowers the molecular weight of the wet gas; ince the compressor is usually centrifugal, this limits the discharge ressure. Lower pressure means less capacity and this can force a eduction in charge or operation at lower conversion. A number of indices relate metal activity to hydrogen and coke roduction. (These indices predate the use of metal passivation in the CC process but are still reliable). The most commonly used index is x Nickel + Vanadium. This indicates that nickel is four times as ctive as vanadium in producing hydrogen. Other indices [9] used are: Jersey Nickel Equivalent Index = 1,000 x (Ni + 0.2 x V + 0.1 x Fe) Shell Contamination Index = 1,000 x (14 x Ni + 14 x Cu + 4 x V + Fe) V Davison Index = Ni + Cu H— 4 V Mobil = Ni + — 4
4
Fluid Catalytic Cracking Handbook
In every equation, nickel is the most active. These indices convert ll metals to a common basis, generally either vanadium or nickel. Metals are most active when they first deposit on the catalyst. With ime, they lose their initial effectiveness through continuous oxidationeduction cycles. On average, about one third of the nickel on the quilibrium catalyst will have the activity to promote dehydrogenaion reactions, A small amount of nickel in the FCC feed has a significant influence on the unit operation. In a "clean" gas oil operation, the hydrogen yield s about 40 standard cubic feet (scf) per barrel of feed (0.07 wt%). This is a manageable rate that most units can handle. If the nickel evel increases to 1.5 ppm, the hydrogen yield increases up to 100 scf per barrel (0.17 wt%). Note that in a 50,000 barrel/day unit, this orresponds to a mere 16 pounds per day of nickel. Unless the catalyst ddition rate is increased or the nickel in the feed is passivated (see Chapter 3), the feed rate or conversion may need to be reduced. The wet gas will become lean and may limit the pumping capacity of he wet gas compressor. In most units, the increase in hydrogen make does not increase coke yield; the coke yield in a cat cracker is constant (Chapter 5). The coke yield does not go up because other unit constraints, such as the egenerator temperature and/or wet gas compressor, force the operator o reduce charge or severity. High hydrogen yield also affects the ecovery of C3+ components in the gas plant. Hydrogen works as an nert and changes the liquid-vapor ratio in the absorbers. On a wt% basis, the increase in hydrogen is negligible, but the sharp ncrease in gas volume impacts unit performance, Catalyst composition and feed chloride have a noticeable impact on hydrogen yield. Catalysts with an active alumina matrix tend to ncrease the dehydrogenation reactions. Chlorides in the feed reactivate ged nickel, resulting in high hydrogen yield. Two common indicators track the effects of nickel on the catalyst. These are: * Hydrogen/methane ratio * Volume of hydrogen per barrel of feed
The H2/CH4 ratio is an indicator of dehydrogenation reactions. However, the ratio is sensitive to the reactor temperature and the type f catalyst. A better indicator of nickel activity is the volume of
FCC Feed Characterization
85
ydrogen per barrel of fresh feed. The typical H2/CH4 mole ratio for gas oil having less than 0.5 ppm nickel is between 0.25 to 0.35, he equivalent H2 make is between 30 and 40 scf/bbl of feed. It is usually more accurate to back-calculate the feed metals from he equilibrium catalyst data than to analyze the feed regularly. If ickel will be a regular component of the feed, passivators are availble. If nickel affects operation and margins, it is often beneficial to se antimony to passivate the nickel. This can be attractive if the ickel on the equilibrium catalyst is greater than 1,000 ppm.,
anadium
Vanadium also promotes dehydrogenation reactions, but less than nickel. Vanadium's contribution to hydrogen yield is 20% to 50% of nickel's ontribution, but vanadium is a more severe poison. Unlike nickel, anadium does not stay on the surface of the catalyst. Instead, it migrates o the inner (zeolite) part of the catalyst and destroys the zeolite crystal ructure. Catalyst surface area and activity are permanently lost. Vanadium occurs as part of organo-metallic molecules of high molecular weight. When these heavy molecules are cracked, coke esidue containing vanadium is left on the catalyst. During regeneraon, the coke is burned off and vanadium is converted to vanadium xides such as vanadium pentoxide (V2O5). V2O5 melts at 1,274°F 690°C), which allows it to destroy zeolite under typical regenerator emperature conditions. V2O5 is highly mobile and can go from one article to another. There are several theories about the chemistry of vanadium poisonng. The most prominent involves conversion of V2O5 to vanadic acid H3VO4) under regenerator conditions. Vanadic acid, through hydrolysis, xtracts the tetrahedral alumina in the zeolite crystal structure, causing to collapse. The severity of vanadium poisoning depends on the following factors: 1. Vanadium Concentration In general, vanadium concentrations above 2,000 ppm on the E-Cat can justify passivation. 2. Regenerator Temperature Higher regenerator temperatures (>1,250°F or 677°C) exceed the melting point of vanadium oxides, increasing their mobility. This
§
Fluid Catalytic Cracking Handbook
allows vanadium to find zeolite sites. This deactivation is in addition to the hydrothermal deactivation caused by higher regenerator temperature alone. 3. Combustion Mode Regenerators operating in full combustion and producing "clean" catalyst (Figure 2-10) increase vanadium pentoxide formation because of the excess oxygen. 4. Sodium Sodium and vanadium react to form sodium vanadates. These mixtures have a low melting point (<1,200°F or 649°C) and increase vanadium mobility. 5. Steam Steam reacts with V2O5 to form volatile vanadic acid. Vanadic acid, through hydrolysis, causes collapse of the zeolite crystal. 6. Catalyst Type The alumina content, the amount of rare-earth, and the type and amount of zeolite affect catalyst tolerance to vanadium poisoning.
80 4-_________)
0
Figure 2-10.
1000
j.
2000
[_
3000
I
]_
4000
5000
6000
Vanadium, ppm Vanadium deactivation varies with regenerator severity [13].
FCC Feed Characterization
67
1. Catalyst Addition Rate A higher catalyst addition rate dilutes the concentration of metals and allows less time for the vanadium to get fully oxidized,
Alkaline Earth Metals
Alkaline earth metals in general, and sodium in particular, are detrimental to the FCC catalyst. Sodium permanently deactivates the catalyst by neutralizing its acid sites. In the regenerator it causes the zeolite to collapse, particularly in the presence of vanadium. Sodium comes from two prime sources: • Sodium in the fresh catalyst « Sodium in the feed
Fresh catalyst contains sodium as part of the manufacturing process. Chapter 3 discusses the drawbacks of sodium that are inherent in the resh catalyst. Sodium in the feed is called added sodium. For all practical purposes, he adverse effects of sodium are the same regardless of its origin. Sodium usually appears in the form of sodium chloride. Chlorides end to reactivate aged metals on the catalyst and allows them to cause more damage. Sodium originates from the following places: • Caustic that is added downstream of the crude oil desalter. Caustic is injected downstream of the desalter to control overhead corrosion. Natural chloride salts in crude decompose to HC1 at typical unit temperatures. Caustic reacts with these salts to form sodium chloride. Sodium chloride is thermally stable at the temperature found in the crude and vacuum unit heaters. This results in sodium chloride being present in either atmospheric or vacuum resids. Most refiners discontinue caustic injection when they process residue to the FCC unit. It can still be present in purchased feedstocks, however. • Water soluble salts that are carried over from the desalter. An effective desalting operation is more important than ever when processing heavy feedstocks to the cat cracker. Chloride salts are usually water soluble and are removed from raw crude in the desalter. However, some of these salts can be carried over with desalted crude. • Processing of the refinery "slop." A number of refiners process the refinery slop in their desalter. This can adversely affect the
8
Fluid Catalytic Cracking Handbook
desalter and carry over salts with the desalted crude. Slop can be fed to the coker or FCC main fractionator with the same result, « Purchased FCC feedstock can be exposed to salt water as ballast, * The me of atomizing steam and/or water that contain sodium. Just about every refiner practices some type of feed atomization using either steam or water. The steam or water can contain varying amounts of sodium depending on the quality of water treatment used in the refinery.
Another problem associated with sodium appears in the form of odium chloride. Chlorides tend to reactivate aged metals by redisibuting the metals on the equilibrium catalyst and allowing them to ause more damage.
Other Metals
Iron is usually present in FCC feed as tramp iron and is not cataytically active. Tramp iron refers to various corrosion by-products om upstream processing and handling. Copper is as active as nickel, but the feed contains much less of it.
ummary
The metals in the FCC feed have many deleterious effects. Nickel auses excess hydrogen production, forcing eventual loss in the conersion or thruput. Both vanadium and sodium destroy catalyst strucure, causing losses in activity and selectivity. Solving the undesirable ffects of metal poisoning involves several approaches: * * * *
Increasing the makeup rate of fresh catalyst Employing some type of metal passivation Adding good quality equilibrium catalysts to flush the metals Employing demetalization technology to remove metals from the equilibrium catalyst
Both demetalization and passivation technology are addressed in hapter 3.
The typical refinery laboratory is not equipped to conduct PONA nd other chemical analyses of the FCC feed on a routine basis.
FCC Feed Characterization
§9
However, physical properties such as °API gravity and distillation are easy to measure. As a result, empirical correlations have been developed by the industry to determine chemical properties from these physical analyses. Characterizing FCC feed provides quantitative and qualitative estimates of the FCC unit's performance. Process modeling uses the feed properties to predict FCC yields and product qualities. The process model should be used in daily unit monitoring, catalyst evaluations, optimization, and process studies. There are no standard correlations. Some companies have proprietary correlations, but this does not mean that these correlations do a better ob at predicting yields. Nonetheless, they all incorporate most or some of the same physical properties. The most widely published correlations in use today are: • • • •
K Factor TOTAL n-d-M Method API Method
K Factor
The K factor is a very useful indication of feed crackability. The K factor relates to the hydrogen content of the feed. It is normally calculated using feed distillation and gravity data, and measures aromaticity relative to paraffinicity. Higher K values indicate increased paraffinicity and more crackability. A K value above 12.0 indicates a paraffinic feed; a K value below 11.0, aromatic. Like aniline point, the K factor differentiates between the highly paraffinic and aromatic stocks. However, within the narrow range (K = 11.5-12.0), the K factor does not correlate between aromatics and naphthenes. Instead, it relates fairly well to the paraffin content (Figure 2-11). The K factor does not provide information as to the ratio of naphthene and paraffin contents. The ratio of naphthenes to paraffins can vary considerably with the same K values (Table 2-8). K value is the ratio of the cube root of a boiling temperature to gravity. There are two widely used methods to calculate the K factor: Kw and the Kuop. The equations used for calculating both factors are as follows:
0
Fluid Catalytic Cracking Handbook 64
60 S I (8 O.
66
52
11.4
11.6
11.8
12
UOP K Factor Figure 2-11. Weight percent paraffins at various UOP K factors.
(MeABP + 46Q)1/3 SG
(CABP + 460)1/3 SG
FCC Feed Characterization
71
(VABP + 460)1/3 SG
KUOP
w A un
(T(10%) + T(30%) + T(50%) + T(70%) + T(90%)) :
V AJar =
= £(F vi xTB! /3 ) 3
Where:
MeABP SG CABP VABP MABP
mj
TBj
vj
T
= = = = = = = = =
Mean Average Boiling Point, °F Specific Gravity at 60°F (see Equation 2-1) Cubic Average Boiling Point, °F Volumetric Average Boiling Point, °F Molar Average Boiling Point, °F Mole Fraction of Component i Normal boiling point of pure component i, °F Volume fraction of component i Temperature, °F
Table 2-8 Variation of C^CP as a Function of Kuop Factor*
Sample No,
1 2 3 4 5 6 7
K
uop Factor
CA + CN (wt%)
CN/CP
11.70 11.69 11.70 11.67 11.70 11.70 11.70
46 45 46 45 45 44 42
0.47 0.44 0.44 0.43 0.39 0.35 0,33
The K factor relates well to aromatics + naphthenes, but not to naphthenes. ource: Andreasson [10]
2
Fluid Catalytic Cracking Handbook
The UOP method uses CABP which, for all practical purposes, is he same as VABP as shown in Appendix 2. The K factor is more opular than Kw because the VABP data are readily available. The use f MeABP in the Watson method generally results in a lower K value han that of UOP. Example 2-1 illustrates steps to calculate the K nd K,,, factors.
Example 2-1
Determine K UOP and Watson Kw using the following FCC feed roperties: Feed Properties
API Gravity G ensity efractive Index iscosity
@ @ @ @ @ @
60°F 60°F 20°C (68°F) 67°C 130°F, SUS 210°F, SUS
ulfur, wt% niline Point, °F(°C)
23.5 0.913 0.90 1.4810 137.0 50.0 0.48 192.0 (88.9)
D-1160 @ 1 atm Vol%
Temp. °F
652 751 835 935 1080
10 30 50 70 90 Procedure
Calculate VABP from distillation data. Calculate the 10%-90% slope. Calculate MeABP and CABP by adding corrections from Appendix 2 to VABP.
tep 1: VABP = 1/5(652 + 751 + 835 + 935 + 1080) VABP = 851°F = 455°C = 728.2°K
FCC Feed Characterization
73
Step 2: 10%-90% slope Mope =
T(90%) - T(10%) 1080 - 652 = 80 80
Slope = 5.3°F/percent off
Step 3: From Appendix 2, corrections to VABP are approximately -34°F for MeABP and -10°F for CABP. Therefore: MeABP = 851 - 34 = 817°F = 436°C CABP = 851 - 10 = 841°F = 449.4°C
A K (817 + 460)1/3 Step 4: Kw = = 11.88
0.913
nstead of using Appendix 2, the MeABP can be determined from the quation below [11]: (
MeABP = VABP + 2 -
(T _T )
^
^
\
U 70+0.075 x VABP J
+1.53
MeABP = 851 + 2-f I080"852 +1.5 (170 + 0.075x851 MeABP = 816°F (435°C)
n the absence of full distillation data, the K factor can be estimated sing the 50% point in place of MeABP.
In summary, the K factor can provide information about the aromaticity or paraffinicity of the feed. However, within the narrow range (K = 11.5-12.0), it cannot differentiate between ratio of paraffins, naphthenes, nd aromatics. To determine these ratios, other correlations, such as TOTAL or n-d-M, should be employed.
4
Fluid Catalytic Cracking Handbook
TOTAL
The TOTAL correlations calculate aromatic carbon content, hydrogen ontent, molecular weight, and refractive index using routine laboratory ests. The TOTAL correlations are listed below and are also in Appenix 3, Example 2-2 illustrates the use of TOTAL correlations. Example 2-2
Molecular Weight (MW) MW = 0.0078312 x (SG)"0'0978 x (AP °C)OJ238 x (VABP °C)L6971 MW = 0,0078312 x (0.913)"0'0978 x (88.9)0*1238 x {455)L6971 MW = 446
Refractive Index (RI) @ 20°C (68°F) RI f20) = 1 + 0.8447 x (SG)1'2058 x (VABP °K)-°'0557 x (MW)~ao°44 RI(20> = 1 + 0.8447 x (0.913)1'2058 x (728.2)"0'0557 x (446y°M44 RI(20) =1.5105
Refractive Index (RI) @ 60°C (140°F) RI(60) = 1 + 0.8156 x (SG)L2392 x (VABP °K)'a0576 x (MW)'0'0007 RI(60) = 1 + 0.8156 x (0.913)L2392 x (728.2)"0'0576 x (446)"0'0007 Rl{60) = 1.4963
Hydrogen (H2) Content, wt% H2 = 52.825 - 14.26 x RI(20) - 21.329 x (SG) - 0.0024 x (MW) - 0.052 x (S) + 0.757 x In (v) H, = 52.825 - 14.26 x 1.5105 - (21.329 x 0.913) - (0.0024 x
FCC Feed Characterization
75
446) - (0.052 x 0.48} - (0.872 - In (7.37)) H2 = 12.23 wt%
Aromatic (CA) Content, wt% CA = -814.136 + (635.192 x RI(20)) - (129.266 x (SG)) + (0.013 x (MW)) - (0.34 x (S)) + (0.872 x In (v)) CA = -814.136 + (635.192 x 1.5105) - (129.266 x 0.913) + (0.013 x 446) - (0.34 x 0.48) + (0.872 x In 7.37) CA = 19.19 wt%
Where:
SG AP °C VABP °C VABP °K S V
= Specific gravity at 20°C = Aniline Point, °C = Volumetric Average Boiling Point, °C = Volumetric Average Boiling Point, °K = Sulfur, wt% = Viscosity at 100°C
For FCC feeds, particularly the ones containing residue, the TOTAL correlation is more accurate at predicting aromatic carbon content than he n-d-M correlation. Table 2-9 illustrates this comparison. One option s to calculate MW, RI(20)» CA, and H2 from the TOTAL correlation, and use either the n-d-M or API method to calculate the wt% naphthene CN) and wt% paraffin (Cp).
n-d-M Method
The n-d-M correlation is an ASTM (D-3238) method that uses efractive index (n), density (d), average molecular weight (MW), and ulfur (S) to estimate the percentage of total carbon distribution in the aromatic ring structure (% CA), naphthenic ring structure (CN), and paraffin chains (% Cp). Both refractive index and density are either measured or estimated at 20°C (68°F). Appendix 4 shows formulas used to calculate carbon distribution. Note that the n-d-M method alculates, for example, the percent of carbon in the aromatic ring
6
Fluid Catalytic Cracking Handbook Table 2-9 Comparison of TOTAL Correlations with Other Methods
Correlation
arbon Content (%C) n-d-M API TOTAL Hydrogen Content (%H) Linden Fein-Wilson-Winn Modified Winn TOTAL Molecular Weight (MW) API Maxwell Kesler-Lee TOTAL efractive Index (RI) API @ 20°C Lindee-Whitter @ 20°C TOTAL @ 20°C TOTAL @ 60°C
Absolute Average Deviation
Bias Maximum Deviation
5.14 2.88 0.93
4.67 2.53 0.00
12.99 9.13 3.45
0.31 0.36 0.19 0.10
-0.05 0.19 0.07 0.00
1.57 1.43 0.86 0.42
Average Deviation
62.0 63.3 61.5 10.6
-62.0 -63.6 -61.1 -0.20
180.9 175.0 176.9 44.4
0.0368
-0.0367
0.0993
0.0315 0.0021 0.0021
-0.0131 0.0 0.0
0.0303 0.0074 0.0074
ource: Dhuleaia [1]
tructure. For instance, if there was a toluene molecule in the feed, he n-d-M method predicts six aromatic carbons (86%) versus the ctual seven carbons. ASTM D-2502 is one of the most accurate methods of determining molecular weight. The method uses viscosity measurements; in the bsence of viscosity data, molecular weight can be estimated using the OTAL correlation. The n-d-M method is very sensitive to both refractive index and ensity. It calls for measurement or estimation of the feed refractive ndex at 20°C (68°F). The problem is that the majority of FCC feeds re virtually solid at 20°C and the refractometer is unable to measure
FCC Feed Characterization
77
he refractive index at this temperature. To use the n-d-M method, efractive index at 20°C needs to be estimated using published corelations. For this reason the n-d-M method is usually employed in onjunction with other correlations such as TOTAL. Example 2-3 can be used to illustrate the use of the n-d-M correlations.
Example 2-3
Using the feed property data in Example 2-1, determine MW, CA, CN nd Cp using the n-d-M method.
Step 1: Molecular weight determination by ASTM method. 1. Obtain viscosity at 100°F (37.8°C) a. Plot viscosities at 130°F (54.4°C) and 210°F (98.89°C). b. Extrapolate to 100°F, VIS = 279 SSU. 2. Convert viscosities from SUS to centistoke (csT): a. From Appendix 6, viscosity @ 100°F = 60.0 cSt. b. Viscosity @ 210°F = 7.37 cSt. 3. Obtain molecular weight: a. From Appendix 5, H = 372 and MW = 430.
Step 2: Calculate refractive index @ 20°C from the TOTAL correlation. RI(20) = 1 + 0.8447 x (SO)1'2056 x (VABP(deg C) + 273.16r°'0557 x (MW)-0'0044 RI(20) = 1 + 0.8447 x (0.913)!'2056 x 728.2^'0557 x 446^0044 RI,20) = 1.5046
Step 3: Calculate n-d-M Factors. V = 2.51 x (RI(20) - 1.4750) - (D20 - 0.8510) = 0.0271 V = 2.51 x (1.5046 - 1.4750) - (0.90 - 0.8510) = +0.0271 w = (D20 - 0.8510) - 1.11 (RI(20J - 1.4750) = +0.0226 w - (0.90 - 0.8510) - 1.11 x (1.5046 - 1.4750) = +0.0226
8
Fluid Catalytic Cracking Handbook Because V is positive: %CA = 20.16
Because w is positive: %CR =820xw-(3xS) + -
10,000
MW 10,000
%CR = 820 x 0.0226 - 3 x 0.48 +
430
The API method is a generalized method that predicts mole fraction f paraffinic, naphthenic, or aromatic compounds for an olefin-free ydrocarbon. The development of the equations is based on dividing he hydrocarbon into two molecular ranges: heavy fractions (200 < MW < 600) and light fractions (70 < MW <200). Appendix 7 contains API correlations applicable to the FCC feed. Example 2-4
Use the feed property data in Example 2-1 to calculate MW, RI(20r XA, XN, and Xp, employing API correlations (see Appendix 7). Calculate MW: MW = a x exp(b x MeABP + C x SG + d x MeABP x SG) x (MeABP)e x (SG/
Where:
onstants a b c d e f
= = = = = =
20.486 1.165 xlO"4 -7.787 1.1582xlO"3 1.26807 4.98308
FCC Feed Characterization
79
Calculate refractive index: I = A x exp(B x MeABP + C x SG + D x MeABP x SG) x MeABP*1 x SGF 1 = 9 341 v 10~2 xexD^'4Wxlo~4x!'276+5pl4^x0'913~3'289xlcr4xl>276x0"913) x(!276)-°-407x(0.913)-3-3333
RI = (1 + 2 x I/I- I) l/2 /
\l/2
_ 1 + 2x0.294 V R1(20) "~l 1-0.294 J RJ(20) = 1 .500
,,^ ,7. . ^ . „ SG - 0.24- 0.222 xlog(v210- 35.5) VG = Viscosity Gravity Constant = ™ 0.755 Q=
0.913 -0.24- 0.022 xlog(50- 35.5) 0.755
VG = 0.8575 XA = g + h(Ri) + i(VG)
XA = -403.8 + 265.7 xf 1.5000-^^1 + 161. Ox 0.8575 XA = 1 1 .5 rnol% XN = d + e(Ri) + f(VG)
XN = 246.4 - 367.0 x 1.5000 XN = 31.8 mol% Xp = a + b(Ri) + i(VG)
2
+ 196.3 x 0.8575
0
Fluid Catalytic Cracking Handbook Xp = 257.37 + 101.33(Ri) + 160.988(0.8575)
Xp = 56,7 mo.1%
Where:
onstants a b c d e f g h i
= = = = = = = =
+2.5737 +1.0133 -3.573 +2.464 -3.6701 +1.96312 4.0377 +2.6568 +1.60988
The findings from TOTAL, n-d-M, and API are summarized in Table -10. The comparison illustrates how sensitive the predicted feed omposition is to the refractive index @ 20°C. For instance, using the OTAL correlation, there is a 35% drop in the aromatic content in sing RI(20) = 1.5000 instead of RI(20) = 1.5105. When using these orrelations, every effort should be made to obtain accurate and onsistent values for the refractive index at 20°C. With the refractive
Table 2-10 Comparison of the Findings Among the 3 Correlations API
n-d-M
efractive Index @ 20°C Molecular Weight
1.5000
413
430
arbon Content:
Mol%
Wt%
Aromatic Naphthene Paraffin
11.5, (14.3)* 31.8, (27.9)* 56.7, (57.8)*
TOTAL
1.5105
446 Wt% t
(20.2)*,(8.8) (20.2)*,(41.1)f (59.6)*,(50.1)t
Uses Rll2i)l from n-d-M correlation to determine composition, Uses RI^0ifrorn API correlation to determine composition.
19.2, (12.5)*
FCC Feed Characterization
81
ndex at any given temperature, the RI(2o) can be calculated from the ollowing equation. Example 2-5 illustrates the use of the equation. RI (2m = RI(t) + 6.25 x (t - 20) x 10"4 t = temp, °C
Example 2-5
With the refractive index @ 78°C = 1.4810, determine the refracive index @ 20°C. RI(2()) = 1.4810 + 6.25 x (67 - 20) x 10"4 RI(20) = 1.5104
Note that the calculated RI(20) closely matches that using the TOTAL correlation.)
Pretreatment of FCC feedstock through hydroprocessing has a number of benefits including: • • • • •
Hydrodesulfurization (HDS) Hydrodenitrogenation (HDN) Hydrodemetallization (HDM) Aromatic Reduction Conradson Carbon Removal
Desulfurization of FCC feedstocks reduces the sulfur content of FCC products and SOX emissions. In the United States, road diesel sulfur can be 500 ppm (0.05 wt%). In some European countries, for example n Sweden, the sulfur of road diesel is 50 ppm or less. In California, he gasoline sulfur is required to be less than 40 ppm. The EPA's complex model uses sulfur as a controlling parameter to reduce toxic missions. With hydroprocessed FCC feeds, about 5% of feed sulfur s in the FCC gasoline. For non-hydroprocessed feeds, the FCC gasoline sulfur is typically 10% of the feed sulfur.
2
Fluid Catalytic Cracking Handbook
The nitrogen compounds in the FCC feed deactivate the FCC atalyst activity resulting in an increase in coke and dry gas. Hydroenitrogenation reduces nitrogen compounds in FCC feeds. In the egenerator, the nitrogen and the attached heterocyclic compounds add nwanted heat to the regenerator causing a low unit conversion. Hydrodemetallization reduces the amount of nickel and, to a lesser xtent, vanadium in FCC feeds. Nickel dehydrogenates feed to molecular ydrogen and aromatics. Removing these metals allows heavier gas il cut points. Polynuclear aromatics (PNA) do not react in the FCC and tend to emain in coke. Adding hydrogen to the outer ring clusters makes them more crackable and less likely to form coke on the catalyst. Hydroprocessing reduces the Conradson carbon residue of heavy ils, Conradson carbon residue becomes coke in the FCC reactor. This xcess coke must be burned in the regenerator, increasing regenerator ir requirements.
It is important to characterize FCC feeds as to their molecular tructure. Once the molecular configuration is known, kinetic models an be developed to predict product yields. The simplified correlations bove do a reasonable job of defining hydrocarbon type and distribuon in FCC feeds. Each correlation provides satisfactory results within he range for which it was developed. Whichever correlation is used, he results should be trended and compared with unit operation. A clear understanding of feed physical properties is essential to uccessful work in the areas of troubleshooting, catalyst selection, unit ptimization, and any planned revamp.
REFERENCES
1. Dhulesia, H., "New Correlations Predict FCC Feed Characterizing Parameters," Oil & Gas Journal, January 13, 1986, pp. 51-54 ASTM, "Standard Test Method for Calculation of Carbon Distribution and Structural Group Analysis of Petroleum Oils by the n-d-M Method," ASTM Standard D-3238-85, 1985. 3. Riazi, M. R., and Daubert, T. E., "Prediction of the Composition of Petroleum Fractions," Ind. Eng. Chem. Process Dev., Vol. 19, No. 2, 1982, pp. 289-294.
FCC Feed Characterization
83
4. ASTM, "Standard Test Method for Estimation of Molecular Weight (Relative Molecular Mass) of Petroleum Oils from Viscosity Measurements," ASTM Standard D-2502-92, 1992, 5. Flanders, R. L., Proceedings of the 35th Annual NPRA Q&A Session on Refining and Petrochemical Technology, Philadelphia, Pa., 1982, p. 59. 6. Wollaslon, E. G., Forsythe, W. L., and Vasalos, I. A., "Sulfur Distribution in FCC Products," Oil & Gas Journal, August 2, 1971, pp. 64-69. 7. Huling, G. P., McKinney, J. D., and Readal, T, C, "Feed-Sulfur Distribution in FCC Products," Oil & Gas Journal, May 19, 1975, pp. 73-79. 8. Campagna, R. J., Krishna, A. S., and Yanik, S. J., "Research and Development Directed at Resid Cracking," Oil and Gas Journal, October 31, 1983, pp. 129-134. 9. Davison Div., W. R. Grace & Co., "Questions Frequently Asked About Cracking Catalyst," Grace Davison Catalagram, No. 64, 1982, p. 29. 0. Andreasson, H. U. and Upson, L. L., "What Makes Octane," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 2223, 1985. 1. Van, K. B., Gevers, A., and Blum, A., "FCC Unit Monitoring and Technical Service," presented at 1986 Akzo Chemicals Symposium, Amsterdam, The Netherlands. 2. Scherzer, J., and McArthur, D. P., "Nitrogen Resistance of FCC Catalysts," presented at Katalistiks' 8th Annual FCC Symposium, Venice, Italy, 1986. 3. Dougan, T. J., Alkemade, V, Lakhampel, B., and Brock, L. T., "Advances in FCC Vanadium Tolerance," presented at NPRA Annual Meeting, San Antonio, Texas, March 20, 1994; reprinted in Grace Davison Catalagram No. 72, 1985.
CHAPTER 3
FCC Catalysts
The introduction of zeolite in commercial FCC catalysts in the early 1960s was one of the most significant advances in the history of cat racking. Zeolite catalysts provided a greater profit with little capital nvestment. Simply stated, zeolite catalysts were and still are the iggest bargain of all time for the refiner. Improvements in catalyst echnology have continued, enabling refiners to meet the demands of heir market with minimum capital investment. Compared to amorphous silica-alumina catalysts, the zeolite catalysts re more active and more selective. The higher activity and selectivity ranslate to more profitable liquid product yields and additional crackng capacity. To take full advantage of the zeolite catalyst, refiners ave revamped older units to crack more of the heavier, loweralue feedstocks. A complete discussion of FCC catalysts would fill another book. This chapter provides enough information to select the proper catalyst nd to troubleshoot the unit's operation. The key topics discussed are: • • • • • • •
Catalyst Components Catalyst Manufacturing Techniques Fresh Catalyst Properties Equilibrium Catalyst Analysis Catalyst Management Catalyst Evaluation Additives
CATALYST COMPONENTS
FCC catalysts are in the form of fine powders with an average article size in the range of 75 microns. A modern cat cracking catalyst as four major components: • Zeolite • Matrix 84
FCC Catalysts
85
* Binder • Filler
Zeolite
Zeolite, or more properly, faujasite, is the key ingredient of the FCC atalyst. It provides product selectivity and much of the catalytic ctivity. The catalyst's performance largely depends on the nature and uality of the zeolite. Understanding the zeolite structure, types, racking mechanism, and properties is essential in choosing the "right" atalyst to produce the desired yields.
Zeolite Structure
Zeolite is sometimes called molecular sieve. It has a well defined attice structure. Its basic building blocks are silica and alumina etrahedra (pyramids). Each tetrahedron (Figure 3-1) consists of a ilicon or aluminum atom at the center of the tetrahedron, with oxygen toms at the four corners. Zeolite lattices have a network of very small pores. The pore diameter f nearly all of today's FCC zeolite is approximately 8.0 angstroms (°A). These small openings, with an internal surface area of roughly 600 square
Figure 3-1.
Silicon/aluminum-oxygen tetrahedron [15].
6
Fluid Catalytic Cracking Handbook
meters per gram, do not readily admit hydrocarbon molecules that have molecular diameter greater than 8.0°A to 10°A. The elementary building block of the zeolite crystal is a unit cell. The unit cell size (UCS) is the distance between the repeating cells n the zeolite structure. One unit cell in a typical fresh Y-zeolite lattice ontains 192 framework atomic positions: 55 atoms of aluminum and 137 atoms of silicon. This corresponds to a silica (SitX,) to alumina A12O3) molal ratio (SAR) of 5. The UCS is an important parameter n characterizing the zeolite structure.
Zeolite Chemistry
As stated above, a typical zeolite consists of silicon and aluminum toms that are tetrahedrally joined by four oxygen atoms. Silicon is n a +4 oxidation state; therefore, a tetrahedron containing silicon is eutral in charge. In contrast, aluminum is in a +3 oxidation state. This ndicates that each tetrahedron containing aluminum has a net charge f -1, which must be balanced by a positive ion. Solutions containing sodium hydroxide are commonly used in ynthesizing the zeolite. The sodium serves as the positive ion to alance the negative charge of aluminum tetrahedron. This zeolite is alled soda Y or NaY. The NaY zeolite is not hydrothermally stable ecause of the high sodium content. The ammonium ion is frequently sed to displace sodium. Upon drying the zeolite, ammonia is vaporized. The resulting acid sites are both the Bronsted and Lewis types. The Bronsted acid sites can be further exchanged with rare earth material, uch as cerium and lanthanum to enhance their strengths. The zeolite ctivity comes from these acid sites.
eolite Types
Zeolites employed in the manufacture of the FCC catalyst are ynthetic versions of naturally occurring zeolites called faujasites. There are about 40 known natural zeolites and over 150 zeolites that ave been synthesized. Of this number, only a few have found commercial pplications. Table 3-1 shows properties of the major synthetic zeolites. The zeolites with applications to FCC are Type X, Type Y, and ZSM-5. Both X and Y zeolites have essentially the same crystalline tructure. The X zeolite has a lower silica-alumina ratio than the Y eolite. The X zeolite also has a lower thermal and hydrothermal
FCC Catalysts
8?
Table 3-1 Properties of Major Synthetic Zeolites
Zeoiite Type
Pore Size Dimensions (°A)
Silica-toAlumina Ratio
Zeolite A Faujasite ZSM-5
4.1 7.4 5.2 x 5.8
2-5 3-6 30-200
Mordenite
6.7 x 7.0
10-12
Applications Detergent manufacturing Catalytic cracking and hydrocracking Xylene isomerization, benzene alkylation, catalytic cracking, catalyst dewaxing, and methanol conversion. Hydro-isomerization, dewaxing
tability than the Y zeolite. Some of the earlier FCC zeolite catalysts contained X zeolite; however, virtually all of today's catalysts contain Y zeolite or variations thereof (Figure 3-2). ZSM-5 is a versatile zeolite that increases olefin yields and octane. ts application is further discussed later in this chapter. Until the late 1970s, the NaY zeolite was mostly ion exchanged with are earth components. Rare earth components, such as lanthanum and
USY Zeolite (~ 7 Al Atoms/u.c.)
nit Cell Dimension =24.25 A (SiO2/AI2O3=54)
Equilibrium REY (-23 Al Atoms/u.c.) Unit Cell Dimension = 24.39 A (SiO2/AI2O3 « 15)
Figure 3-2. Geometry of USY and REY zeolites [14].
8
Fluid Catalytic Cracking Handbook
erium, were used to replace sodium in the crystal. The rare earth lements, being trivalent, simply form "bridges" between two to three cid sites in the zeolite framework. Bridging protects acid sites from eing ejected and stabilizes the zeolite structure. Rare earth exchange dds to the zeolite activity and thermal and hydrothermal stability. The reduction of lead in motor gasoline in 1986 created the need or a higher FCC gasoline octane. Catalyst manufacturers responded y adjusting the zeolite formulations, an alteration that involved xpelling a number of aluminum atoms from the zeolite framework. The removal of aluminum increased SAR, reduced UCS, and in the rocess, lowered the sodium level of the zeolite. These changes ncreased the gasoline octane by raising its olefinicity. This aluminumeficient zeolite was called ultrastable Y, or simply USY, because of s higher stability than the conventional Y.
eolite Properties
The properties of the zeolite play a significant role in the overall erformance of the catalyst. Understanding these properties increases ur ability to predict catalyst response to changes in unit operation. rom its inception in the catalyst plant, the zeolite must retain its atalytic properties under the hostile conditions of the FCC operation. The reactor/regenerator environment can cause significant changes in hemical and structural composition of the zeolite. In the regenerator, or instance, the zeolite is subjected to thermal and hydrothermal reatments. In the reactor, it is exposed to feedstock contaminants such s vanadium and sodium. Various analytical tests determine zeolite properties. These tests upply information about the strength, type, number, and distribution f acid sites. Additional tests can also provide information about urface area and pore size distribution. The three most common arameters governing zeolite behavior are as follows: • Unit Cell Size • Rare Earth Level « Sodium Content
Unit Cell Size (UCS). The UCS is a measure of aluminum sites or he total potential acidity per unit cell. The negatively-charged aluminum toms are sources of active sites in the zeolite. Silicon atoms do not
FCC Catalysts
89
possess any activity. The UCS is related to the number of aluminum toms per cell (N Af ) by [1]: NA, + 111 x (UCS - 24.215)
The number of silicon atoms (Nsi) is; Nsi = 192 - NA,
The SAR of the zeolite can be determined either from the above two quations or from a correlation such as the one shown in Figure 3-3. The UCS is also an indicator of zeolite acidity. Because the alumium ion is larger than the silicon ion, as the UCS decreases, the acid ites become farther apart. The strength of the acid sites is determined y the extent of their isolation from the neighboring acid sites. The lose proximity of these acid sites causes destabilization of the zeolite tructure. Acid distribution of the zeolite is a fundamental factor ffecting zeolite activity and selectivity. Additionally, the UCS measurement can be used to indicate octane potential of the zeolite. A lower UCS presents fewer active sites per unit cell. The fewer acid ites are farther apart and, therefore, inhibit hydrogen transfer reactions, which in turn increase gasoline octane as well as the production of C3 and lighter components (Figure 3-4). The octane increase is due o a higher concentration of olefins in the gasoline. Zeolites with lower UCS are initially less active than the conventional rare earth exchanged zeolites (Figure 3-5). However, the ower UCS zeolites tend to retain a greater fraction of their activity under severe thermal and hydrothermal treatments, hence the name ultrastable Y. A freshly manufactured zeolite has a relatively high UCS in the ange of 24,50°A to 24.75°A. The thermal and hydrothermal environment of the regenerator extracts alumina from the zeolite structure and, herefore, reduces its UCS. The final UCS level depends on the rare arth and sodium level of the zeolite. The lower the sodium and rare arth content of the fresh zeolite, the lower UCS of the equilibrium atalyst (E-cat).
Rare Earth Level. Rare earth (RE) elements serve as a "bridge" o stabilize aluminum atoms in the zeolite structure. They prevent the
0
Fluid Catalytic Cracking Handbook
Figure 3-3.
Silica-alumina ratio versus zeolite unit cell size,
luminum atoms from separating from the zeolite lattice when the atalyst is exposed to high temperature steam in the regenerator. A fully rare-earth-exchanged zeolite equilibrates at a high UCS, whereas a non-rare-earth zeolite equilibrates at a very low UCS in the ange of 24.25 [3]. All intermediate levels of rare-earth-exchanged eolite can be produced. The rare earth increases zeolite activity and
FCC Catalysts
24.24
24.28
91
24.32
24.36
24.32
24.36
Unit Cell Size, A
6.0 5.5
>s t
1 5.0 o 4.5 «*>
4.0
24.20
24.24
24.28 Unit Cell Size, A
Figure 3-4.
Effects of unit cell size on octane and C3-gas make [4].
2
Fluid Catalytic Cracking Handbook
90 80
1520°F,20% steam in air.
0
10
20
30
40
50
60
70
80
90 100
Time, hrs
Figure 3-5. Comparison of activity retention between rare-earth-exchanged eolites versus USY zeolites. (Source: Grace Davison Octane Handbook.)
gasoline selectivity with a loss in octane (Figure 3-6). The octane loss s due to promotion of hydrogen transfer reactions. The insertion of rare arth maintains more and closer acid sites, which promotes hydrogen ransfer reactions. In addition, rare earth improves thermal and hydrohermal stability of the zeolite. To improve the activity of a USY zeolite, he catalyst suppliers frequently add some rare earth to the zeolite.
Sodium Content. The sodium on the catalyst originates either from eolite during its manufacture or from the FCC feedstock. It is important or the fresh zeolite to contain very low amounts of sodium. Sodium decreases the hydrothermal stability of the zeolite. It also eacts with the zeolite acid sites to reduce catalyst activity. In the egenerator, sodium is mobile. Sodium ions tend to neutralize the trongest acid sites. In a dealuminated zeolite, where the UCS is low 24.22°A to 24.25°A), the sodium can have an adverse affect on the gasoline octane (Figure 3-7). The loss of octane is attributed to the drop in the number of strong acid sites. FCC catalyst vendors are now able to manufacture catalysts with a odium content of less than 0.20 wt%. Sodium is commonly reported as
FCC Catalysts
5
§3
Yield of Gasoline, % ~ I
4
3
2
Gasoline octane (R+MV2
1
0 0
2
4
6
8
10
12
Rare Earth, wt% Figure 3-6. Effects of rare earth on gasoline octane and yield.
he weight percent of sodium or soda (Na2O) on the catalyst. The proper way to compare sodium is the weight fraction of sodium in the zeolite, This is because FCC catalysts have different zeolite concentrations. UCS, rare earth, and sodium are just three of the parameters that re readily available to characterize the zeolite properties. They proide valuable information about catalyst behavior in the cat cracker. f required, additional tests can be conducted to examine other eolite properties.
Matrix
The term matrix has different meanings to different people. For ome, matrix refers to components of the catalyst other than the eolite. For others, matrix is a component of the catalyst aside from he zeolite having catalytic activity. Yet for others, matrix refers to he catalyst binder. In this chapter, matrix means components of the atalyst other than zeolite and the term active matrix means the omponent of the catalyst other than zeolite heaving catalytic activity.
4
Fluid Catalytic Cracking Handbook MOTOR OCTANE VS. SODIUM OXIDE 81.5
-
,
;
0.3
0.4
0.5
81.0
O 5 80.5
80.0 0.2
0.6
Na2O, wt% on catalyst
RESEARCH OCTANE VS. SODIUM OXIDE
94
93
92
91
0
1
2
3
4
5
Na2O, wt% on zeolite
igure 3-7. Effects of soda on motor and research octanes: motor octane s. sodium oxide [11]; research octane vs. sodium oxide [4].
FCC Catalysts
95
Alumina is the source for an active matrix. Most active matrices used n FCC catalysts are amorphous. However, some of the catalyst suppliers ncorporate a form of alumina that also has a crystalline structure. Active matrix contributes significantly to the overall performance f the FCC catalyst. The zeolite pores are not suitable for cracking f large hydrocarbon molecules generally having an end point > 900°F 482°C); they are too small to allow diffusion of the large molecules o the cracking sites. An effective matrix must have a porous structure o allow diffusion of hydrocarbons into and out of the catalyst, An active matrix provides the primary cracking sites. The acid sites ocated in the catalyst matrix are not as selective as the zeolite sites, ut are able to crack larger molecules that are hindered from entering he small zeolite pores. The active matrix precracks heavy feed moleules for further cracking at the internal zeolite sites. The result is a ynergistic interaction between matrix and zeolite, in which the activity ttained by their combined effects can be greater than the sum of their ndividual effects [2]. An active matrix can also serve as a trap to catch some of the anadium and basic nitrogen. The high boiling fraction of the FCC eed usually contains metals and basic nitrogen that poison the zeolite. One of the advantages of an active matrix is that it guards the zeolite rom becoming deactivated prematurely by these impurities.
Filler and Binder
The filler is a clay incorporated into the catalyst to dilute its activity. Kaoline [Al2(OH)2, Si2O5] is the most common clay used in the CC catalyst. One FCC catalyst manufacturer uses kaoline clay as a keleton to grow the zeolite in situ. The binder serves as a glue to hold the zeolite, matrix, and filler ogether. Binder may or may not have catalytic activity. The importance f the binder becomes more prominent with catalysts that contain high oncentrations of zeolite. The functions of the filler and the binder are to provide physical ntegrity (density, attrition resistance, particle size distribution, etc.), heat transfer medium, and a fluidizing medium in which the more mportant and expensive zeolite component is incorporated. In summary, zeolite will effect activity, selectivity, and product uality. An active matrix can improve bottoms cracking and resist
6
Fluid Catalytic Cracking Handbook
vanadium and nitrogen attacks. But a matrix containing very small ores can suppress strippablity of the spent catalyst and increase ydrogen yield in the presence of nickel. Clay and binder provide hysical integrity and mechanical strength.
The manufacturing process of modern FCC catalyst is divided into wo general groups—incorporation and "in-situ" processes. All catalyst uppliers manufacture catalyst by an incorporation process that requires making zeolite and matrix independently and using a binder to hold hem together. In addition to the incorporation process, Engelhard also manufactures FCC catalyst using an "in-situ" process in which the eolite component is grown within the pre-formed miscrospheres. The ollowing sections provide a general description of zeolite synthesis.
Conventional Zeolite (KEY, REHY, HY)
NaY zeolite is produced by digesting a mixture of silica, alumina, nd caustic for several hours at a prescribed temperature until crystalization occurs (Figure 3-8). Typical sources of silica and alumina are odium silicate and sodium aluminate. Crystallization of Y-zeolite ypically takes 10 hours at about 210°F (100°C). Production of a uality zeolite requires proper control of temperature, time, and pH f the crystallization solution. NaY zeolite is separated after filtering nd water-washing of the crystalline solution. A typical NaY zeolite contains approximately 13 wt% Na2O. To nhance activity and thermal and hydrothermal stability of NaY, the odium level must be reduced. This is normally done by the ion xchanging of NaY with a medium containing rare earth cations and/ r hydrogen ions. Ammonium sulfate solutions are frequently employed s a source for hydrogen ions. At this state of the catalyst synthesis there are two approaches for urther treatment of NaY. Depending on the particular catalyst and the atalyst supplier, further treatment (rare earth exchanged) of NaY can e accomplished either before or after its incorporation into the matrix. Post-treatment of the NaY zeolite is simpler, but may reduce ion xchange efficiency.
Filtrate to waste treatment
Figure 3-8, Typical manufacturing steps to produce FCC catalyst.
Na^aoHte rystallization 00 F, 12-24 Hr
8
Fluid Catalytic Cracking Handbook
USY Zeolite
An ultrastable or a dealuminated zeolite (USY) is produced by eplacing some of the aluminum ions in the framework with silicon. The conventional technique (Figure 3-9) includes the use of a high emperature (1,300-1,500°F [704-816°C]) steam calcination of HY zeolite, (13%Na O,A,c
NAY
NHf4
2
24.68 A)
- EXCHANGES
NHY
I
(3%Na£ 90)
STEAM CALCINE 114OO DEG. F
USY
JVH.+
(3%Na O,A0* 24,50A!
. EXCHANGES
LOW-SODA USY
(
Figure 3-9. Synthesis of USY zeolite (NAY),
FCC Catalysts
9§
Acid leaching, chemical extraction, and chemical substitution are all orms of dealumination that have become popular in recent years. The main advantage of these processes over conventional dealumination s the removal of the nonframework or occluded alumina from the zeolite cage structure. A high level of occluded alumina residing in he crystal is thought to have an undesirable impact on product selectivities by yielding more light gas and LPG; however, this has not been proven commercially. In the manufacturing of USY catalyst, the zeolite, clay, and binder are slurried together. If the binder is not active, an alumina component having catalytic properties may also be added. The well-mixed slurry solution is then fed to a spray dryer. The function of a spray dryer is o form microspheres by evaporating the slurry solution, through the use of atomizers, in the presence of hot air. The type of spray dryer and the drying conditions determine the size and distribution of catalyst particles.
Engelhard Process
Engelhard's "in-situ" FCC catalyst technology is mainly based on growing zeolite within the kaolin-based particles as shown in Figure 3-9A. The aqueous solution of various kaolins is spray dried to form microspheres. The microspheres are hardened in a high-temperature (1,300°F/704°C) calcination process. The NaY zeolite is produced by digestion of the microspheres, which contain metakaolin, and mullite with caustic or sodium silicate. Simultaneously, an active matrix is ormed with the microspheres. The crystallized microspheres are iltered and washed prior to ion exchange and any final treatment.
With each shipment of fresh catalyst, the catalyst suppliers typically mail refiners an inspection report that contains data on the catalyst's physical and chemical properties. This data is valuable and should be monitored closely to ensure that the catalyst received meets the agreed specifications. A number of refiners independently analyze random samples of the fresh catalyst to confirm the reported properties. In addition, quarterly review of the fresh catalyst properties with the catalyst vendor will ensure that the control targets are being achieved.
00
Fluid Catalytic Cracking Handbook
he particle size distribution (PSD), sodium (Na), rare earth (RE), and urface area (SA) are some of the parameters in the inspection sheet hat require close attention.
article Size Distribution (PSD)
The PSD is an indicator of the fluidization properties of the catalyst, n general, fluidization improves as the fraction of the 0-40 micron articles is increased; however, a higher percentage of 0-40 micron articles will also result in greater catalyst losses. The fluidization characteristics of an FCC catalyst largely depend on he unit's mechanical configuration. The percentage of less than 40 microns in the circulating inventory is a function of cyclone efficiency. n units with good catalyst circulation, it may be economical to minimize he fraction of less than 40-micron particles. This is because after a few ycles, most of the 0-40 microns will escape the unit via the cyclones, The catalyst manufacturers control PSD of the fresh catalyst, mainly hrough the spray-drying cycle. In the spray dryer, the catalyst slurry must be effectively atomized to achieve proper distribution. As illurated in Figure 3-10, the PSD does not have a normal distribution hape. The average particle size (APS) is not actually the average size f the catalyst particles, but rather the median value.
urface Area (SA), M2/g
The reported surface area is the combined surface area of zeolite nd matrix. In zeolite manufacturing, the measurement of the zeolite urface area is one of the procedures used by catalyst suppliers to ontrol quality. The surface area is commonly determined by the mount of nitrogen adsorbed by the catalyst. The surface area correlates fairly well with the fresh catalyst activity. Upon request, catalyst suppliers can also report the zeolite surface area. his data is useful in that it is proportional to the zeolite content of he catalyst.
odium (Na), wt%
Sodium plays an intrinsic part in the manufacturing of FCC catalysts. s effects are well known and, because it deactivates the
FCC Catalysts
Figure 3-10.
101
Particle size distribution of a typical FCC catalyst.
zeolite and reduces the gasoline octane, every effort should be made o minimize the amount of sodium in the fresh catalyst. The catalyst nspection sheet expresses sodium or soda (Na2O) as the weight percent on the catalyst. When comparing different grades of catalysts, t is more practical to express the sodium content on the zeolite.
Rare earth (RE) is a generic name for 14 metallic elements of the Ianthanide series. These elements have similar chemical properties and are usually supplied as a mixture of oxides extracted from ores such as bastnaesite or monazite. Rare earth improves the catalyst activity (Figure 3-11) and hydrohermal stability. Catalysts can have a wide range of rare earth levels.
02
Fluid Catalytic Cracking Handbook
0
1
2
3
Rare Earth, wt% Figure 3-11. Effect of rare earth on catalyst activity.
epending on the refiner's objectives. Similar to sodium, the inspection heet shows rare earth or rare earth oxide (RE2O3) as the weight ercent of the catalyst. Again, when comparing different catalysts, the oncentration of RE on the zeolite should be used.
EQUILIBRIUM CATALYST ANALYSIS
Refiners send E-cat samples to catalyst manufacturers on a regular asis. As a service to the refiners, the catalyst suppliers provide nalyses of the samples in a form similar to the one shown in Figure -12. Although the absolute E-cat results may differ from one vendor o another, the results are most useful as a trend indicator. The tests performed on E-cat samples provide refiners with valuable nformation on unit conditions. The data can be used to pinpoint otential operational, mechanical, and catalyst problems because the hysical and chemical properties of the E-cat provide clues on the nvironment to which it has been exposed. The following discussion describes each test briefly and examines the ignificance of these data to the refiner. The E-cat results are divided nto catalytic properties, physical properties, and chemical analysis.
C wt% 0.23 0.23 0.16 0.23 0.22 0.20 0.24 0.15 0.24 0 0 0 0 0 0 0 0 0
1.3 1.2 1.2
Fe ppm 5600 5600 5600 5600 5600 5600 5600 5600 5600
9
1.3 1.4 1.3 1.2 1.4 1.2
G.F.
2.2 1.9 3.1 2.6 3.2 2.6 2.3 2.8 2.9
C.F.
9 0 9 8 9 9 7 0
ata
0 0 0 2 0 0 2 0 4
0-40 wt% 10 7 8 9 6 9 10 7 10
r TM
0-80 wt% 63 61 67 69 65 67 71 64 67 Cu Sb UCS RE203 ppm ppm 24.27 1.79 25 416 23 1,80 446 24 440 ' 24.27 1.79 24 r™44i 24 1.79 24.25 445 23 [_ 420 1.80 179 24 24.27 458 25 432 1.79 24 24.27 409 1.76
0-20 wt%
Figure 3-12. Typical E-cat analysis.
V Ni ppm ppm 1997 4106 4093 1948 4051 1940 4099 1974 4017 1942 3962 1910 3892 1893 3893 L_J885 1873 3875
S.A. P.V. ABD m2/gm, cc/gm gm/cc 147 0.30 0.83 148 0.83 0.28 0.84 147 0.29 148 0.83 0.29 0.28 0.83 148 0,29 0.84 150 148 0.28 0.85 148 0.85 0.29 0.28 0.84 148
130 130 130 130 130 132 131 130 130]
Z
70 72 69 68 70 69 67 71 69
APS
M
1 1 1 1 1 1 1 1 1
AI2O3 ppm 28. 29. 29. 28. 28. 28. 28. 28. 28.
04
Fluid Catalytic Cracking Handbook
Catalytic Properties
The activity, coke, and gas factors are the tests that reflect the elative catalytic behavior of the catalyst.
Conversion (Activity)
The first step in E-cat testing is to burn the carbon off the sample. he sample is then placed in a MAT unit (Figure 3-13), the heart of which is a fixed bed reactor. A certain amount of a standard gas oil eedstock is injected into the hot bed of catalyst. The activity is eported as the conversion to 430°F (221°C) material. The feedstock's uality, reactor temperature, catalyst-to-oil ratio, and space velocity are our variables affecting MAT results. Each catalyst vendor uses slightly ifferent operating variables to conduct microactivity testing, as ndicated in Table 3-2. In commercial operations, catalyst activity is affected by operating onditions, feedstock quality, and catalyst characteristics. The MAT eparates catalyst effects from feed and process changes. Feed conaminants, such as vanadium and sodium, reduce catalyst activity. -cat activity is also affected by fresh catalyst makeup rate and egenerator conditions.
Coke Factor (CF), Gas Factor (GF)
The CF and GF represent the coke- and gas-forming tendencies of an -cat compared to a standard steam-aged catalyst sample at the same onversion. The CF and GF are influenced by the type of fresh catalyst nd the level of metals deposited on the E-cat. Both the coke and gas actors can be indicative of the dehydrogenation activity of the metals on he catalyst. The addition of amorphous alumina to the catalyst will tend o increase the nonselective cracking, which forms coke and gas.
hysical Properties
The tests that reflect physical properties of the catalyst are surface rea, average bulk density, pore volume, and particle size distribution.
urface Area (SA), M2/g
For an identical fresh catalyst, the surface area of an E-cat is an ndirect measurement of its activity. The SA is the sum of zeolite and
FCC Catalysts Star dard FCCl Feed n-Q Equilibrium Catalysts
!
L
105
Syringe Pump
J -"1^3 way Valve
@
Coke Burn Off -
"I I Reactor Furnace
control
Temp
Tempcontrol
ri i— is l~W
Preheat Zone Catalyst y
rJFlow Meter L^«—<•) Purge N,
Gas Product Sample
1 \
lllllllnl
„„
n
\
.„ „
1
1 i
/ ^^^
/ / / pent Catalyst
/
-''tL \
[""F7',
/'
T
Cold Bath j salt Solution / i Gas Collector J-i ' / Gas Volume . / Determination ,
r
Liquid Product To Leco Analyzer for To Gas Chromatograph Coke Determination for Analysis of \ Light Hydrocarbons and Simulated Distillation
/
Gas Product f To Gas Chromatograth for Component Analysis
N
Material Balance Detailed Produce Yields Activity Gas Factor, Coke Factor
Figure 3-13.
Typical MAT equipment [3].
matrix surface areas. Hydrothermal conditions in the cat cracker estroy the zeolite cage structure, thus reducing its surface area. They lso dealuminate the zeolite framework. Hydrothermal treatment has ess effect on the matrix surface area, but the matrix surface area is ffected by the collapse of small pores to become larger pores.
06
Fluid Catalytic Cracking Handbook Table 3-2 Equilibrium Microactivity Test Conditions
Tester (United States)
AKZO Davisoe ngelhard
Temp. °F/C°
Cat-toOil Wt. Ratio
998/537 980/527 910/488
3.0 4.0 5.0
WHSV hr1
Catalyst Contact Time, Seconds
Feed Source
Reactor Type
NA 30 15
75 30 48
(1) (2) (3)
Isothermal Isothermal Isothermal
) Kuwait vacuum gas oil ) Sour import heavy gas oil ) Mid-continent Alzo uses a fluid bed testing unit
MAT Gas Oil Properties
Properties
Akzo*
Davison*
Engelhard**
API Gravity D-1160 IBP, °F 50%, °F 90%, °F onearbon, wt% ulfur, wt% otal Nitrogen, ppmv API Procedure 2B4.1 Aromatics, vol% Naphthenes, vol% Paraffins
20.4
22.5
28.6
674 883 934 0.17 3.18 1,009
423 755 932 0.25 2.59 860
373 732 899 0.22 0.52 675
21.9 25.4 52.6
21.7 19.6 58.7
30 28 42
*Azko Private Communication, July 1997 f Grace Davison Catalagram, No. 79, 1989 *Engelhard Catalyst Report, No. Tl-825
pparent Bulk Density (ABD), g/cc
Bulk density can be used to troubleshoot catalyst flow problems. A oo-high ABD can restrict fluidization, and a too-low ABD can result n excessive catalyst loss. Normally, the ABD of the equilibrium atalyst is higher than the fresh catalyst ABD due to thermal and ydrothermal changes in pore structure that occur in the unit.
FCC Catalysts
101
Pore Volume (PV), cc/g
Pore volume is an indication of the quantity of voids in the catalyst articles and can be a clue in detecting the type of catalyst deactivation hat takes place in a commercial unit. Hydrothermal deactivation has ery little effect on pore volume, whereas thermal deactivation decreases ore volume.
Pore Diameter (°A)
The average pore diameter (APD) of a catalyst can be calculated rom the E-cat analysis sheet by using the following equation: PV x 4 x 10,000 APD(°A) =
SA Example 3-1
For an E-cat with a PV = 0.40 cc/g and SA = 120 m2/g, determine APD, APD = 133 °A
Particle Size Distribution (PSD)
PSD is an important indicator of the fluidization characteristics f the catalyst, cyclone performance, and the attrition resistance of he catalyst. A drop in fines content indicates the loss of cyclone fficiency. This can be confirmed by the particle size of fines collected ownstream of the cyclones. An increase in fines content of the E-cat ndicates increased catalyst attrition. This can be due to changes in resh catalyst binder quality, steam leaks, and/or internal mechanical roblems, such as those involving the air distributor or slide valves.
Chemical Properties
The key elements that characterize chemical composition of the catalyst re alumina, sodium, metals, and carbon on the regenerated catalyst.
lumina (A12O3)
The alumina content of the E-cat is the total weight percent of lumina (active and inactive) in the bulk catalyst. The alumina content
08
Fluid Catalytic Cracking Handbook
f the E-cat is directly related to the alumina content of the fresh atalyst. When changing catalyst grades, the alumina level of the Eat is often used to determine the percent of new catalyst in the unit.
odium (Na)
The sodium in the E-cat is the sum of sodium added with the feed nd sodium on the fresh catalyst. A number of catalyst suppliers report odium as soda (Na2O). Sodium deactivates the catalyst acid sites and auses collapse of the zeolite crystal structure. Sodium can also reduce he gasoline octane, as discussed earlier,
ickel (Ni), Vanadium (V), Iron (Fe), Copper (Cu)
These metals, when deposited on the E-cat catalyst, increase coke nd gas-making tendencies of the catalyst. They cause dehydrogenation eactions, which increase hydrogen production and decrease gasoline ields. Vanadium can also destroy the zeolite activity and thus lead o lower conversion. The deleterious effects of these metals also epend on the regenerator temperature: the rate of deactivation of a metal-laden catalyst increases as the regenerator temperature increases. These contaminates originate largely from the heavy (1,050+ °F/ 66+ °C), high-molecular weight fraction of the FCC feed. The uantity of these metals on the E-cat is determined by their levels in he feedstock and the catalyst addition rate. Essentially, all these metals n the feed are deposited on the catalyst. Most of the iron on the Eat comes from metal scale from piping and from the fresh catalyst. Metals content of the E-cat can be determined fairly accurately by onducting a metals balance around the unit: Metalsin - Metalsout = Metals Accumulated
his is a first order differential equation. Its solution is: Me = A + [M0 - A] x e-Ca
Where:
Me = = W = Mf =
E-cat Metals Content, ppm (W x Mf)/Ca Feed rate, Ib/day Feed Metals, ppm
X
^
FCC Catalysts
109
Ca = Catalyst Addition Rate, Ib/day M0 = Initial Metals on the E-cat, ppm = Time, day = Catalyst Inventory, Ib
At steady state, the concentration of any metal on catalyst is: M = A =
(W x M f )
141 5
'
M
=
-x 350.4 x Mf B
= Catalyst addition rate, pounds of catalyst per barrel of feed
igure 3-14 is the graphical solution to the above equation and can e employed to estimate metals content of the E-cat, based on feed metals and catalyst addition rate.
Carbon (C)
The deposition of carbon on the E-cat during cracking will temporrily block some of the catalytic sites. The carbon, or more accurately he coke, on the regenerated catalyst (CRC) will lower the catalyst ctivity and, therefore, the conversion of feed to valuable products Figure 3-15). The CRC is an important parameter for a unit operator to monitor eriodically. Most FCC units check for CRC on their own, usually aily. The CRC is an indicator of regenerator performance. If the CRC hows signs of increasing, this could reveal malfunction of the regenrator's air/spent catalyst distributors. It should be noted that the MAT umbers reported on the E-cat sheet are determined after the CRC has een completely burned off.
CATALYST MANAGEMENT
Depending on the design of a cat cracker, the circulating inventory an contain 30-1,200 tons of catalyst. Fresh catalyst is added to the nit continually to replace the catalyst lost by attrition and to maintain
Fluid Catalytic Cracking Handbook
10
9000
8000
2000
0
0.05
0,1
0.15
0.2
0.2S
0.3
0.35
0.4
0.45
0.5
O.S5
0.6
0.66
07
Catalyst Additions, Ib/bbl
igure 3-14. Catalyst metals content versus catalyst addition rate for 22°API Gravity Feed. (Source: Katalystics' Regional Technology Seminar, New Orleans, ouisiana, December 15, 1998.)
atalyst activity. The daily makeup rate is typically 1% to 2% of nventory or 0.1 to 0.3 pounds (0.045 to 0.14 kg) of catalyst per barrel f fresh feed. In cases where the makeup rate for activity maintenance xceeds catalyst losses, part of the inventory is periodically withdrawn
FCC Catalysts
111
100
90 80 70 60 50 40
Figure 3-15.
0.0
0.5
1.0
1.5
CRC(wt%)
Catalyst activity retention vs. carbon on regenerated catalyst [12].
rom the unit to control the catalyst level in the regenerator. Catalyst ines leave the unit with the regenerator flue gas and the reactor vapor. The catalyst ages in the unit, losing its activity and selectivity. The eactivation in a given unit is largely a function of the unit's mechanical onfiguration, its operating condition, the type of fresh catalyst used, nd the feed quality. The primary criterion for adding fresh catalyst s to arrive at an optimum E-cat activity level. A too-high E-cat activity will increase delta coke on the catalyst, resulting in a higher regenrator temperature. The higher regenerator temperature reduces the atalyst circulation rate, which tends to offset the activity increase. The amount of fresh catalyst added is usually a balance between atalyst cost and desired activity. Most refiners monitor the MAT data rom the catalyst vendor's equilibrium data sheet to adjust the fresh atalyst addition rate. It should be noted that MAT numbers are based n a fixed-bed reactor system and, therefore, do not truly reflect the ynamics of an FCC unit. A catalyst with a high MAT number may r may not produce the desired yields. An alternate method of measurng catalyst performance is dynamic activity. Dynamic activity is alculated as shown below:
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Fluid Catalytic Cracking Handbook
. . . . (Second Order Conversion) rv Dynamic Activity = (Coke Yield, Wt% of Feed)
Where; Second Order Conversion =
(MAT Conversion, Vol%) (100 - MAT Conversion, Vol%)
or example, a catalyst with a MAT number of 70 vol% and a 3.0 wt% coke yield will have a dynamic activity of 0.78. However, another atalyst with a MAT conversion of 68 vol% and 2.5 wt% coke yield will have a dynamic activity of 0.85. This could indicate that in a ommercial unit the 68 MAT catalyst could outperform the 70 MAT atalyst, due to its higher dynamic activity. Some catalyst vendors have egun reporting dynamic activity data as part of their E-cat inspecon reports. The reported dynamic activity data can vary significantly om one test to another, mainly due to the differences in feedstock uality between MAT and actual commercial application. In addition, he coke yield, as calculated by the MAT procedure, is not very ccurate and small changes in this calculation can affect the dynamic ctivity appreciably. The most widely accepted model to predict E-cat activity is based n a first-order decay type [7]: A A
-
(t)
AA
(0)
X xe
c~ (S+K)t ,
A
0 X S v n _, e-
At steady state, the above equation reduces to: A(t) ~~ - AA(0) xx ce -(S+K)t I A°c
A
X S
, v O + IS.
x f'1^ - p-tK+S)') '
Where:
(t)
0
= = = =
Catalyst microactivity at anytime Catalyst microactivity at starting time Time after changing catalyst or makeup rate Daily fractional replacement rate = addition rate/inventory Deactivation constant = ln(At - A0)/-t
FCC Catalysts
113
xample 3-2 below illustrates the use of the above equations. Example 3-2
Assume a 50,000 bpd cat cracker with: * » * *
Catalyst inventory of 300 tons Makeup rate of 4.0 tons per day Fresh catalyst MAT conversion 80 vol% E-cat MAT number 71.5 vol%
Determine: a. New E-cat MAT conversion if the addition rate is reduced to 3.0 tons/day: si ?5 —
40 — U.ulj33 days 300
t = 300/4 = 75 days Deactivation Constant = K = ^——
-75.0
^ = 0.001498 days'!
30 300
New Fractional Replacement = —— = 0.01 days"1
The revised E-cat =
'• = 69.5 vol% 0.01333 + 0.001498
b. The new E-cat MAT conversion if the fresh catalyst MAT number is reduced from 80 vol% to 75 vol%: A =
75x0.01333 ^ ,„ = 67 vol% 0.01333 + 0.001498
When a refiner changes the FCC catalyst, it is often necessary to etermine the percent of the new catalyst in the unit. The following quation, which is based on a probability function, can be used to esmate the percent changeover. p = 1 - e~fxsxt
14
Fluid Catalytic Cracking Handbook
Where: = Fractional Changeover = Retention Factor, usually in the range of 0.6 to 0.9 = Replacement Rate = Addition rate/Inventory = Time, day Example 3-3
The 300- ton inventory unit in Example 3-2 is changing catalyst type nd planning to add 3.5 tons per day of new catalyst. Determine the ercent of changeover after 60 days of operation. Assume a retention actor of 0.7. -0.7x0.0117x60
P = 0,387 or 38.7%
Another method of calculating the % changeover is by the use of lumina balance. Example 3-4
The example below illustrates the application of the alumina balance method. For the same 300-ton inventory unit, assume the alumina (A12O3) ontents of the present and new fresh catalysts are 48 wt% and 38 wt%, respectively. Sixty (60) days after the catalyst switch, the alumina ontent of E-cat is 43 wt%. Determine % changeover: Al2O3 (new) - Al2O3 (equil.) Functional Changeover = 1 - —-—--—A12O3 (new) - Al 2 O 3 (old) Fractional Changeover = 1 - 38 — 48
-
= 0.5 = —>}50%
This method can also be used to calculate the catalyst retention actor. The above equations assume steady-state operation, constant nit inventory, and constant addition and loss rate.
FCC Catalysts
115
Catalyst management is a very important aspect of the FCC process. election and management of the catalyst, as well as how the unit is perated, are largely responsible for achieving the desired product. roper choice of a catalyst will go a long way toward achieving a uccessful cat cracker operation. Catalyst change-out is a relatively imple process and allows a refiner to select the catalyst that maximizes he profit margin. Although catalyst change-out is physically simple, requires a lot of homework as discussed later in this section. As many catalyst formulations are available, catalyst evaluation hould be an ongoing process. However, it is not an easy task to valuate the performance of an FCC catalyst in a commercial unit ecause of continual changes in feedstocks and operating conditions, n addition to inaccuracies in measurements. Because of these limitaons, refiners sometimes switch catalyst without identifying he objectives and limitations of their cat crackers. To ensure that a roper catalyst is selected, each refiner should establish a methodology hat allows identification of "real" objectives and constraints and nsures that the choice of the catalyst is based on well-thoughtut technical and business merits. In today's market, there are over 20 different formulations of FCC catalysts. Refiners should evalute catalyst mainly to maximize profit opportunity and to minimize sk. The "right" catalyst for one refiner may not necessarily be "right" or another, A comprehensive catalyst selection methodology will have the ollowing elements: 1. Optimize unit operation with current catalyst and vendor a. Conduct test run b. Incorporate the test run results into an FCC kinetic model c. Identify opportunities for operational improvements d. Identify unit's constraints e. Optimize incumbent catalyst with vendor 2. Issue technical inquiry to catalyst vendors a. Provide test run results b. Provide E-cat sample c. Provide processing objectives d. Provide unit limitations
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Fluid Catalytic Cracking Handbook
3, Obtain vendor responses a. Obtain catalyst recommendation b. Obtain alternate recommendation c. Obtain comparative yield projection 4, Obtain current product price projections a. For present and future four-quarters 5, Perform economic evaluations on vendor yields a. Select catalysts for MAT evaluations 6, Conduct MAT of selected list a. Perform physical and chemical analyses b. Determine steam deactivation conditions c. Deactivate incumbent fresh catalyst to match incumbent E-cat d. Use same deactivation steps for each candidate catalyst 7, Perform economic analysis of alternatives a. Estimate commercial yield from MAT evaluations 8, Request commercial proposals a. Consult at least two vendors b. Obtain references c. Check references 9, Test the selected catalysts in a pilot plant a. Calibrate the pilot plant steaming conditions using incumbent E-cat b. Deactivate the incumbent and other candidate catalysts c. Collect at least two or three data points on each by varying the catalyst-to-oil ratio 10. Evaluate pilot plant results a. Translate the pilot plant data b. Use the kinetic model to heat-balance the data c. Identify limitations and constraints 11. Make the catalyst selection a. Perform economic evaluation b. Consider intangibles-research, quality control, price, steady supply, manufacturing location c. Make recommendations 12. Post selection a. Monitoring transition-% changeover b. Post transition test run c. Confirm computer model
FCC Catalysts
117
13. Issue the final report a. Analyze benefits b. Evaluate selection methodology
There is a redundancy of flexibility in the design of FCC catalysts. Variation in the amount and type of zeolite, as well as the type of ctive matrix, provide a great deal of catalyst options that the refiner an employ to fit its needs. For smaller refiners, it may not be practical o employ pilot plant facilities to evaluate different catalysts. In this ase, the above methodology can still be used with emphasis shifted oward using the MAT data to compare the candidate catalysts. It is mportant that MAT data are properly corrected for temperature, soaking time," and catalyst strippability effects.
For many years, cat cracker operators have used additive compounds or enhancing cat cracker performance. The main benefits of these dditives (catalyst and feed additives) are to alter the FCC yields and educe the amount of pollutants emitted from the regenerator. The dditives discussed in this section are CO promoter, SOX reduction, SM-5, and antimony.
CO Promoter
The CO promoter is added to most FCC units to assist in the ombustion of CO to CO2 in the regenerator. The promoter is added o accelerate the CO combustion in the dense phase and to minimize he higher temperature excursions that occur as a result of afterburning n the dilute phase. The promoter allows uniform burning of coke, articularly if there is uneven distribution between spent catalyst and ombustion air. Regenerators operating in full or partial combustion can utilize the enefits of the CO promoter. The addition of the promoter tends to ncrease the regenerator temperature and NOx emission. The metallurgy f the regenerator internals should be checked for tolerance of the igher temperature. The active ingredients of the promoter are typically the platinum roup metals. The platinum, in the concentration of 300 ppm to 800
18
Fluid Catalytic Cracking Handbook
pm, is typically dispersed on a support. The effectiveness of the romoter largely depends on its activity and stability. Promoter is frequently added to the regenerator two to three times day, normally at a rate of 3 to 5 pounds (1 to 2.3 kg) promoter per on of fresh catalyst. The concentration of platinum required in a unit nventory is about 0.5 to 1.5 ppm. The promoter addition rate may e increased if antimony solution is being used to passivate the nickel. The use of CO promoter, particularly during unit start-ups, improves he stability of the regeneration operation. However, not every cat racker can justify combustion-promoted operation. Heat balance, vailability of combustion air, NOX emission metallurgical limits, and he presence of CO boiler are some of the factors that should be onsidered before using combustion promoter. For example, in units perating with low oxygen levels and partial combustion, a promoted ystem could increase carbon on regenerated catalyst (CRC). This is ecause CO combustion reaction competes with carbon burning reacion for the available oxygen. The combustion of CO to CO2 will also ncrease NOX emissions. This is largely due to the oxidation of intermediates such as ammonia and cyanide gases into nitrogen oxide (NO).
SOX Additive
The coke on the spent catalyst entering the regenerator contains ulfur. In the regenerator, the sulfur in the coke is converted to SO2 nd SO3. The mixture of SO2 and SO3 is commonly referred to as SOX, nd approximately 80% to 90% of SOX is SO2, with the rest being SO3, The SOX leaves the regenerator with the flue gas and is eventually ischarged to the atmosphere. Coke yield, thiophenic sulfur content f the feed, the regenerator operating condition, and the type of FCC atalyst are the major factors affecting SOX emissions. The environmental impact of SOX emissions has gained much ttention over the past ten years. The United States Environmental rotection Agency (EPA) New Source Performance Standards (NSPS) went into effect in 1989. The ruling covers new, modified, and recontructed FCC units since January 1994. It should be noted that the outhern California Air Quality Management District (SCAQMD) oard has established a limit of 60 kilograms of SOX per 1,000 barrels f feed for the existing FCC units.
FCC Catalysts
119
There are three common methods for SOX abatement. These are flue as scrubbing, feedstock desulfurization, and SOX additive. The SOX dditive is often the least costly alternative, which is the approach racticed by many refiners. The SOX additive, usually a metal oxide, is added directly to the atalyst inventory. The additive works by adsorbing and chemically onding with SO3 in the regenerator. This stable sulfate species is arried with the circulating catalyst to the riser, where it is reduced r "regenerated" by hydrogen or water to yield H2S and metal oxide. able 3-3 shows the postulated chemistry of SOX reduction by a OX agent. To achieve the highest efficiency of SOX additive, it is imporant that: * Excess oxygen be available; oxygen promotes the SO2 to SO3 reaction. SOX additive will only form a metal sulfate from SOV * The regenerator temperature be lower; lower temperature favors SO2 + 1/2 O2 -> SO3 * The capturing agent be physically compatible with the FCC catalyst and easily regenerated in the riser and stripper. * CO promoter be used, which oxidizes SO2 to SO3. * There be a uniform distribution of air and spent catalyst. Air/ catalyst mixing in the regenerator can significantly affect the SOX pick-up efficiency.
Table 3-3 Mechanism of Catalytic SOX Reduction
. In the Regenerator Sulfur in Coke (S) + O2 SO2 + l/2 O2 MXO + SO,
—» —> —>
MXSO4
—> -> —>
MXS + 4H2O MXO + H2S + 3 H2O MXO + H2O
SO2 + SO3 SO3
. In the Reactor and Stripper MXSO4 + 4H2
Mxso4 + 4H2 MXS + H2O
ource: Thiel [9]
20
Fluid Catalytic Cracking Handbook
* Operation of the reactor stripper be efficient. The stripper efficiency is very important to allow the release of sulfate and the formation of H2S.
Since most of the regenerators operating in full combustion mode sually operate with 1% to 3% excess oxygen, the capturing efficiency f SOX additive is often greater in full combustion than in partial ombustion units.
ZSM-5
ZSM-5 is Mobil Oil's proprietary shape-selective zeolite that has a ifferent pore structure from that of Y-zeolite. The pore size of ZSMis smaller than that of Y-zeolite (5.1°A to 5.6°A versus 8°A to 9°A), n addition, the pore arrangement of ZSM-5 is different from Y-zeolite, s shown in Figure 3-16. The shape selectivity of ZSM-5 allows
Figure 3-16.
Comparison of Y faujasite and ZSM-5 zeolites [13].
FCC Catalysts
121
referential cracking of long-chain, low-octane normal paraffins, as well as some olefins, in the gasoline fraction. ZSM-5 additive is added to the unit to boost gasoline octane and o increase light olefin yields. ZSM-5 accomplishes this by upgrading ow-octane components in the gasoline boiling range (C7 to C1O) into ght olefins (C3, C4, C5), as well as isomerizing low-octane linear lefins to high-octane branched olefins, ZSM-5 inhibits paraffin ydrogenation by cracking the C7+ olefins. ZSM-5's effectiveness depends on several variables. The cat crackers hat process highly paraffinic feedstock and have lower base octane ill receive the greatest benefits of using ZSM-5. ZSM-5 will have ttle effect on improving gasoline octane in units that process naphthenic eedstock or operate at a high conversion level. When using ZSM-5, there is almost an even trade-off between FCC asoline volume and LPG yield. For a one-number increase in the esearch octane of FCC gasoline, there is a 1 vol% to 1.5 vol% ecrease in the gasoline and almost a corresponding increase in the PG, This again depends on feed quality, operating parameters, and ase octane. The decision to add ZSM-5 depends on the objectives and conraints of the unit. ZSM-5 application will increase load on the wet as compressor, FCC gas plant, and other downstream units. Most efiners who add ZSM-5 do it on a seasonal basis, again depending n their octane need and unit limitations. The concentration of the ZSM-5 additive should be greater than 1 % f the catalyst inventory to see a noticeable increase in the octane. An octane boost of one research octane number (RON) will typically equire a 2% to 5% ZSM-5 additive in the inventory. It should be oted that the proper way of quoting percentage should be by SM-5 concentration rather than the total additive because the activity nd attrition rate can vary from one supplier to another. There are new enerations of ZSM-5 additives that have nearly twice the activity of he earlier additives. In summary, ZSM-5 provides the refiner the flexibility to increase asoline octane and light olefins. With the introduction of reformulated asoline, ZSM-5 could play an important role in producing isoutylene, used as the feedstock for production of methyl tertiary butyl ther (MTBE).
22
Fluid Catalytic Cracking Handbook
Metal Passivation
As discussed in Chapter 2, nickel, vanadium, and sodium are the metal compounds usually present in the FCC feedstock. These metals eposit on the catalyst, thus poisoning the catalyst active sites. Some f the options available to refiners for reducing the effect of metals n catalyst activity are as follows: • • • • • •
Increasing the fresh catalyst makeup rate Using outside E-cat Employing metal passivators Incorporating metal trap into the FCC catalyst Using demetalizing technology to remove the metals from the catalyst The MagnaCat separation process (demetalizing technology), which allows discarding the "older" catalyst particles containing higher metal levels
Metal passivation in general, and antimony in particular, are discussed n the following section. In recent years, several methods have been patented for chemical assivation of nickel and vanadium. Only some of the tin compounds ave had limited commercial success in passivating vanadium. Although n has been used by some refiners, it has not been proven or as widely ccepted as antimony. In the case of nickel, antimony-based comounds have been most effective in reducing the detrimental effects f nickel poisoning. It should be noted that, although the existing ntimony-based technology is the most effective method of reducing he deleterious effects of nickel, the antimony is fugitive and can be onsidered hazardous. In this case, a bismuth-based passivator may be better choice.
ntimony
Antimony-based passivation was introduced by Phillips Petroleum n 1976 to passivate nickel compounds in the FCC feed. Antimony is njected into the fresh feed, usually with the help of a carrier such as ght cycle oil. If there are feed preheaters in the unit, antimony should e injected downstream of the preheater to avoid thermal decomposion of the antimony solution in the heater tubes. The effects of antimony passivation are usually immediate. By orming an alloy with nickel, the dehydrogenation reactions that are
FCC Catalysts
123
aused by nickel are often reduced by 40% to 60%. This is evidenced y a sharp decline in dry gas and hydrogen yield. Nickel passivation is generally economically attractive when the ickel content of the E-cat is greater than 1,000 ppm. The Phillips etroleum secondary antimony patent position is due to expire in late 1999, At that time, antimony passivation can become economically ttractive at a lower nickel level than 1,000 ppm. The antimony solution should be added in proportion to the amount f nickel present in the feed. The optimum dosage normally coresponds to an antimony-to-nickel ratio of 0.3 to 0.5 on the E-cat. Antimony's retention efficiency on the catalyst is in the range of 75% o 85% without the recycling of slurry oil to the riser. If slurry recycle s being practiced, the retention efficiency is usually greater than 90%. Any antimony not deposited on the circulating catalyst ends up in the ecanted oil and the catalyst fines from the regenerator. It is often a ood practice to discontinue antimony injection about one month prior o a scheduled unit shutdown to make sure the exposure to catalyst ust containing antimony is reduced to a minimum when wearing a alf-faced respirator.
SUMMARY
The introduction of zeolite into the FCC catalyst in the early 1960s was one of the most significant developments in the field of cat racking. The zeolite greatly improved selectivity of the catalyst, esulting in higher gasoline yields and indirectly allowing refiners to rocess more feed to the unit. With the introduction of reformulated asoline, new formulations in FCC catalyst will again help refiners meet new requirements in gasoline quality. Since there are over 120 different FCC catalyst formulations in the market today, it is important that the refinery personnel involved in at cracker operations have some fundamental understanding of catalyst echnology. This knowledge is useful in areas such as proper troublehooting and customizing a catalyst that would match the refiner's eeds. The additive technology will be expanding in coming years. he need to produce reformulated gasoline will increase demand for he shape-selective zeolite, such as ZSM-5. The pressure from environmental agencies to reduce SOX and NOX will further increase the emand for additives that reduce emissions.
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Fluid Catalytic Cracking Handbook
REFERENCES 1. Breck, D. W., Zeolite Molecular Sieves: Structure, Chemistry, and Use, New York: Wiley Interscience, 1974. 2. Hayward, C. M. and Winkler, W. S., "FCC: Matrix/Zeolite," Hydrocarbon Processing, February 1990, pp. 55-56. 3. Upson, L. L., "What FCC Catalyst Tests Show," Hydrocarbon Processing November 1981, pp. 253-258. 4. Pine, L. A., Maher, P. J., and Wachter, W. A., "Prediction of Cracking Catalyst Behavior by a Zeolite Unit Cell Size Model," Journal of Catalysis, No. 85, 1984, pp. 466-476. 5. Magnusson, J. and Pudas, R., "Activity and Product Distribution Characteristics of the Currently Used FCC Catalyst Systems," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 2223, 1985. 6. John G. S. and Mikovsky, R. J., "Calculation of the Average Activity of Cracking Catalysts," Chemical Engineering Science, Vol. 15, 1961, pp. 172-175. 7. Gaughan, J. R., "Effect of Catalyst Retention on Inventory Replacement," Oil & Gas Journal, December 26, 1983, pp. 141-145. 8. Tamborski, G. A., Magnabosco, L. M., Powell, J. W., and Yoo, J. S., "Catalyst Technology Improvements Make SOX Emissions Control Affordable," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 22-23, 1985. 9. Thiel, P. G., Blazek, J. J., "Additive R," Grace Davison Catalagram, No. 71, 1985. 10. Engelhard Corporation, "Reduced Unit Cell Size Catalysts Offer Improved Octane for FCC Gasoline," The Catalyst Report, TI-762. 11. Engelhard Corporation, "Increasing Motor Octane by Catalytic Means Part 2," The Catalyst Report, EC6100P. 12. Engelhard Corporation, "The Chemistry of FCC Coke Formation," The Catalyst Report, Vol. 7, Issue 2. 13. Majon, R. J. and Spielman, J., "Increasing Gasoline Octane and Light Olefin Yields with ZSM-5," The Catalyst Report, Vol. 5, Issue 5, 1990. 14. Davison Div., W.R. Grace & Co., Grace Davison Catalagram, No. 72, 1985. 15. Humphries, Adrian P., "Zeolite Fundamentals and Synthesis," Akzo Chemicals, 1987. 16. Davison Octane Handbook. 17. G. Yaluris and A. W. Peters, "Studying the Chemistry of the FCCU Regenerator in the Laboratory Under Realistic Conditions," Grace Davison, Columbia, MD, 1998.
CHAPTER 4
Chemistry of FCC Reactions
A complex series of reactions (Table 4-1) take place when a large as-oil molecule comes in contact with a 1,200°F to 1,400°F (650°C o 760°C) FCC catalyst. The distribution of products depends on many actors, including the nature and strength of the catalyst acid sites. lthough most cracking in the FCC is catalytic, thermal cracking eactions also occur. Thermal cracking is caused by factors such as on-ideal mixing in the riser and poor separation of cracked products the reactor. The purpose of this chapter is to: • Provide a general discussion of the chemistry of cracking (both thermal and catalytic). • Highlight the role of the catalyst, and in particular, the influence of zeolites. • Explain how cracking reactions affect the unit's heat balance.
Whether thermal or catalytic, cracking of a hydrocarbon means the reaking of a carbon to carbon bond. But catalytic and thermal crackg proceed via different routes. A clear understanding of the different echanisms involved is beneficial in areas such as: • Selecting the "right" catalyst for a given operation • Troubleshooting unit operation • Developing a new catalyst formulation Topics discussed in this chapter are: • Thermal cracking • Catalytic cracking • Thermodynamic aspects
125
26
Fluid Catalytic Cracking Handbook Table 4-1 Important Reactions Occurring in FCC
. Cracking: Paraffins cracked to olefins and smaller paraffins Olefins cracked to smaller olefins
C9Hl8 -> C4H8 + C5H10
Aromatic side-chain scission
ArC10H21
Naphthenes (cyclo-paraffins) cracked to olefins and smaller ring compounds
Cyclo-C1oH20 -> C6H12
CSH1
ArC5H9
C4H8
. Isomerization: Olefin bond shift
1-C4H8 -^ trans-2-C4H8
Normal olefins to iso-olefin
n-C5H10 —> iso-C5H10
Normal paraffins to iso-paraffin
n-C4H10
Cyclo-hexane to cyclo-pentane
C6H12 + C5H9CH3
. Hydrogen Transfer:
iso-C4Hlo
Naphthene + Olefin -» Aromatic + Paraffin
Cyclo-aromatization
. Trans-alkylation/Alkyl-group Transfer
2C6H5CH3
C6H6
. Cyclization of Olefins to Naphthenes
C7H14 -^ CH3-cyclo-C6H ii
. Dehydrogenation
n-C8H18
. Dealkylation
H6 lso-C3H 7-C6H5 -» C6H6 -\. C C3 H
. Condensation
Ar-C3H == CH2 + R,CH = CHR2 v AT Ar + 2H
~^
C
8H16
+ H
2 3
7
6
jT\I
THERMAL CRACKING
Before the advent of the catalytic cracking process, thermal cracking was the primary process available to convert low-value feedstocks into ghter products. Refiners still use thermal processes, such as delayed oking and visibreaking, for cracking of residual hydrocarbons.
Chemistry of FCC Reactions
127
Thermal cracking is a function of temperature and time. The reaction ccurs when hydrocarbons in the absence of a catalyst are exposed to high emperatures in the range of 800°F to 1,200°F (425°C to 650°C). The initial step in the chemistry of thermal cracking is the formation f free radicals. They are formed upon splitting the C-C bond. A free adical is an uncharged molecule with an unpaired electron. The upturing produces two uncharged species that share a pair of elecrons. Equation 4-1 shows formation of a free radical when a paraffin molecule is thermally cracked.
R2 V "D IV,,
H
f*
7 IV
| H
I
*/"*
V-
T
t>
V- '
.IV
' H
(4-0
H
Free radicals are extremely reactive and short-lived. They can ndergo alpha scission, beta scission, and polymerization. (Alphacission is a break one carbon away from the free radical; betacission, two carbons away.) Beta-scission produces an olefin (ethylene) and a primary free adical (Equation 4-2), which has two fewer carbon atoms [1]: J\
""""' V.'..il.^ —~ VvlT'} "•"*— V-, —
Lisy ""' "
°~~~
Vx
S~M.-y T IT'iV--- ~™ V_-ilo
\iT"1"jiii. )
The newly formed primary free radical can further undergo betacission to yield more ethylene. Alpha-scission is not favored thermodynamicaily but does occur. Alpha-scission produces a methyl radical, which can extract a ydrogen atom from a neutral hydrocarbon molecule. The hydrogen xtraction produces methane and a secondary or tertiary free radical Equation 4-3).
-» CH4 + R-CH2-CH2-CH2-CH2-'CH-CH2-CH3
(4-3)
This radical can undergo beta-scission. The products will be an lpha-olefin and a primary free radical (Equation 4-4).
28
Fluid Catalytic Cracking Handbook
R-CH2-CH2-CH2-CH2-'CH-CH2-CH3 -» R-CH2-CH2-'CH2 + H2C=CH-CH2-CH3
(4-4)
Similar to the methyl radical, the R-*CH2 radical can also extract a hydrogen atom from another paraffin to form a secondary free radical nd a smaller paraffin (Equation 4-5). R,-'CH 2 + R-CH2-CH2-CH2-CH2-CH2-CH2-CH3 -> R,-CH 3 + R-CH2-CH2-CH2-CH2-CH2-*CH-CH3
R-*CH? is more stable than H3*C. Consequently, the hydrogen extracion rate of R-*CH2 is lower than that of the methyl radical. This sequence of reactions forms a product rich in C} and G,, nd a fair amount of alpha-olefins. Free radicals undergo little branchng (isomerization). One of the drawbacks of thermal cracking in an FCC is that a high percentage of the olefins formed during intermediate reactions polymerize and condense directly to coke. The product distribution from thermal cracking is different rom catalytic cracking, as shown in Table 4-2. The shift in product istribution confirms the fact that these two processes proceed via ifferent mechanisms,
CATALYTIC CRACKING Catalytic reactions can be classified into two broad categories: * Primary cracking of the gas oil molecules • Secondary rearrangement and re-cracking of cracked products
Before discussing mechanisms of the reactions, it is appropriate to eview FCC catalyst development and examine its cracking properties. An in-depth discussion of FCC catalyst was presented in Chapter 3.
FCC Catalyst Development
The first commercial fluidized cracking catalyst was acid-treated atural clay. Later, synthetic silica-alumina materials containing 10 to
Chemistry of FCC Reactions
129
Table 4-2 Comparison of Products of Thermal and Catalytic Cracking
Hydrocarbon Type
-Paraffms
Thermal Cracking
Catalytic Cracking
C2 is major product, with C3 to C6 is major product; much C3 and C3, and C4 to few n-olefins above C4; C16 olefins; little branching much branching
Olefins
Slow double-bond shifts and little skeletal isomerization; H-transfer is minor and nonselective for tertiary olefins; only small amounts of aromatics formed from aliphatics at 932°F (500°C)
Rapid double-bond shifts, extensive skeletal isomerization, H-transfer is major and selective for tertiary olefins; large amounts of aromatics formed from aliphatics at 932°F (500°O
Naphthenes
Crack at slower rate than paraffins
If structural groups are equivalent, crack at about the same rate as paraffins
Alkyl-aromatics
Cracked within side chain
Crack next to ring
ource: Venuto [2]
15 percent alumina replaced the natural clay catalysts. The synthetic ilica-alumina catalysts were more stable and yielded superior products. In the mid-1950s, alumina-silica catalysts, containing 25 percent lumina, came into use because of their higher stability. These synthetic atalysts were amorphous; their structure consisted of a random array f silica and alumina, tetrahedrally connected. Some minor improvements in yields and selectivity were achieved by switching to catalysts uch as magnesia-silica and alumina-zirconia-silica.
mpact of Zeolites
The breakthrough in FCC catalyst was the use of X and Y zeolites uring the early 1960s. The addition of these zeolites substantially ncreased catalyst activity and selectivity. Product distribution with a eolite-containing catalyst is different from the distribution with an morphous silica-alumina catalyst (Table 4-3). In addition, zeolites are 1,000 times more active than the amorphous silica alumina catalysts.
130
Fluid Catalytic Cracking Handbook Table 4-3 Comparison of Yield Structure for Fluid Catalytic Cracking of Waxy Gas Oil over Commercial Equilibrium Zeolite and Amorphous Catalysts
Yields, at 80 vol% Conversion
Hydrogen, wt% C1's + C2's, wt%
Amorphous, High Alumina
Zeolite, XZ-25
Change from Amorphous
0.08 3.8
0.04 2.1
-0.04 -1.7
Propylene, vol% Propane, vol% Total C3's
16.1 1.5 17.6
11.8 1.3 13.1
-4.3 -0.02 -4.5
Butenes, vol% -Butane, vol% -Butane, vol% Total C4's
12.2 7.9 0,7 20.8
7.8 7.2 0.4 15.4
-4.4 -0.7 -0.3 -5.4
C5-390 at 90% ASTM asoline, vol%
55.5
62.0
+6.5
Light Fuel Oil, vol% Heavy Fuel Oil, vol% Coke, wt%
4.2 15.8 5.6
6.1 13.9 4.1
+1.9 -1.9 -1.5
Gasoline Octane No.
94
89.8
-4.2
ource: Venuto [2]
The higher activity comes from greater strength and organization of he active sites in the zeolites. Zeolites are crystalline alumina-silicates having a regular pore tructure. Their basic building blocks are silica and alumina tetrahedra. Each tetrahedron consists of silicon or aluminum atoms at the center of he tetrahedron with oxygen atoms at the corners. Because silicon and luminum are in a +4 and 4-3 oxidation state, respectively, a net charge f -1 must be balanced by a cation to maintain electrical neutrality. The cations that replace the sodium ions determine the catalyst's ctivity and selectivity. Zeolites are synthesized in an alkaline environment such as sodium hydroxide, producing a soda-Y zeolite. These oda-Y zeolites have little stability but the sodium can be easily
Chemistry of FCC Reactions
131
exchanged. Ion exchanging sodium with cations, such as hydrogen or rare earth ions, enhances acidity and stability. The most widely used rare earth compounds are lanthanum (La3*) and cerium (Ce3+). The catalyst acid sites are both Bronsted and Lewis type. The catalyst can have either strong or weak Bronsted sites; or, strong or weak Lewis sites. A Bronsted-type acid is a substance capable of donating a proton. Hydrochloric and sulfuric acids are typical Bronsted acids. A Lewis-type acid is a substance that accepts a pair of electrons. Lewis acids may not have hydrogen in them but they are still acids. Aluminum chloride is the classic example of a Lewis acid. Dissolved in water, it will react with hydroxyl, causing a drop in solution pH. Catalyst acid properties depend on several parameters, including method of preparation, dehydration temperature, silica-to-alumina ratio, and the ratio of Bronsted to Lewis acid sites,
Mechanism of Catalytic Cracking Reactions
When feed contacts the regenerated catalyst, the feed vaporizes. Then positive-charged atoms called carbocations are formed. Carbocation is a generic term for a positive-charged carbon ion. Carbocations can be either carbonium or carbenium ions. A carbonium ion, CH5+, is formed by adding a hydrogen ion (H+) to a paraffin molecule (Equation 4-6), This is accomplished via direct attack of a proton from the catalyst Bronsted site. The resulting molecule will have a positive charge with 5 bonds to it. R — CH2 — CH2 — CH2 — CH3 + H+ (proton attack) -» R — C+H — CH2 — CH2 — CH3 + H2
(4-6)
The carbonium ion's charge is not stable and the acid sites on the catalyst are not strong enough to form many carbonium ions. Nearly all the cat cracking chemistry is carbenium ion chemistry. A carbenium ion, R-CH2+, comes either from adding a positive charge to an olefin or from removing a hydrogen and two electrons from a paraffin (Equations 4-7 and 4-8). R — CH. = CH — CH2 — CH2 — CH3 + H* (a proton @ Bronsted site) —>_^ jp^ ~™_™. ^^ |"j
« v-'Jcin ——• v^JrJ'-j"""
v^in.'-) —"""— v^-ii-t
\£|.— / j
32
Fluid Catalytic Cracking Handbook
R — CH2 -— CH2 — CH2 — CH3 (removal of H~ @ Lewis site) _» R _ c+H — CH2 — CH2 — CH3
(4-8}
Both the Bronsted and Lewis acid sites on the catalyst generate arbenium ions. The Bronsted site donates a proton to an olefin molecule and the Lewis site removes electrons from a paraffin moleule. In commercial units, olefins come in with the feed or are prouced through thermal cracking reactions. The stability of carbocations depends on the nature of alkyl groups tached to the positive charge. The relative stability of carbenium ions as follows [2] with tertiary ions being the most stable: Tertiary .
C ""~ V.-C+
'"•"""" V--
> P
V_^
Secondary P
\*s
P+
V~"
P
V-"'
>
Primary R
JLX.
P
V-'
> Ethyl > Methyl P+
V_--
P
V--
P+
*—•
P* V,,'
c
One of the benefits of catalytic cracking is that the primary and econdary ions tend to rearrange to form a tertiary ion (a carbon with hree other carbon bonds attached). As will be discussed later, the ncreased stability of tertiary ions accounts for the high degree of ranching associated with cat cracking. Once formed, carbenium ions can form a number of different eactions. The nature and strength of the catalyst acid sites influence he extent to which each of these reactions occur. The three dominant eactions of carbenium ions are: * The cracking of a carbon-carbon bond * Isomerization * Hydrogen transfer
Cracking Reactions
Cracking, or beta-scission, is a key feature of ionic cracking. Betacission is the splitting of the C-C bond two carbons away from the ositive-charge carbon atom. Beta-scission is preferred because the nergy required to break this bond is lower than that needed to break he adjacent C-C bond, the alpha bond. In addition, short-chain hydroarbons are less reactive than long-chain hydrocarbons. The rate of
Chemistry of FCC Reactions
133
he cracking reactions decreases with decreasing chain length. With short chains, it is not possible to form stable carbenium ions. The initial products of beta-scission are an olefin and a new carbenium on (Equation 4-9). The newly-formed carbenium ion will then continue a series of chain reactions. Small ions (four-carbon or five-carbon) can ransfer the positive charge to a big molecule, and the big molecule can crack. Cracking does not eliminate the positive charge; it stays until two ions collide. The smaller ions are more stable and will not crack, They survive until they transfer their charge to a big molecule, R _ " V,' ri+"H 11 IV. "
CH V*' 1. !••->
PH V--' 1 !••) — PH V--J. .I')
"
CH \.~^ I to
-* CH3 — CH = CH2 + C+H2 — CH2 — CH2R
(4-9)
Because beta-scission is mono-molecular and cracking is endohermic, the cracking rate is favored by high temperatures and is not equilibrium-limited.
somerization Reactions
Isomerization reactions occur frequently in catalytic cracking, and nfrequently in thermal cracking. In both, breaking of a bond is via beta-scission. However, in catalytic cracking, carbocations tend to earrange to form tertiary ions. Tertiary ions are more stable than secondary and primary ions; they shift around and crack to produce branched molecules (Equation 4-10). (In thermal cracking, free radicals yield normal or straight chain compounds.)
CH3 — CH, -— C+H — CH, — CH2R -» CH3 — C+ — CH — CH2R H
CR
or CH — CH2 — CH2R
(4-10) Some of the advantages of isomerization are:
34
Fluid Catalytic Cracking Handbook
* Higher octane in the gasoline fraction. Isoparaffins in the gasoline boiling range have higher octane than normal paraffins. * Higher-value chemical and oxygenate feedstocks in the C3/C4 fraction. Isobutylene and isoamylene are used for the production of methyl tertiary butyl ether (MTBE) and tertiary amyl methyl ether (TAME). MTBE and TAME can be blended into the gasoline to reduce auto emissions. * Lower cloud point in the diesel fuel. Isoparaffins in the light cycle oil boiling range improve the cloud point.
Hydrogen Transfer Reactions
Hydrogen transfer is more correctly called hydride transfer. It is a imolecular reaction in which one reactant is an olefin. Two examples re the reaction of two olefins and the reaction of an olefin and naphthene. In the reaction of two olefins, both olefins must be adsorbed on ctive sites that are close together. One of these olefins becomes a araffin and the other becomes a cyclo-olefin as hydrogen is moved om one to the other. Cyclo-olefin is now hydrogen transferred with nother olefin to yield a paraffin and a cyclodi-olefin. Cyclodi-olefin will then rearrange to form an aromatic. The chain ends because romatics are extremely stable. Hydrogen transfer of olefins converts hem to paraffins and aromatics (Equation 4-11). 4 CrlH2n -» 3 Cn H2n+2 + CnH2n^ olefins
—> paraffins
+ aromatic
(4-11)
In the reaction of naphthenes with olefins, naphthenic compounds re hydrogen donors. They can react with olefins to produce paraffins nd aromatics (Equation 4-12). 3 C n H 2n + CmH2m
-» 3 Cn H 2n+2
olefins
—> paraffins
+ naphthene
+ Cm H2m^6 + aromatic
(4-12)
A rare-earth-exchanged zeolite increases hydrogen transfer reactions. n simple terms, rare earth forms bridges between two to three acid ites in the catalyst framework. In doing so, the rare earth protects
Chemistry of FCC Reactions
135
hose acid sites. Because hydrogen transfer needs adjacent acid sites, bridging these sites with rare earth promotes hydrogen transfer reactions. Hydrogen transfer reactions usually increase gasoline yield and stability. The reactivity of the gasoline is reduced because hydrogen ransfer produces fewer olefins. Olefins are the reactive species in gasoline for secondary reactions. Therefore, hydrogen transfer reactions indirectly reduce "overcracking'1 of the gasoline. Some of the drawbacks of hydrogen transfer reactions are: • * • *
Lower gasoline octane Lower light olefin in the LPG Higher aromatics in the gasoline and LCO Lower olefin in the front end of gasoline
Other Reactions
Cracking, isomerization, and hydrogen transfer reactions account for he majority of cat cracking reactions. Other reactions play an imporant role in unit operation. Two prominent reactions are dehydrogenation and coking.
Dehydrogenation. Under ideal conditions (i.e., a "clean" feedstock and a catalyst with no metals), cat cracking does not yield any appreciable amount of molecular hydrogen. Therefore, dehydrogenation eactions will proceed only if the catalyst is contaminated with metals uch as nickel and vanadium.
Coking. Cat cracking yields a residue called coke. The chemistry of coke formation is complex and not very well understood. Similar o hydrogen transfer reactions, catalytic coke is a "bimolecular" eaction. It proceeds via carbenium ions or free radicals. In theory, coke yield should increase as the hydrogen transfer rate is increased. t is postulated [4] that reactions producing unsaturates and multi-ring aromatics are the principal coke-forming compounds. Unsaturates such as olefins, diolefins, and multi-ring polycyclic olefins are very reactive and can polymerize to form coke. For a given catalyst and feedstock, catalytic coke yield is a direct unction of conversion. However, an optimum riser temperature will minimize coke yield. For a typical cat cracker, this temperature is
36
Fluid Catalytic Cracking Handbook
bout 950°F (510°C). Consider two riser temperatures, 850°F and ,050°F (454°C and 566°C), at the extreme limits of operation. At 50°F, a large amount of coke is formed because the carbenium ions o not desorb at this low temperature. At 1,050°F (566°C), a large mount of coke is formed, largely due to olefin polymerization. The minimum coking temperature is within this range.
THERMODYNAMIC ASPECTS
As stated earlier, catalytic cracking involves a series of simultaneous eactions. Some of these reactions are endothermic and some are xothermic. Each reaction has a heat of reaction associated with it Table 4-4). The overall heat of reaction refers to the net or combined eat of reaction. Although there are a number of exothermic reactions, he net reaction is still endothermic. The regenerated catalyst supplies enough energy to heat the feed o the riser outlet temperature, to heat the combustion air to the flue as temperature, to provide the endothermic heat of reaction, and to ompensate for any heat losses to atmosphere. The source of this nergy is the burning of coke produced from the reaction. It is apparent that the type and magnitude of these reactions have an mpact on the heat balance of the unit. For example, a catalyst with less ydrogen transfer characteristics will cause the net heat of reaction to be more endothennic. Consequently this will require a higher catalyst circuation and, possibly, a higher coke yield to maintain the heat balance.
UMMARY
Although cat cracking reactions are predominantly catalytic, some onselective thermal cracking reactions do take place. The two proesses proceed via different chemistry. The distribution of products learly confirms that both reactions take place, but that catalytic eactions predominate. The introduction of zeolites into the FCC catalyst in the early 1960s rastically improved the performance of the cat cracker reaction roducts. The catalyst acid sites, their nature, and strength have a major influence on the reaction chemistry. Catalytic cracking proceeds mainly via carbenium ion intermediates. he three dominant reactions are cracking, isomerization, and hydrogen
n-C10H22 -> n-C7H16 + C3H6 1~C8H16 -> 2C4Hg 4C6H12 -» 3C6H14 + C6H6 cyclo-C6Hl2 + 3 1-C5H!0 -> 3n-C5H12 + C6H6 1-C4H8 -» trans-2-C4H8 n-C6H10 -» iso-C4H10 o-C6H4(CH3)2 -> m-C6H4(CH3)2 cyclo-C6H12 -» CH3-cyclo-C5H9 C6H6 + m-C6H4(CH3)2 -> 2C6H5CH3 1-C7H14 -» CH3-cyclo-C6H11 iso-C3H7-C6H5 -> C6H6 + C3H6 n-C6H14 ^ 1-C6H12 + H2 3C2H4 —> 1-C6H12 1-C4H8 + iso-C4H10 -> iso-C8H18
Specific Reaction 850°F
2.46 2.10 11.09 10.35 0.25 -0.23 0.30 1.09 0.65 1.54 0.88 -1.52 — —
950°F
— 1.05 — -1.2 3.3
980°F
— 1.10 0.65
— 0.09 -0.36
— 2.23 —
2.04 1.68 12.44 11.22 0.32 -0.20 0.33 1.00 0.65 2.11 0.41 -2.21 — —
Log KE (equilibrium constant)
H
Table 4-4 Some Thermodynamic Data for Idealized Reactions of Importance in Catalytic Crac
n Class
g
n transfer
ation
ylation ion ation genation ization Alkylation
enuto [2]
38
Fluid Catalytic Cracking Handbook
ransfer. Finally, the type and degree of reactions occurring will nfluence the unit heat balance.
REFERENCES*
. Gates, B. C., Katzer, J. R., and Schuit, G. G., Chemistry of Catalytic Processes. New York: McGraw-Hill, 1979. . Venuto, P. B. and Habib, E. T., Fluid Catalytic Cracking with Zeolite Catalysts. New York: Marcel Dekker, Inc., 1979, . Broekhoven, E. V. and Wijngaards, H., "Investigation of the Acid Site Distribution of FCC Catalysts with Ortho-xylene as a Model Compound," 1988 Akzo Chemicals FCC Symposium, Amsterdam, The Netherlands, . Koerroer, G. and Deeba, M., "The Chemistry of FCC Coke Formation," Engelhard Corporation, The Catalyst Report, Vol. 7, Issue 2, 1991.
he author also expresses appreciation to Messrs. Terry Reid of Akzo Nobel and Tom abib of Davison Div., W. R. Grace & Co., for their many helpful comments.
CHAPTER 5
Unit Monitoring and Control
The only proper way to monitor the performance of a cat cracker s by periodic material and heat balance surveys on the unit. By arrying out these tests frequently, one can collect, trend, and evaluate he unit operating data. Additionally, meaningful technical service to ptimize the unit operation should be based on regular test runs. Understanding the operation of a cat cracker also requires in-depth nowledge of the unit's heat balance. Any changes to feedstock quality, perating conditions, catalyst, or mechanical configuration will impact he heat balance. Heat balance is an important tool in predicting and valuating the changes that will affect the quantity and the quality of CC products. Finally, before the unit can produce one barrel of product, it must irculate catalyst smoothly. One must be familiar with the dynamics f pressure balance and key process controls. The main topics discussed in this chapter are: • • • •
Material Balance Heat Balance Pressure Balance Process Control Instrumentation
n the material and heat balance sections, the discussions include: • Two methods for performing test runs • Some practical steps for carrying out a successful test run • A step-by-step method for performing a material and heat balance survey • An actual case study
139
40
Fluid Catalytic Cracking Handbook
n the pressure balance section, the significance of the pressure balance n debottlenecking the unit is discussed. Finally, fundamentals of both basic" and "advanced" process controls are presented. This chapter presents the entire procedure for performing heat and weight balances. The last section of the chapter discusses the use of he distributed control system and computer in automating the process,
MATERIAL BALANCE
Complete data collection should be carried out weekly. Since changes n the unit are continuous, regular surveys permit distinction among he effects of feedstock, catalyst, and operating conditions. An accurate ssessment of a cat cracker operation requires reliable plant data. A easonable weight balance should have a 98% to 102% closure. In any weight balance exercise, the first step is to identify the input nd output streams. This is usually done by drawing an envelope(s) round the input and output streams. Two examples of such envelopes re shown in Figure 5-1. One of the key pieces of data is the composition of products leaving he reactor. The reactor effluent vapor entering the main fractionator ontains hydrocarbons, steam, and inert gases. By weight, the hydroarbons in the reactor overhead stream are equal to the fresh feed plus ecycle minus the portion of the feed that has been converted to coke. the feed can contain water, it should be analyzed for and corrected. The sources of steam in the reactor vapor are: lift steam to the andpipe, atomization steam to the feed nozzles, dome steam, and ripping steam. Some units may have other streams and the feed may ontain water. Depending on the reactor pressure, approximately 25% o 50% of the stripping steam is entrained with the spent catalyst owing to the regenerator, which should be deducted. Inert gases such as nitrogen and carbon dioxide enter the riser ntrained with the regenerated catalyst. The quantity of these inert asses is directly related to catalyst circulation rate. These gases flow hrough the gas plant and leave the unit with the off-gas from the ponge oil absorber column. They are not significant for the weight alance, but they are usually the only source of inerts in the off-gas nd should be deducted. FCC products are commonly reported, on an inert-free basis, as the olume and weight fractions of the fresh feed. In a rigorous weight
Unit Monitoring and Control
141
External Streams^-"
Figure 5-1. FCC unit input/output streams.
balance, gasoline and light cycle oil (LCO) yields and unit converion are reported based on fixed end points. The common end points are 430°F (221 °C) TBP for gasoline and 700°F TBP for LCO, Other popular cut points are 430°F (221°C) ASTM D-86 for gasoline and 650°F (343°C) or 670°F (354°C) ASTM D-86 for LCO. Using fixed
42
Fluid Catalytic Cracking Handbook
ut points isolates the reactor system from the distillation sysem performance. Conversion is defined as the volume or weight percent of feedstock onverted to gasoline and other lighter products, including coke. However, conversion is typically calculated by subtracting the volume ercent or weight percent of liquid products heavier than gasoline from resh feed, and dividing by the volume or weight of fresh feed. This s shown as follows: ~ , Feed - (light cycle oil + heavy cycle oil + decanted oil) . ,,„ Conversion m% = ^-^ —x 100
Depending on seasonal demands, the gasoline end point can range rom 380°F to 450°F (193°C to 232°C). Undercutting of gasoline ncreases the LCO product and can appear as low conversion. Thereore, it is necessary to distinguish between the apparent and true onversion. The apparent conversion is calculated before the gasoline nd point adjustment is made, and the true conversion is calculated fter the adjustment.
Testing Methods
The material balance around the riser requires the reactor effluent omposition. Two techniques are used to obtain this composition. Both echniques require that the coke yield be calculated. The first technique is to draw an envelope with the reactor effluent s the inlet stream and the product flows as the outlet streams. Streams rom other units must be included. The flow rates and composions of the entering and leaving streams are then totaled. The net is he reactor effluent. This is the method practiced by most refiners. The second technique involves direct sampling of the reactor effluent Figure 5-2). In this technique, a sample of reactor effluent is collected n an aluminized polyester bag for separation and analysis. There are several advantages and disadvantages to reactor effluent ampling;
dvantages of Reaction Mix Sampling • Allows data gathering on different sets of conditions without waiting for the recovery side to equilibrate.
lop container
Sample probe
Figure 5-2,
Reaction mix sampling [2].
Gate and ball valves
44
Fluid Catalytic Cracking Handbook
* Eliminates concern about rate and compositions of extraneous streams entering the gas plant because they are not included in the overall balance. * Eliminates concern about correcting for end points because the effluent sample is cut at the desired TBP end point. * Eliminates concern about obtaining a 100% weight balance.
Disadvantages of Reaction Mix Sampling * Possible leaks during sampling. * Possible inaccurate measurement of volume of gas and weight of liquid. * Requires qualified individuals to perform the test. » Requires separate lab to perform analyses. * Can require special procedures and be expensive.
Recommended Procedures for Conducting a Test Run
A successful test run requires a clear definition of objectives, careful lanning, and proper interpretation of the results. The following steps an be used as a guide to ensure a smooth and successful test run,
rior to the Test Run 1. Issue a memo to the involved departments: operations, laboratory, maintenance, and oil movement. Communicate the purpose, duration, and scope of the test run. Include a list of samples and the required analyses (Table 5-1). 2, Inform the units feeding the FCC. The composition of FCC feedstock should remain relatively constant during the test run. Flow meters should be zeroed and calibrated. Sample taps should be checked, particularly those that are not used regularly. 5, The sample bombs used to collect gas and LPG products should be purged, marked, and ready.
ata Collection 1. The duration of a test run is usually 8 to 12 hours. 2. Operating parameters should be specified. It should be documented which constraints (i.e., blower, wet gas compressor, etc.) the unit is operating against.
Unit Monitoring and Control
145
Table 5-1 Typical Laboratory Analysis of FCC Streams
Tests °API D-86 D-1160
Gas Oil
Sulfur
Viscosity
Metals
/
/
GC
/
/
/
Slurry Recycle
/
/
/
Decanted Oil Product
/
/
/
LCO Product
/
/
/
Gasoline Product
/
/
/
/
/
/
Feedstock /
LPG
C.,\s and C4's
Tail Gas
/
3. The sample taps must be bled adequately before samples are collected. A reliable flue gas analysis is important; an extra sample can be collected. The laboratory should retain the unused samples until all analyses are verified. 4. Pertinent operating data must be collected. A form similar to the one shown in Table 5-2 can be used to gather the data.
Mass Balance Calculations 1. The orifice plate meter factor should be adjusted for actual operating parameters. For liquid streams, the flow meters should be adjusted for °API gravity, temperature, and viscosity. For gas streams, the flow rate should be adjusted for the operating temperature, pressure, and molecular weight. 2. Chromatographs of each stream must be normalized to 100%. The GC of the off-gas must include accurate analysis of hydrogen, 3. The coke yield should be calculated using air rate and flue gas composition.
46
Fluid Catalytic Cracking Handbook Table 5-2 Operating Data
eed and Product Rates resh Feed Rate oker Off Gas CC Tail Gas PG asoline CD O
Other Pertinent Flow Rates ispersion Steam eactor Stripping Steam eactor Dome Steam ir to Regenerator
50,000 bpd (331 nrVhr) 3,000,000 scfd (3,540 m3/hr) 16,000,000 scfd (18,878 m3/hr) 11,565 bpd (77 mVhr) 30,000 bpd (199 m3/hr) 10,000 bpd (66 m3/hr) 3,000 bpd (20 mVhr) 9,000 Ib/hr (4,082 kg/hr) 13,000 Ib/hr (5,897 kg/hr) 1,200 Ib/hr (544 kg/hr) 90,000 scf/min (152,912 m3/hr)
emperature,°F/°C iser Inlet iser Outlet lower Discharge egen. Dense Phase egen. Flue Gas mbient
594/312 972/522 374/190 1,309/709 1,330/721 80/27
ressure, psig/Kp lower Discharge egen. Dome eactor Dome egenerated Catalyst Slide Valve, AP pent Catalyst Slide Valve, AP
43/296 34/234 33/227 5.8/40 6.0/41
lue Gas Analysis, Mol% O, O, O O2 2 + AT Miscellaneous Data elative Humidity resh Catalyst Makeup -Cat MAT
1.5 15.4 0.0
500 ppm -> 0.05 mol% 83.05 80% 4 tons/day 68%
Unit Monitoring and Control
147
4. The flow rate of each stream should be converted to weight units. 5. The quantity of inert gases and extraneous streams should be subtracted from the FCC gas plant products. 6. The raw mass balance should be reported, including the error, Then the feed/products should be normalized to 100%. The error will be distributed in proportion to flow rates or a known inaccurate meter will be adjusted. 7. Gasoline and LCO rates will be adjusted to standard cut points. 8. The feed characterization correlations discussed in Chapter 2 should be used to determine the composition of fresh feed.
Analysis of Results 1. The yields and quality of the desired products should be reported and compared with the unit targets. 2. The results of this test run should be compared with the results of previous test runs; any significant changes in the yields and/ or operating parameters should be highlighted. 3. The final step is to perform simple economics of the unit operation and make recommendations that improve short- and longterm unit operation.
The following case study demonstrates a step-by-step approach to performing a comprehensive material and heat balance.
A test run is conducted to evaluate the performance of a 50,000 bpd 331 m3/hr) FCC unit. The feed to the unit is gas oil from the vacuum nit. No recycle stream is processed; however, the off-gas from the elayed coker is sent to the gas recovery section. Products from he unit are fuel gas, LPG, gasoline, LCO, and decanted oil (DO). Tables 5-2 and 5-3 contain stream flow rates, operating data, and aboratory analyses. The meter factors have been adjusted for actual perating conditions. The mass balance is performed as follows: 1. Identification of the input and output streams used in the overall mass balance equation. 2. Calculation of the coke yield.
48
Fluid Catalytic Cracking Handbook Table 5-3 Feed and Product Inspections
Feed
API Gravity ulfur, Wt% Analine Point, °F/ °C RI @ 67°C Viscosity, SSU @ 150°F (65.5°C) @ 210°F (98.9°C) Distillation, °F Vol% 10 30 50 70 90 EP
Component
H, CH4 C, 2= C3
,=
C4 NC4 C4 C5+ H2S N2
O2
otal p. Gravity
Decanted Oil
Gasoline
LCO
58.5
21.5
2.4
D-86
D-86
D-1160
125 160 213 285 369 433
477 514 547 576 627 666
646 687 720 771 846 1,055
25.2 0.5 208/97.8 1.4854 109 54 D-1160 682 766 835 901 1,001 1,060
Mole% Composition of FCC Gas Plant Streams
FCC Tail Gas
15.5 35.8 17.1 11.0 1.6 4.7 0.7 0.2 1.3 1.0 2.1 7.2 1.8 100.0 0.78
LPG
17.9 31.3 16.1 10.9 23.8
100.0 0.55
FCC Gasoline
0.4 2.0 4.4 93.2
100.0
Coker Off-Gas
8.0 47.2 14.9 2.5 8.4 4.4 0.9 3.2 3.4 4.9 2.0 0.2 100.0 0.96
Unit Monitoring and Control
3, 4, 5, 6,
149
Conversion of the flow rates to weight units (e.g., Ib/hr). Normalization of the data to obtain a 100% weight balance. Determination of the component yields. Adjustment of the gasoline, LCO, and decanted oil yields to standard cut points.
nput and Output Streams in the Overall Mass Balance
As shown in Envelope 1 of Figure 5-1, the input hydrocarbon streams are fresh feed and coker off-gas. The output streams are FCC ail gas (minus inerts), LPG, gasoline, LCO, DO, and coke.
Coke Yield Calculations
As discussed in Chapter 1, a portion of the feed is converted to coke n the reactor. This coke is carried into the regenerator with the spent catalyst. The combustion of the coke produces H2O, CO, CO2, SO2, and traces of NOx. To determine coke yield, the amount of dry air to he regenerator and the analysis of flue gas are needed. It is essential o have an accurate analysis of the flue gas. The hydrogen content of coke relates to the amount of hydrocarbon vapors carried over with he spent catalyst into the regenerator, and is an indication of the eactor-stripper performance. Example 5-1 shows a step-by-step calulation of the coke yield. Example 5-1 Determination of the Unit's Coke Yield
Given: Wet air = 90,000 SCFM, Relative Humidity = 80%, Ambient Temperature = 80°F (26.7°C)
Figure 5-3 can be used to obtain percent dry air as a function of ambient emperature and relative humidity. For this example, the percentage of dry ir is 97.1% or: A- = Ami 90,000SCF x Imole x 60 Min = ,,,,., ,. • nDry Air 0.971 x — 13,817 moles/hr Min 379.5 SCF 1 hr
Flue gas rate (dry basis) is calculated from the dry air rate using nitrogen nd argon as tie elements.
50
Fluid Catalytic Cracking Handbook
„ ,, , . N (13,817 moles/hrx 0.7901) i a i _ , „ * Flue gas rate (dry = 13,145 moles/hr J basis)= 0.8305
.7901 and 0.8305 are concentrations of (nitrogen + argon) in atmospheric ry air and flue gas (from analysis), respectively.
he flow rates of each component in the flue gas stream are: « * * *
O2 out = 0.015 x 13,145 moles/hr = 197 moles/hr CO2 out = 0.154 x 13,145 moles/hr = 2,024 moles/hr SOj out = 0.0005 x 13,145 moles/hr = 7 moles/hr (N2~ + Ar) out = 0.8305 x 13,145 moles/hr = 10,917 moles/hr
n oxygen balance can be used to calculate water formed by the combuson of coke: * O2 out = 197 + 2,024 +7 = 2,228 moles/hr * 1)3 in = 0.2095 x 13,817 moles/hr = 2,895 moles/hr * O2 used for combustion of hydrogen = 2,895 - 2,228 = 667 raoles/hr
ince for each mole of O2, two moles of water are formed, the amount of ater is: * H2O formed = 667 x 2 = 1,334 moles/hr
omponents of coke are carbon, hydrogen, and sulfur. Their rates are calcuted as follows: * * « *
Carbon = 2,024 moles/hr x 12 Ibs/mole = 24,288 Ibs/hr Hydrogen = 1,334 moles/hr x 2.02 Ibs/mole = 2,695 Ibs/hr Sulfur = 7 moles/hr x 32.1 Ibs/moles = 225 Ibs/hr Coke = 24,288 + 2,695 + 225 = 27,208 Ibs/hr
* H-, content of coke, wt% = —: — x 100 = 9.9 27,231 Ibs/hr The hydrogen content of coke indicates the amount of hydrocarbon vapors arried through the stripper with the spent catalyst.
Conversion to Unit of Weight, Ibs/hr
The next step is to convert the flow rate of each stream in the verall mass balance equation to the unit of weight (e.g., Ibs/hr). xample 5-2 shows these conversions for gas and liquid streams.
Figure 5-3,
50
80
100
Dry air versus relative humidity and temperature.
Temperature ,Deg F
70
Dry Air versus Relative Humidity & Temperature
1tO
52
Fluid Catalytic Cracking Handbook
Example 5-2 Conversion of Input and Output Streams to the Unit of Weight (Ib/hr)
„ , „ . 50,000bbl 1day 141.5 350.3 Ib • Fresh Feed = —-—-- x - - x • -- x - —— day 24 hr (131.5 + 25.2) bbl = 658,964 lb/hr »
„,
3,000,000 SCF A.V 1day 1mole 27.8 lbs ni-—«___Q« 1|«,„,,. , •* . *„._„,„_ _^ V _. ,. , . . . . u _.-_...-... Sf\ x1U/m/nr A. v ._.....,.._L..._, JvJ.O 111 day 24 hr 379.5 SCF 1mole
f^r\lrpir gas CTQG — v^UJVCI —*"_!_ " !_ ........
.™ .. 16,000,000 SCF Iday Imole • FCC tail gas = —-- -x - ^x day 24 hr 379.5 SCF
22.6 Ibs x Imole
= 39,701 Ib/hr
he amount of inerts in the FCC tail gas is: 16,000,OOOSCF 1day _ _ _ „ Imole KT •N x - i- x 0.072 x 2 = —-- -day 24 hr 379.5 SCF „ 16,000,OOOSCF A n - t 1day 1mole . CO2 = —!- -x 0.021 x - ^-x - day 24 hr 379.5 SCF
281bs 3,542lb/hr x= 3,542 Ib/hr Imole 441bs , , » « , , „ x= 1,623 Ib/hr Imole
• Inert-free FCC tail gas = 39,701 - (3,542 + 1,623) = 34,537 Ib/hr .
LpG= H.565bbl x lday x
day
24 hr
141.5 X35031b = (131.5 + 123.5) bbl
„ r 30,000bbl day 141.5 350.31b • Gasoline = —-- x --x x day 24 hr (131.5 + 58.5) bbl = 326, 102 Ib/hr
day
24hr
141.5 (131.5 + 21.5)
f
bbl
3,000bbl 1day 141.5 350.31b ., «--,.,, = —x - i-x x -- = 46,2731b/hr day 24 hr (131.5 + 2.4) bbl
Unit Monitoring and Control
153
Normalization of the Data
Because a preliminary weight balance seldom has a 100% closure, t is necessary to normalize the yield to obtain a 100% weight balance, Example 5-3 shows the preliminary overall weight balance.
Example 5-3 Preliminary Overall Weight Balance
nput = Fresh Feed + Coker Off-Gas Output = FCC tail gas + LPG + Gasoline + LCO + DO + Coke • Input = 658,814 + 9,182 = 667,996 Ib/hr • Output = 34,617 + 93,656 + 326,124 + 134,973 + 46,270 + 27,231 = 662,871 lb/hr « Difference = 667,996 - 662,871 = 5,125 lb/hr
Error in mass balance = 0.8 wt%
The products are adjusted upward in proportion to theilr rates to obtain a 100% weight balance. The normalized rates: • • • • « •
Tail gas LPG Gasoline LCO DO Coke
= 34,883 Ib/hr = 94,460 lb/hr = 328,766 Ib/hr = 136,054 lb/hr = 46,626 Ib/hr = 27,440 Ib/hr
= = = =
11,658 bpd 30,230 bpd 10,077 bpd 3,023 bpd
Component Yield
The reactor yield is then determined by performing a component balance. The amount of C5+ in the gasoline boiling range is calculated by subtracting the C4 and lighter components from the total gas plant products. Example 5-4 shows the step-by-step calculation of the component yields. The summary of the results, normalized but unadusted for the cut points is shown in Table 5-4.
54
Fluid Catalytic Cracking Handbook Example 5-4 Calculation of Individual Components 0.155 x 16 MMSCFDx 2.02 ~~" 379.5x24
0.08 x 3 MMSCFDx 2.02 379,5x24
=
2
_„ 0.358x16 MMSCFDx 16 0.472x3.0 MMSCFDx 16 - C M 1 U r t CH 44 = --= 7,585 Ib/hr 379.5x24 379.5x24 C 2 ~~~ = 0-171 xl6MMSCFDx30 379.5x24 /„ _ 0.11x16 MMSCFDx 28 2
379.5x24
0.149 x 3 MMSCFDx 30 379.5x24 0.025 x 3 MMSCFD x 28 379.5x24
0.016x!6MMSCFDx44 | 0.179x 11,65]8BPDx 175.3 + ~~ 379.5x24 " 24 =
3
0.084x3 MMSCFDx 44 = 15,262 lb/hr 379.5 x 24 ^_ 0.047x16 MMSCFDx 42 0.313x11,658 BPDx 181.8 • Cr3 = + 379.5 x 24 24 0.044x3 MMSCFDx 42 O A C A / I I U / , = 30,504 Ib/hr 379.5 x 24 '
4
^0.002x16MMSCFDx58 "~ 379.5x24
| +
0.109x11,658BPDx204.6 24
0.02 x 30,230 x 204.6 MMSCFD x 42 0.032 x 3 MMSCFD x 58 24 379.5x24 =15,579 Ib/hr 0.007x16MMSCFDx58 '~ 379.5x24 =
4
[ +
0.161x11658BPDx 197.2 24
0.004x30,230x204.6 BPDx 197.2 24 = 16,95 8 Ib/hr
0.009x3 MMSCFDx 58 379.5x24
Unit Monitoring and Control
155
0.013xl6MMSCFDx56 0.238x1 l,658BPDx213.4 + 379.5x24 24 0.044x30,230x213.4 0.034x3MMSCFDx56 = 37,150 Ib/hr 24 379.5x24
Table 5-4
Normalized FCC Weight Balance Summary with Coker Gas Subtracted
Stream
Fresh Feed
bpd
ib/hr
Vol% of Feed
Wt% of Feed
50,000
658,814
100.00
100.00
Products
H7 C,
C,
c; Total C2 and lighter
H2S
C3
C?
C4 NC4
c;
Gasoline (Cs+)
LCO
DO Coke Total Apparent Conversion nerts
0.07 1.15 1.15 0.79 3.16
497 7,585 7,549 5,187 20,818
2,090 4,027 2,064 1,827 4,178
1,032 15,262 30,504 16,958 15,579 37,150
4.18 8.05 4.13 3.65 8.36
0.16 2.32 4.63 2.57 2.36 5.64
28,650 10,077
311,437 136,008
57.30 20.15
47.27 20.64
3,023
46,626 27,440 658,814
6.05
7.08 4.17 100.00 72.28
55,936
5,143
111.87 73.8
56
Fluid Catalytic Cracking Handbook
djustment of Gasoline and LCO Cut Points
As discussed earlier in this chapter, gasoline and LCO yields are enerally corrected to a constant boiling range basis. The most commonly used bases are 430°F TBP gasoline and 640°F TBP LCO end oints. Since TBP distillations are not routinely performed, they are ften estimated from the D-86 distillation data. The adjustments to the nd points involve the following: * Adding to the raw LCO all the 430°F+ in the raw gasoline and subtracting the 430°F in the LCO stream. « Adding to the raw LCO all the 650°F~ in the raw decanted oil and subtracting the 650°F~ in the decant oil stream. * Adding to the raw gasoline all the 430°F~ in the raw LCO and subtracting the 430°F* in the gasoline stream. • Adding to the raw decanted oil all the 650°F+ in the raw LCO and subtracting the 650°F~ in the decant oil stream.
Table 5-5 illustrates steps used to convert ASTM D-86 data to TBP. he laboratory usually converts D-1160 and reports the data as D-86, xtrapolation of the TBP data indicates the following: « « « •
The The The The 514
430°F+ content of the FCCU gasoline is 3 vol%, or 859 bpd. gasoline (430°F~) content of LCO is 8 vol%, or 806 bpd. 650°F+ content ofLCO is 12 vol%, or 1,209 bpd. LCO (650°F~) content of the decanted oil is 17 vol%, or bpd.
herefore, the adjusted rates are as follows: Gasoline (C5+ to 430°F TBP end point) = 28,650 - 859 + 806 = 28,597 bpd LCO (430°F to 650°F TBP end point) = 10,077 + 514 - 1,209 - 806 + 859 = 9,435 bpd DO (650°F+) = 3,023 + 1,209 - 514 = 3,718 bpd
able 5-6 shows the normalized FCC weight balance with the adjusted ut points.
Unit Monitoring and Control
Table 5-5 Conversion of ASTM Distillation to TBP Distillation for Gasoline, LCO, and Decanted Oil Gasoline TBP (From Appendix 9, TBP 50% point = 213°F) Given D-86 50% 30% 10% 70% 90% EP -
- 30% = - 10% = - IBP = - 50% = - 70% = 90% =
From Appendix 10 53°F 35°F 25°F 72°F 84°F 64°F
30% TBP = 10% TBP = IBP TBP = 70% TBP = 90% TBP = EP TBP =
140°F 77°F 26°F 297°F 383°F 501°F
LCO TBP (From Appendix 9: TBP 50% point = 561°F) Given D-86 50% - 30% = 30% - 10% = 10% - IBP = 70%-50% = 90% - 70% = EP-90% =
From Appendix 10 33°F 4FF 73°F 29°F 51°F 39°F
30%TBP = 511°F 10% TBP = 441°F IBP TBP = 343°F 70%TBP = 601°F 90% TBP = 660°F EPTBP = 712°F
Decanted Oil TBP (From Appendix 9: TBP 50% point = 744°F) Given D-86
From Appendix 10
50% 30% 10% 70% 90%
30% TBP = 694°F 10% TBP = 624°F IBP TBP = 425°F 70% TBP = 807°F 90% TBP = 886°F
-
30% = 33°F 10% = 41°F IBP = 236°F 50% = 51°F 70% = 75°F
157
58
Fluid Catalytic Cracking Handbook Table 5-6 Normalized and Adjusted FCC Weight Balance Summary
Stream
Fresh Feed
bpd
ib/hr
Vol% of Feed
Wt% of Feed
50,000
658,814
100.00
100,00
Products
497 7,585 7,549 5,187 20,818
H, C, C,
= Total C2 and lighter
H 2S C,=
C
3
C4 NC4
c;
Gasoline (C5+ to 430°F TBP)
CO (430°F TBP to 650°F TBP)
DO (65Q°F+ TBP)
Coke Total
rue Conversion nerts
0.07 1.15 1.15 0.79 3.16
2,090 4,027 2,064 1,827 4,178
1 ,032 15,262 30,504 16,958 15,579 37,150
4.18 8.05 4.13 3.65 8.36
0.1.6 2 32 4.63 2.57 2.36 5.64
28,597
312,073
57.19
47,37
9,435
1 26,004
18.87
19.13
3,718
55,994
7.44
8.50
55,936
27,440 658,814
111.87
4.17 100.00
73.7
72.3
5,143
HEAT BALANCE
A cat cracker continually adjusts itself to stay in heat balance. This means that the reactor and regenerator heat flows must be equal Figure 5-4). Simply stated, the unit produces and burns enough coke o provide energy to:
Unit Monitoring and Control
Steam
Steam Oil Feed Figure 5-4. Reactor-regenerator heat balance.
159
60
Fluid Catalytic Cracking Handbook
• Increase the temperature of the fresh feed, recycle, and atomizing steam from their preheated states to the reactor temperature « Provide the zendothermic heat of cracking • Increase the temperature of the combustion air from the blower discharge temperature to the regenerator flue gas temperature • Make up for heat losses from the reactor and regenerator to the surroundings • Provide for miscellaneous heat sinks, such as stripping steam and catalyst cooling
A heat balance can be performed around the reactor, around the tripper-regenerator, and as an overall heat balance around the reactoregenerator. The stripper-regenerator heat balance can be used to alculate the catalyst circulation rate and the catalyst-to-oil ratio.
Heat Balance Around Stripper-Regenerator
If a reliable spent catalyst temperature is not available, the stripper s included in the heat balance envelope (II) as shown in Figure 5-4, The combustion of coke in the regenerator satisfies the following eat requirements: « Heat to raise air from the blower discharge temperature to the regenerator dense phase temperature • Heat to desorb the coke from the spent catalyst • Heat to raise the temperature of the stripping steam to the reactor temperature • Heat to raise the coke on the catalyst from the reactor temperature to the regenerator dense phase temperature • Heat to raise the coke products from the regenerator dense temperature to flue gas temperature • Heat to compensate for regenerator heat losses • Heat to raise the spent catalyst from the reactor temperature to the regenerator dense phase temperature
Using the operating data from the case study, Example 5-5 shows eat balance calculations around the stripper-regenerator. The results re used to determine the catalyst circulation rate and the delta coke. Delta coke is the difference between coke on the spent catalyst and oke on the regenerated catalyst.
Unit Monitoring and Control
161
Example 5-5 Stripper-Regenerator Heat Balance Calculations
I. Heat generated in the regenerator: C to CO2 = 24,288 Ib/hr x 14,087 Btu/lb = 342 x 106 Btu/hr H2 to H2O = 2,695 Ib/hr x 51,571 Btu/lb = 139 x 106 Btu/hr S to SO2 = 225 Ib/hr x 3,983 Btu/lb = 0.9 x 106 Btu/hr Total heat released in the regenerator: 342 + 139 + 0.9 = 482 x 106 Btu/hr II. Required heat to increase air temperature from blower discharge to the regenerator dense phase temperature: From Figure 5-5, enthalpies of air at 374°F and at 1,309°F are 90 Btu/lb and 355 Btu/lb. Therefore, the required heat is = 407,493 Ib/hr x (355 - 90) Btu/lb = 108.0 x 106 Btu/hr
III. Energy to desorb coke from the spent catalyst: Desorption of coke = 27,208 Ib/hr x 1,450 Btu/lb = 39.5 x 106 Btu/hr IV. Energy to heat the stripping steam: Enthalpy of 50 psig-saturated steam = 1,179 Btu/lb Enthalpy of 50 psig at 972°F =1,519 Btu/lb Change of enthalpy = 13,000 Ib/hr x (1,519 - 1,179) Btu/lb = 4.4 x 106 Btu/hr V. Energy to heat the coke on the spent catalyst: 27,231 Ibs/hr x 0.4 Btu/lb-°F x (1,309 - 972)°F = 3.7 x 106 Btu/hr
VI. Energy to heat the flue gas from regenerator dense phase to regenerator flue gas temperature: From Figure 5-5, enthalpy of flue gas at 1,309°F = 365 Btu/lb and at 1,330°F = 370 Btu/lb. The required heat is therefore = 433,445 Ib/hr x (370 - 355)°F = 2.6 x 106 Btu/hr
VII, Heat loss to surroundings: Assume heat loss from the stripper-regenerator (due to radiation and convection) is 4% of total heat of combustion, i.e., 0.04 x 482.4 MM Btu/hr = 19.3 x 106 Btu/hr
62
Fluid Catalytic Cracking Handbook
VIII. Energy required to heat the spent catalyst from its reactor to the regenerator temperature = 481.9 - 108.0 - 39.5 - 4.4 - 3.7 - 2.6 - 19.3 = 304.4 x 106 Btti/hr IX. Calculation of catalyst circulation ^ , „. . . Catalyst Circulation =
304.4 xl0 6 Btu/hr (0.285 Btu/°F-lb) x (1,309 - 972)°F
= 3.169 x 106 Ibs/hr = 26.4 short tons/min. Where: 0.285 is the catalyst heat capacity (see Figure 5-6) Cat/oil ratio = 3.169 x 106/658,914 = 4.8 .„ , Coke Yield, wt% 4.2 A 0_ „ ACoke = = — = 0.87 wt% cat/oil ratio 4,8
Reactor Heat Balance
The hot regenerated catalyst supplies the bulk of the heat required o vaporize the liquid feed (and any recycle) to provide the overall ndothermic heat of cracking, and to raise the temperature of disperion steam and inert gases to the reactor temperature. Heat In
Heat Out
Fresh Feed Recycle Air Steam
Reactor Vapors Flue Gas Losses
The calculation of heat balance around the reactor is illustrated in Example 5-6. As shown, the unknown is the heat of reaction. It is alculated as the net heat from the heat balance divided by the feed ow in weight units. This approach to determining the heat of reaction s acceptable for unit monitoring. However, in designing a new cat racker, a correlation is needed to calculate the heat of reaction. The eat of reaction is needed to specify other operating parameters, such
W
I
f
«M#
0J
E
re
to
of the FCC
20
30
40
as a function of the
Alumina Content, Wt.%
50
60
content,
Unit Monitoring and Control
165
as preheat temperature. Depending on conversion level, catalyst type, and feed quality, the heat of reaction can vary from 120 Btu/lb to 220 Btu/lb. In the unit, the heat of reaction is a useful tool. It is an indirect indication of heat balance accuracy. Trending the heat of reaction on a regular basis provides insight into reactions occurring in the riser and the effects of feedstock and catalyst changes. Example 5-6 Reactor Heat Balance
I. Heat into the reactor 1. Heat with regenerator catalyst = 3.169 x 106 Ib/hr x 0.285 Btu/lb-°F x 1,309°F = 1,182.4 x 106 Btu/hr = 1,182.4 x 106 Btu/hr 2. Heat with the fresh feed: At a feed temperature of 594°F, °API gravity = 25.2 and K factor = 12.08, the feed liquid enthalpy is 405 Btu/lb (see Figure 5-7), therefore, heat content of the feed is = 658,914 Ib/hr x 405 Btu/lb = 266.9 x 106 Btu/hr. 3. Heat with atomizing steam: From steam tables, enthalpy of 150 Ib saturated steam = 1,176 Btu/lb, therefore, heat with steam = 10,000 Ib/hr x 1,176 Btu/lb = 11.8 x 106 Btu/hr. 4 Heat of adsorption: The adsorption of coke on the catalyst is an exothermic process; the heat associated with this adsorption is assumed to be the same as desorption of coke in the regenerator (i.e., 35.3 x 106 Btu/hr). Total heat in = 1,182.4 + 266.9 + 11.8 + 35.3 = 1,496.4 x 106 Btu/hr. II. Heat out of the reactor 1, Heat with spent catalyst = 3,169 x 106 Ib/hr x 0.285 Btu/lb-°F x 972°F = 878 x 106 Btu/hr. 2, Heat required to vaporize feed: From Figure 5-8, enthalpy reactor vapors = 778 Btu/lb, therefore, heat content of the vaporized products = 658,814 Ib/hr x 778 Btu/lb = 512.6 x 106 Btu/hr. 3. Heat content of steam: Enthalpy of steam @ 972°F = 1,519 Btu/lb, therefore, heat content of steam = 10,000 Ib/hr x 1,519 Btu/lb = 15.2 x 106 Btu/hr. 4. Heat loss to surroundings: Assume heat loss due to radiant and convection to be 2% of heat with the regenerated catalyst (i.e., 0.02 x 304.4 = 6.1 x 106 Btu/hr)
61
Fluid Catalytic Cracking Handbook
I. Calculation of heat of reaction Total heat out = total heat in Total heat out = 878 x 106 + 512.6 x 106 + 15.2 x 106 + 6.1 x 1C)6 + overall heat of reaction = Total heat in = 1,499.6 x 106Btu/hr Overall endothermic heat of reaction = 84.5 x 106 Btu/hr or —» 128.2 Btu/lb of feed.
Analysis of Results
Once the material and heat balances are complete, a report must be written. It will first present the data. It will then discuss factors ffecting product quality and any abnormal results. It will then discuss he key findings and recommendations to improve unit operation. In the previous examples, the feed characterizing correlations in Chapter 2 are used to determine composition of the feedstock. The esults show that the feedstock is predominantly paraffinic (i.e., 61.6% araffins, 19.9% naphthenes, and 18.5% aromatics). Paraffinic feedocks normally yield the most gasoline with the least octane. This onfirms the relatively high FCC gasoline yield and low octane bserved in the test run. This is the kind of information that should e included in the report. Of course, the effects of other factors, such s catalyst and operating parameters, will also affect the yield structure nd will be discussed. The coke calculation showed the hydrogen content to be 9.9 wt%. As discussed in Chapter 1, every effort should be made to minimize he hydrogen content of the coke entering the regenerator. The hydroen content of a well-stripped catalyst is in the range of 5 wt% to wt%. A 9.9 wt% hydrogen in coke indicates either poor stripper peration and/or erroneous flue gas analysis.
RESSURE BALANCE
Pressure balance deals with the hydraulics of catalyst circulation in he reactor/regenerator circuit. The pressure balance starts with the atic pressures and differential pressures that are measured. The arious pressure increases and decreases in the circuit are then calulated. The object is to:
5 I
920
980
-*-K«11
960
1000
-*-K-t2
Deg. F
1040
-*-K»13
1020
1060
Figyre 5-8. Hydrocarbon vapor enthalpies at various Watson K factors.
940
1080
Unit Monitoring and Control
« * * *
169
Maximize catalyst circulation Ensure steady circulation Maximize the available pressure drop at the slide valves Minimize the loads on the blower and the wet gas compressor
A clear understanding of the pressure balance is extremely important n "squeezing" the most out of a unit. Incremental capacity can come rom increased catalyst circulation or from altering the differential pressure between the reactor-regenerator to "free up" the wet gas compressor or air blower loads. One must know how to manipulate he pressure balance to identify the "true" constraints of the unit. Using the drawing(s) of the reactor-regenerator, the unit engineer must be able to go through the pressure balance and determine whether t makes sense. He or she needs to calculate and estimate pressures, densities, pressure buildup in the standpipes, etc. The potential for mprovements can be substantial.
Basic Fluidization Principals
A fluidized catalyst behaves like a liquid. Catalyst flow occurs in he direction of a lower pressure. The difference in pressure between any two points in a bed is equal to the static head of the bed between hese points, multiplied by the fluidized catalyst density, but only if he catalyst is fluidized. FCC catalyst can be made to flow like a liquid, but only if the pressure force is transmitted through the catalyst particles and not the vessel wall. The catalyst must remain in a fluidized state as it makes a loop through the circuit. To illustrate the application of the above principals, the role of each major component of the circuit is discussed in the following sections, ollowed by an actual case study. As a reference, Appendix 8 contains luidization terms and definitions commonly used in the FCC.
Major Components of the Reactor-Regenerator Circuit
The major components of the reactor-regenerator circuit that either produce or consume pressure are as follows: * Regenerator catalyst hopper * Regenerated catalyst standpipe
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• • • • •
Fluid Catalytic Cracking Handbook
Regenerated catalyst slide (or plug) valve Riser Reactor-stripper Spent catalyst standpipe Spent catalyst slide (or plug) valve
egenerator Catalyst Hopper
In some FCC units, the regenerated catalyst flows through a hopper rior to entering the standpipe. The hopper is usually internal to the egenerator and often of an inverted cone design. It provides sufficient me for the regenerated catalyst to be deaerated before entering the tandpipe. This causes the catalyst entering the standpipe to have maximum flowing density. The higher the density, the greater the ressure buildup in the standpipe. In some FCC designs, the regenerated atalyst hopper is external with fluffing aeration to control the catalyst ensity entering the standpipe.
egenerated Catalyst Standpipe
The standpipe's height provides the driving force for transferring he catalyst from the regenerator to the reactor. The elevation differnce between the standpipe entrance and the slide valve is the source f this pressure buildup. For example, if the height difference is 30 eet (9.2 meters) and the catalyst density is 40 lb/ft3 (641 kg/m3), the ressure buildup is: 40 1h 1 ft2 Pressure Gain = 30 ft x -^3 x , = 8.3 psi (57Fkp) ft 144 in2
The key to obtaining maximum pressure gain is to keep the catalyst uidized over the length of the standpipe. Longer standpipes will equire external aeration. This compensates for compression of the ntrained gas as it travels down the standpipe. Aeration should be dded evenly along the length of the standpipe. In shorter standpipes ufficient flue gas is often carried down with the regenerated catalyst o keep it fluidized and supplemental aeration is unnecessary. Overeration leads to unstable catalyst flow and must be avoided.
Unit Monitoring and Control
171
Aside from proper aeration, the flowing catalyst must contain sufficient 0-40 micron fines to avoid defluidization.
Regenerated Catalyst Slide Valve
The purpose of the regenerated catalyst slide valve is threefold: to regulate the flow of regenerated catalyst to the riser, to maintain pressure head in the standpipe, and to protect the regenerator from a flow reversal. Associated with this control and protection is usually a I psi to 8 psi (7 Kp to 55 Kp) pressure drop across the valve.
Riser
The hot-regenerated catalyst is transported up the riser and into the reactor-stripper. The driving force to carry this mixture of catalyst and vapors comes from a higher pressure at the base of the riser and the ow density of the catalyst/vapor mix. The large density difference between the fluidized catalyst on the regenerator side (approximately 40 b/ft3) and the mixture of cracked hydrocarbon vapors and catalyst on the riser side (approximately 1 lb/ft3) drives the system. As for the pressure balance, this transport of catalyst results in a pressure drop in a range of 5 psi to 9 psi (35 Kp to 62 Kp). This drop is due to static head and, to a esser extent, friction and acceleration of the fluid. In an existing riser, operating changes, such as higher catalyst circulation or lower vapor velocity, can affect the density of reaction mixture and increase pressure drop. This will affect the slide valve differential and percent opening.
Reactor-Stripper The catalyst bed in the reactor-stripper is important for three reasons: « to provide enough residence time for proper stripping of the entrained hydrocarbon vapors prior to entering the regenerator; • to provide adequate static head for flow of the spent catalyst to the regenerator; and • to provide sufficient backpressure to prevent reversal of hot flue gas into the reactor system.
Assuming a stripper with a 20-ft bed level and a catalyst density of 40 lb/ft 3 , the static pressure is: 3 , 40 lbs/ft .. . 20n ft x f— r 2 2 = 5.5 psi
144 in /ft
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Fluid Catalytic Cracking Handbook
pent Catalyst Standpipe
From the bottom of the stripper, the spent catalyst flows into the pent catalyst standpipe. Sometimes the catalyst is partially defluidized n the stripper cone. To counter this, "dry" steam is usually added hrough a distributor) to fluidize the catalyst prior to its entering the tandpipe. The loss of fluidization in the stripper cone can cause a uildup of dense phase catalyst along the cone walls. This buildup can estrict catalyst flow into the standpipe, causing erratic flow and educing pressure buildup in the standpipe. Like the regenerated catalyst standpipe, the spent catalyst standpipe may require supplemental aeration to obtain optimum flow charcteristics. "Dry" stearn is the usual aeration medium.
pent Catalyst Slide or Plug Valve
The spent catalyst slide valve is located at the base of the standpipe. controls the stripper bed level and regulates the flow of spent atalyst into the regenerator. As with the regenerated catalyst slide alve, the catalyst level in the stripper generates pressure as long as is fluidized. The pressure differential across the slide valve will be t the expense of consuming a pressure differential in the range of psi to 6 psi (20 kp to 40 kp). In earlier Model II and Model III FCC units, spent catalyst was ansported into the regenerator using 50% to 100% of combustion ir. This spent cat riser was designed for a minimum air velocity of 0 ft/sec (9.1 m/sec).
Case Study
A survey of the reactor-regenerator circuit of a 50,000 bpd (331 m3/hr) cat cracker produced these results:
eactor dilute phase (dome) pressure eactor catalyst dilute phase bed level eactor-stripper catalyst bed level eactor-stripper catalyst density pent, catalyst standpipe elevation ressure above the spent catalyst slide valve pent catalyst slide valve AP (@ 55% opening)
= = = = = = =
19.0 psig/131 Kp 25.0 ft/7.6 m 18.0 ft/5.5 m 40 Ib/ft3/640 kg/m3 14.4 ft/4.4 m 26.1 psig/180 Kp 4.0 psi/27.6 Kp
Unit Monitoring and Control
Regenerator dilute phase catalyst level Regenerator dense phase catalyst bed level Catalyst density in the regenerator dense phase Regenerated catalyst standpipe elevation Pressure above the regenerated catalyst slide valve Regenerated catalyst slide valve AP (@ 30% opening) Reactor-regenerator pressure AP
173
= 27.0 ft/8.2 m = 15.0 ft/4.6 m = 25 Ib/ft3/400 kg/m3 = 30.0 ft/9.1 m = 30.5 psig/210.3 Kp = 5.5 psi/37.9 Kp = 3 . 0 psi/20.7 Kp
Also, see Figure 5-9 for a graphical representation of the preliminary esults. Starting with the reactor dilute pressure as the working point, the pressure head corresponding to 25 feet (7.6 m) of dilute catalyst ines is: (25 ft) x (0.6 lb/ft3) x (1 ft2/!44 in 2 ) = 0.1 psig (0.7 Kp)
Therefore, the pressure at the top of the stripper bed is: 19.0 + 0.1 = 19.1 psig (131.7 Kp)
The static-pressure head in the stripper is: (18 ft) x (40 lb/ft3) x (1 ft/144 in 2 ) = 5.0 psig (34.5 Kp)
The pressure above the spent catalyst standpipe is: 19.1 + 5.0 = 24.1 psig (166.2 Kp)
The pressure buildup in the spent catalyst standpipe is:
26.1 - 24.1 =2 psi (13.8 K p )
The pressure below the spent catalyst slide valve is: 26.1 -4.0 = 22.1 psig (152 Kp)
The pressure head corresponding to 28 feet (8.5 m) of dilute catalyst ines in the regenerator is: (28 ft) x (1 lb/ft3) x (1 ft2/144 in2) = 0.2 psig (1.4 Kp)
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Fluid Catalytic Cracking Handbook
Rx Vapor
Reactor
Psi diff.
Oil Feed Figure 5-9. Preliminary pressure balance survey.
Unit Monitoring and Control
175
The pressure in the regenerator dome is: 22. J - 0,2 = 21.9 psig (151.0 KP)
The static pressure head in the regenerator is: (18 ft) x (25 lb/ft3) x (1 ft2/!44 in2) = 3.1 psig (21.4 Kp)
The pressure above the regenerated catalyst standpipe is: 22.1 + 3.1 = 25.2 psig (173.7 Kp)
The pressure buildup in the regenerated catalyst standpipe is: 30.5 - 25.2 = 5.3 psi (36.5 Kp)
The pressure below the regenerated catalyst slide valve is: 30.5 - 5.5 = 25 psig (172.4 Kp)
The pressure drop in the vertical riser is: 25 - 19 = 6 psi (41.4 Kp)
The catalyst density in the spent catalyst standpipe is: (2.0 lb/in2) x (144 in2/ft2)/(14.4 ft) = 20 lb/ft3 = 320 kg/m3
The catalyst density in the regenerated catalyst standpipe is: (5.3 lb/in 2 ) x (144 in2/ft2)/(30 ft) = 25.4 lb/ft3 = 407 kg/m3
Figure 5-10 shows the results of the above pressure balance survey.
Analysis of the Findings
The pressure balance survey indicates that neither the spent nor the egenerated catalyst standpipe is generating "optimum" pressure head. This is evidenced by the low catalyst densities of 20 lb/ft3 (320 kg/m3) nd 25.4 lb/ft3 (407 kg/m3), respectively. As indicated in Chapter 8, everal factors can cause low pressure, including "under" or "over"
76
Fluid Catalytic Cracking Handbook
Rx Vapor Reactor
Psi diff.
Oil Feed
igure 5-10. Pressure balance survey with calculated standpipe densities.
Unit Monitoring and Control
177
aeration of the standpipes. In a well-fluidized standpipe, the expected catalyst density is in the range of 35 - 45 lb/ft3 (561 kg/m 3 to 721 kg/m3). If the catalyst density in the spent catalyst standpipe was 40 lb/ft'* (640 kg/m3) instead of 20 lb/ft3 (320 kg/m3), the pressure buildup would have been 4.0 psi instead of 2.0 psi. The extra 2 psi (13.8 Kp) can be used to circulate more catalyst or to lower the reactor pressure. In the regenerated catalyst standpipe, a 40 lb/ft3 (640 kg/m3) catalyst ^ ' 3 density versus a 25.4 lb/ft (407 kg/m") density produces 3 psi (20,7 Kp) more pressure head, again allowing an increase in circulation or a reduction in the regenerator pressure (gaining more combustion air).
Process control instrumentation controls the FCC unit in a safe, monitored mode with limited operator intervention. Two levels of process control are used: • Basic supervisory control • Advanced process control (APC)
Basic Supervisory Control
The primary controls in the reactor-regenerator section are flow, emperature, pressure, and catalyst level. The flow controllers are often used to set desired flows for the fresh eed, stripping steam, and dispersion steam. Each flow controller usually has three modes of control: manual, auto, and cascade. In manual mode, the operator manually opens or closes a valve to the desired percent opening. In auto mode, the operator enters the desired low rate as a set-point. In cascade mode, the controller set-point is an input from another controller. The reactor temperature is controlled by a temperature controller that egulates the regenerated catalyst slide valve. The regenerator temperaure is not automatically controlled but depends on its mode of operation. In partial combustion, the regenerator temperature is conrolled by adjusting the flow of combustion air to the regenerator. In ull burn, the regenerator temperature is a function of operating conditions such as reactor temperature and slurry recycle.
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Fluid Catalytic Cracking Handbook
The reactor pressure is not directly controlled; instead, it floats on he main column overhead receiver. A pressure controller on the verhead receiver controls the wet gas compressor and indirectly ontrols the reactor pressure. The regenerator pressure is often conolled directly by regulating the flue gas slide or butterfly valve. In ome cases, the flue gas slide or butterfly valve is used to control the ifferential pressure between the regenerator and reactor. The reactor or stripper catalyst level controller is controlled with level controller that regulates the movement of the spent catalyst lide valve. The regenerator level is manually controlled to maintain atalyst inventory.
egenerated and Spent Catalyst Slide Valve ow Differential Pressure Override
Normally, the reactor temperature and the stripper level controllers egulate the movement of the regenerated and spent catalyst slide alves. The algorithm of these controllers can drive the valves either ully open or fully closed if the controller set-point is unobtainable. is extremely important that a positive and stable pressure differential e maintained across both the regenerated and spent catalyst slide alves. For safety, a low differential pressure controller overrides the emperature/level controllers should these valves open too much. The hutdown is usually set at 2 psi (14 Kp). The direction of catalyst flow must always be from the regenerator o the reactor and from the reactor back to the regenerator. A negative ifferential pressure across the regenerated catalyst slide valve can llow hydrocarbons to back-flow into the regenerator. This is called flow reversal and can result in an uncontrolled afterburn and ossible equipment damage. A negative pressure differential across he spent catalyst slide valve can allow air to back-flow from the egenerator into the reactor with equally disastrous consequences. To protect the reactor and the regenerator against a flow reversal, ressure differential controllers are used to monitor and control the ifferential pressures across the slide valves. If the differential pressure alls below a minimum set-point, the pressure differential controller PDIC) overrides the process controller and closes the valve. Only fter the PDIC is satisfied will the control of the slide valve return to he process.
Unit Monitoring and Control
179
To maximize the unit's profit, one must operate the unit simulaneously against as many constraints as possible. Examples of these onstraints are limits on the air blower, the wet gas compressor, eactor/regenerator temperatures, slide valve differentials, etc. The onventional regulatory controllers work only one loop at a time and hey do not talk to one another. A skilled operator can "push" the unit gainst more than one constraint at a time, but the constraints change ften. To operate closer to multiple constraints, a number of refiners ave installed an advanced process control (APC) package either within their DCS or in a host computer. The primary advantages of an APC are: * It provides more precise control of the operating variables against the unit's constraints and, therefore, obtains incremental throughput or cracking severity. * It is able to respond quickly to ambient disturbances, such as cold fronts or rainstorms. It can run a day/night operation, taking advantage of the cooler temperatures at night. * It pushes against two or more constraints rather than one single constraint. It can maximize the air blower and wet gas compressor capacities.
As mentioned above, there are two options for installing an APC. One option is to install an APC within the DCS framework, and the ther is to install a multivariable modeling/control package in a host omputer. Each has advantages and disadvantages, as indicated below,
dvantages of Multivariable Modeling and Control
The multivariable modeling/control package is able to hold more ghtly against constraints and recover more quickly from disturbances. his results in an incremental capacity used to justify multivariable ontrol. An extensive test run is necessary to measure the response f unit variables. In APC on DCS framework, the control structure must be designed, onfigured, and programmed for each specific unit. Modifying the ogic can be an agonizing process. Wiring may be necessary. It is ifficult to both document the programming and to test.
80
Fluid Catalytic Cracking Handbook
With a host computer framework, the control package is all in the oftware. Changing the program can still be agonizing, but the program an be tested off-line. There is more flexibility in the computer system, which can be used for many other purposes, including on-line heat nd weight balances.
Disadvantages of Multivariable Modeling and Control
A multivariable model is like a "black box." The constraints go in nd the signals come out. Operators do not trust a system that takes he unit away from them. Successful installations require good training nd continual communication. The operators must know the interconections in the system. The model may need expensive work if changes are made during a urnaround. If the feed gets outside the range the unit was modeled or, results can be at best unpredictable. An upset can happen for which he system was not programmed. The DCS-based APC is installed in a modular form, meaning operators an understand what the controlled variable is tied to more easily. The host computer-based system may have its own problems, includng computer-to-computer data links. In any APC, the operators must be educated and brought into it efore they can use it. The control must be properly designed, meaning he model must be configured and properly "tuned." The operators hould be involved early and all of them should be consulted since ll four shifts may be running the unit differently.
SUMMARY
The only proper method to evaluate the performance of a cat cracker s by conducting a material and heat balance. One balance will tell where the unit is; a series of daily or weekly balances will tell where he unit is going. The heat and weight balance can be used to evaluate revious changes or predict the result of future changes. As discussed n the next chapter, material and heat balances are the foundation for etermining the effects of operating variables. The material balance test run provides a standard and consistent pproach for daily monitoring. It allows for accurate analysis of yields nd trending of unit performance. The reactor effluent can be deter-
Unit Monitoring and Control
181
mined by direct sampling of the reactor overhead line or by conducting unit test run, The heat balance exercise provides a tool for in-depth analysis of he unit operation. Heat balance surveys determine catalyst circulation ate, delta coke, and heat of reaction. The procedures described in this hapter can be easily programmed into a spreadsheet program to alculate the balances on a routine basis. The pressure balance provides an insight into the hydraulics of atalyst circulation. Performing pressure balance surveys will help the nit engineer identify "pinch points." It will also balance two common onstraints: the air blower and the wet gas compressor. Finally, process control systems allow the unit to operate smoothly nd safely. At the next level, an APC package (whether within the DCS ramework or as a host-based multivariable control system) provides more precise control of operating variables against the unit's contraints. It will gain incremental throughput or cracking severity.
REFERENCES
. Davison Div., W.R. Grace & Co., "Cat Cracker Heat and Material Balance Calculations," Grace Davison Catalagmm, No. 59, 1980. . Hsieh, C. R. and English, A. Ar., "Two Sampling Techniques Accurately Evaluate Fluid-Cat-Cracking Products," Oil & Gas Journal, June 23, 1986, pp. 38-43.
CHAPTER 6
Products and Economics
The previous chapters explained the operation of a cat cracker. However, the purpose of the FCC unit is to maximize profitability for he refinery. The cat cracker provides the conversion capacity that very refinery needs to survive. All crudes have heavy gas oils and uel oil; unfortunately, the market for these products has disappeared. FCC economics makes the refinery a viable entity. Over the years, efineries without cat crackers have been shut down because they were ot profitable. Understanding the economics of the unit is as important as undertanding the heat and pressure balance. The dynamics of FCC economics hanges daily and seasonally. It is dependent on market conditions and he availability of feedstocks. The 1990 Clean Air Act Amendment CAAA) has imposed greater restrictions on quality standards for asoline and diesel. The FCC is the major contributor to the gasoline nd diesel pool and is significantly affected by these new regulations. This chapter discusses the factors affecting yields and qualities of FCC product streams. The section on FCC economics describes several ptions that can be used to maximize FCC performance and the efinery's profit margin.
FCC PRODUCTS
The cat cracker converts less valuable gas oils to more valuable roducts. A major objective of most FCC units is to maximize the onversion of gas oil to gasoline and LPG. The products from the cat racker are: • Dry Gas • LPG • Gasoline 182
Products and Economics
• « • •
183
LCO HCO Decanted Oil Coke
Dry Gas
The gas (C2 and lighter) leaving the sponge oil absorber is commonly referred to as dry gas. Its main components are hydrogen, methane, ethane, ethylene, and hydrogen sulfide (H2S). Once the gas s amine-treated for removal of H2S and other acid gases, it is blended nto the refinery fuel gas system. Depending on the volume percent f hydrogen in the dry gas, some refiners recover hydrogen using rocesses such as cryogenics, pressure-swing absorption, or membrane eparation. The recovered hydrogen is often used in hydrotreating. Dry gas is an undesirable by-product of the FCC unit; excessive ields load up the wet gas compressor (WGC) and are often a contraint. The dry gas yield is primarily due to thermal cracking, metals n the feed, and nonselective catalytic cracking. The main factors that ontribute to the increase of dry gas are: • Increase in the concentration of metals (nickel, vanadium, etc.) on the catalyst • Increase in reactor or regenerator temperatures • Increase in the residence time of hydrocarbon vapors in the reactor • Decrease in the performance of the feed nozzles 8 Increase in the aromaticity of the feed
The overhead stream from the debutanizer or stabilizer is a mix of C3's and C4's, usually referred to as LPG (liquefied petroleum gas). t is rich in olefins, propylene, and butylene. These light olefins play n important role in the manufacture of reformulated gasoline (RFC), Depending on the refinery's configuration, the cat cracker's LPG is sed in the following areas: * Chemical sale, where the LPG is separated into C3's and C4's. The C3's are sold as refinery or chemical grade propylene. The C4 olefins are polymerized or alkylated.
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Fluid Catalytic Cracking Handbook
* Direct blending, where the C4's are blended into the refinery's gasoline pool to regulate vapor pressure and to enhance the octane number. However, new gasoline regulations require reduction of the vapor pressure, thus displacing a large volume of C4's for alternative uses. * Alkylation, where the olefins are reacted with isobutane to make a very desirable gasoline blending stock. A Iky late is an attractive blending component because it has no aromatics or sulfur, low vapor pressure, low end point, and high research and motor octane ratings. * MTBE, where isobutylene is reacted with methanol to produce an oxygenate gasoline additive called methyl tertiary butyl ether (MTBE). MTBE is added to gasoline to meet the minimum oxygen requirement for "reformulated" gasoline.
The LPG yield and its olefinicity can be increased by: * Changing to a catalyst, which minimizes "hydrogen transfer" reactions » Increasing the conversion * Decreasing residence time, particularly the amount of time product vapors spend in the reactor housing before entering the main column » Adding ZSM-5 catalyst additive
An FCC catalyst containing zeolite with a low hydrogen transfer ate reduces resaturation of the olefins in the riser. As stated in Chapter , primary cracking products in the riser are highly olefinic. Most of hese olefins are in the gasoline boiling range; the rest appear in the LPG and LCO boiling range. The LPG olefins do not crack further, but they can become saturated y hydrogen transfer. The gasoline and LCO range olefins can be racked again to form gasoline range olefins and LPG olefins. The lefins in the gasoline and LCO range can also cyclize to form ycloparaffins. The cycloparaffins can react through H2 transfer with lefins in the LPG and gasoline to produce aromatics and paraffins. Therefore, a catalyst that inhibits hydrogen transfer reactions will ncrease olefinicity of the LPG, The conversion increase is accomplished by manipulating the folowing operating conditions: * Increasing the reactor temperature. Increasing the reactor temperature beyond the peak gasoline yield results in overcracking
Products and Economics
185
of the gasoline and LCO fractions. The rate of production and olefinicity of the LPG will increase, Increasing feed/catalyst mix zone temperature. Conversion and LPG yield can be increased by injecting a portion of the feed, or naphtha, at an intermediate point in the riser (see Figure 6-1). Splitting or segregation of the feed results in a high-mix zone temperature, producing more LPG and more olefins. This practice
30% of Feed
70% of Feed
Figure 6-1.
A typical feed segregation scheme.
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Fluid Catalytic Cracking Handbook
is particularly useful where the reactor temperature is already maximized due to a metallurgy constraint. » Increasing catalyst to oil ratio. The catalyst to oil ratio can be increased through several knobs including: reducing the FCC feed preheat temperature, optimizing the stripping and dispersion steam rate, and using a catalyst that deposits less coke on the catalyst,
Reduction of the catalyst/hydrocarbon time in the riser, coupled with he elimination of post-riser cracking, reduces the saturation of the already produced olefins" and allows the refiner to increase the eaction severity. The actions enhance the olefin yields and still operate within the wet gas compressor constraints. Elimination of post-riser esidence time (direct connection of the reactor cyclones to the riser) r reducing the temperature in the dilute phase virtually eliminates ndesired thermal and nonselective cracking. This reduces dry gas and iolefin yields. Adding ZSM-5 catalyst additive is another process available to the efiner to boost production of light olefins. ZSM-5 at a typical conentration of 0.5 to 3.0 wt% is used in a number of FCC units to ncrease the gasoline octane and light olefins. As part of the cracking f low octane components in the gasoline, ZSM-5 also makes C3, C4, nd C5 olefins (see Figure 6-2). Paraffinic feedstocks respond the most o ZSM-5 catalyst additive.
Gasoline
FCC gasoline has always been the most valuable product of a cat racker unit. FCC gasoline accounts for about 35 vol% of the total U.S. gasoline pool. Historically, the FCC has been run for maximum asoline yield with the highest octane.
Gasoline Yield For a given feedstock, gasoline yield can be increased by: • Increasing catalyst-to-oil ratio by decreasing the feed preheat temperature • Increasing catalyst activity by increasing fresh catalyst addition or fresh catalyst activity • Increasing gasoline end point by reducing the main column top pumparound rate
Products and Economics
0
5
10
15
ZSM-5 wt% in Catalyst Inventory Figure 6-2. The effect of ZSM-5 on light-ends yield [5].
* Increasing reactor temperature (if the increase does not over-crack the already produced gasoline)
Gasoline Quality
The Clean Air Act Amendment (CAAA), passed in November 1990, as set new quality standards for U.S. gasoline. A complete discussion f the new gasoline formulation requirements can be found in Chaper 10.
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Fluid Catalytic Cracking Handbook
The key components affecting FCC gasoline quality are octane, enzene, and sulfur and are discussed in the following sections.
Octane. An octane number is a quantitative measure of a fuel mixture's resistance to "knocking." The octane number of a particular ample is measured against a standard blend of n-heptane, which has ero octane, and iso-octane, which has 100 octane. The percent of isoctane that produces the same "knock" intensity as the sample is eported as the octane number. Two octane numbers are routinely used to simulate engine perormance: the research octane number (RON) simulates gasoline erformance under low severity (@600 rpm and 120°F (49°C) air emperature), whereas the motor octane number (MON) reflects more evere conditions (@900 rpm and 300°F (149°C) air temperature). At he pump, road octane, which is the average of RON and MON, s reported. Factors affecting gasoline octane are: A. Operating Conditions 1. Reactor Temperature. As a rule, an increase of 18°F (10°C) in the reactor temperature increases the RON by 1.0 and MON by 0.4. However, the MON contribution comes from the aromatic content of the heavy end. Therefore, at high severity, the MON response to the reactor temperature can be greater than 0.4 number per 18°F. 2. Gasoline End Point, The effect of gasoline end point on its octane number depends on the feedstock quality and severity of the operation. At low severity, lowering the end point of a paraffinic feedstock may not impact the octane number; however, reducing gasoline end point produced from a naphthenic or an aromatic feedstock will lower the octane. 3. Gasoline Reid Vapor Pressure (RVP). The RVP of the gasoline is controlled by adding C4's, which increase octane. As a rule, the RON and MON gain 0.3 and 0.2 numbers for a 1.5 psi (10.3 Kp) increase in RVP. B. Feed Quality 1. °API Gravity, The higher the °API gravity, the more paraffins in the feed and the lower the octane (Figure 6-3). 2. K Factor. The higher the K factor, the lower the octane.
Products and Economics
93
92 U
Z
o 91
90 20
22
24
26
Feed Gravity, "API
82
81 Q Z O
s
80
79 20
22
24
26
Feed Gravity, "API Figure 6-3. Feed gravity comparisons (MON and RON) [7].
189
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Fluid Catalytic Cracking Handbook
3, Aniline Point. Feeds with a higher aniline point are less aromatic and more paraffinic. The higher the aniline point, the lower the octane. 4. Sodium. Additive sodium reduces unit conversion and lowers octane (Figure 6-4). C. Catalyst 1. Rare Earth. Increasing the amount of rare earth oxide (REO) on the zeolite decreases the octane (Figure 6-5). 2. Unit Cell Size. Decreasing the unit cell size increases octane (Figure 6-6). 3. Matrix Activity. Increasing the catalyst matrix activity increases the octane. 4. Coke on the Regenerated Catalyst. Increasing the amount of coke on the regenerated catalyst lowers its activity and increases octane.
Benzene. Most of the benzene in the gasoline pool comes from eformate. Reformate, the high-octane blending component from a eformer unit, comprises about 30 vol% of the gasoline pool. Dependng on the reformer feedstock and severity, reformate contains 3 vol% o 5 vol% benzene. FCC gasoline contains 0.5 to 1.3 vol% benzene. Since it accounts or about 35 vol% of the gasoline pool, it is important to know what ffects the cat cracker gasoline benzene levels. The benzene content n the FCC gasoline can be reduced by: • Short contact time in the riser and in the reactor dilute phase • Lower cat-to-oil ratio and lower reactor temperature • A catalyst with less hydrogen transfer
Sulfur. The major source of sulfur in the gasoline pool comes from CC gasoline. Sulfur in FCC gasoline is a strong function of the feed ulfur content (Figure 6-7). Hydrotreating the FCC feedstock reduces ulfur in the feedstock and, consequently, in the gasoline (Figure 6-8). Other factors that can lower sulfur content are: • Lower gasoline end point (see Figure 6-9) • Lower reactor temperature (see Figure 6-10) • Increased matrix activity of the catalyst (text continued on page 195)
Products and Economics RONC vs. SODIUM COMMERICAL DATA
0.40
0.60
EQUILIBRIUM CAT. SODIUM, WT.%
< 80.5 §80.0 -
79.5 79.0 78.5 _
78.0
0.20
0.40
0.80
EQUILIBRIUM CAT SODIUM, WT. %
Figure 6-4.
Effect of sodium on gasoline octane [8J.
191
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Fluid Catalytic Cracking Handbook
84 PILOT PLANT DATA
83
C5-265T/C5-129"C 265-430"F/129-221*C
82
2; 81 o S
80 79
—8
78 77 0 .0
1.0
2.0
3.0
4.0
REO, WT. % Figure 6-5. Effect of fresh REO on MON [9].
82
95
94
81
93
UJ
z 80 O O o: 79 O
91 au on
I 78
89 88
24.20
I
I
I
24.24
Figure 6-6.
I
24.28
I
I
24.32
UNIT CELL SIZE, A
I
I
24.36
I
77
24,20
24.24
24.28
I
24.32
i
i
24.36
UNIT CELL SIZE, A
Effects of unit cell size on research and motor octane [10].
Products and Economics
" £
0.1
1
0.03
0
High N VGO
0.01
O
5
o.oos
34% Recycle
0,001 0.05
0.1
0.2
0.5
1
FCCU Feed Sulfur, wt% Figure 6-7.
FCC gasoline sulfur yield [4].
2,000
Non-Hydrotreated 1.000
o. a.
c
U
Z
O
<
0
Hydrotreated
O O
FCCU FEEDSTOCK SULFUR (wt%) Figure 6-8.
Hydrotreating reduces FCC gasoline sulfur [4].
113
194
Fluid Catalytic Cracking Handbook
1,000
Guff Coast FCCU Feed. 0.62 w% S I
HT FCCU Feed, 0.68 w% S 350
450
FCC Gasoline End Point (°F) Figure 6-9.
FCC gasoline sulfur increases with end point [4],
400
Q. ^"350
Octane Catalyst
3 300
o Octane BBL Catalyst 250 475
500
525
FCC Reactor Isothermal Temperature (°C) Figure 6-10.
FCC gasoline sulfur increases with temperature [4j.
550
Products and Economics
195
ext continued from page 190)
• Increase in the catalyst activity and hydrogen transfer properties • Increase in catalyst-to-oil ratio (Figure 6-11) • Increase in the use of main column overhead reflux rate instead of top pumparound to control the top temperature
LCO
The emphasis on gasoline yield has sometimes overshadowed the mportance of other FCC products, particularly LCO. LCO is widely used as a blending stock in heating oil and diesel fuel. This is paricularly important during winter, when the price of light cycle oil can be higher than gasoline. Under these circumstances, many refiners adjust he FCC operation to increase LCO yield at the expense of gasoline.
LCO Yield
A refiner has several options to increase LCO yield. Since it is often desirable to maintain a maximum cracking severity while maximizing ight cycle oil yield, the simplest way to increase LCO yield is to educe the gasoline end point. Gasoline end point is usually reduced
450
OCTANE OCTANE BBL Feed Sulfur = 0,46 wt%
.400 (-
t
5 F
3 350
w
P
c 300
!
r 250
Figure 6-11,
4
5
6
7
Catalyst to Oil Ratio (W/W) Increased catalyst-to-oil ratio decreases gasoline sulfur [4],
j
96
Fluid Catalytic Cracking Handbook
y lowering the top temperature on the main column by increasing he top pumparound or the top reflux rate. The LCD distillation range is typically 430°F to 67()°F (221°C to 54°C) ASTM. Undercutting the gasoline end point drops the heavy end f the gasoline fraction to be withdrawn with LCO. This affects only the pparent conversion and does not cause changes in the flow rate of other roducts. Reducing the gasoline end point usually increases the octane ecause of the lower-octane components in the heavy end of gasoline. A better method of increasing LCO yield is through better fraconation upstream. The removal of the fraction under 650°F from the eed requires better stripping. The total refinery yield of diesel will ncrease when the light ends are fractionated from the feed (Table 6-1). Some of the catalytic routes to maximize LCO yield are: » * « •
Decrease in the reactor temperature Decrease in the catalyst-to-oil ratio Decrease in catalyst activity Increase in HCO recycle
CO Quality
The CAAA of 1990 set fuel standards for the new "over-the-road iesel fuel." It requires a maximum sulfur of 0.05 wt% (500 ppm) and Table 6-1 Effects of Feed Fractionation on Total Distillate Yield
nitial Boiling Point, °F/°C inal Boiling Point, °F/°C 35°F/224°C to 660°F/349°C Content, wt% onversion, wt% CO, wt% otential FCC LCO, wt% otal Potential Refinery Distillate
ource: Engelhard [6]
Feedstock "Raw" Gas Oil
"Fractionated" Gas Oil
435/224 1,080/582 8
660/349 1,080/582 0
75.9 15.4 15.4 15.4
75.9 14.0 (0.92 x 14.0) = 12 .9 (12.9 + 8.0) = 20.,9
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maximum aromatics of 35 vol%. The California Air Resources Board CARB) has imposed a tighter restriction on diesel aromatics of 10 ol% for larger refineries (over 50,000 bpd capacity) and 20 vol% for maller refineries. A minimum cetane number (CN) is also required. The aromatics content of a diesel fuel affects its cetane number; the ollowing section discusses the cetane number and the principal factors ffecting it.
Cetane. Like octane number, cetane number is a numerical indicaon of the ignition quality of a fuel. However, the two numbers work n reverse. A gasoline engine is spark-ignited and an important fuel uality is to prevent premature ignition during the compression stroke. A diesel engine is compression-ignited and it must ignite when comressed. Unfortunately, components that increase octane will decrease etane. For example, normal paraffinic hydrocarbons have a low octane umber but a very high cetane number. Aromatics have a high octane umber but a very low cetane number. The adjustments in the reactor ield mentioned above to improve LCO yield and quality will all lower asoline yield and quality. Cetane number is measured in a single-cylinder laboratory engine, ut cetane index (CI) is more commonly used. Cetane index is a alculated value and correlates adequately with the cetane number. Most refiners use the ASTM equation (Method D-976-80) to calculate he cetane index. The equation uses the 50% boiling point and °API ravity (Example 6-1). Example 6-1 Cetane Index Equation CI = 65.01 (log T)2 + [0.192(°API) x log T] + 0.16(°API)2 - 0.0001809 (T)2 - 420.34 = 65,01 (log 550)2+ [0.192(19)(/0g 550) 0.16(19)2 - 0.0001809 (550)2- 420.34 = 65.01(2.74)2 + [0.192(19)(2.74)J + 0.16(361) 0.0001809(302,500) - 420.34 = 488.2 + 10.0 + 5.8 - 54.7 - 420.34 CI = 28.9
98
Fluid Catalytic Cracking Handbook
Where:
CI = Cetane index T = Mid boiling temperature, °F API = °API gravity at 60°F
or Example:
T = 550°F (288°C) API = 19 CI = 28.9
ource: ASTM Standards, Method D-976-80
LCO is highly aromatic (50 wt% to 75 wt%) and has a low cetane ndex (20-30). The cetane number and sulfur content determine the mount of LCO that can be blended into the diesel or heating oil pool. Most (30-50 wt%) of the aromatics in the LCO are di- and triromatic molecules. Hydrotreating the LCO can increase its cetane umber. The degree of improvement depends on the severity of the ydrotreating. Mild hydrotreating (500 to 800 psig/3500 to 5500 Kp) an partially hydrogenate some of the di- and tri-aromatics and ncrease cetane by a 1 to 5 number. Severe hydrotreating conditions > 1,500 psig/10,300 Kp) can increase the cetane number above 40. Other conditions-that improve cetane are: » « * «
Undercutting the FCC gasoline Reducing the unit conversion Using an "octane" catalyst Processing paraffinic feedstock
HCO and Decant Oil
HCO is the sidecut stream from the main column that boils between CO and decanted oil (DO). HCO is often used as a pumparound ream to transfer heat to the fresh feed and/or to the debutanizer reboiler. HCO is recycled to extinction, withdrawn as a product and processed n a hydrocracker, or blended with the decant oil. DO is the heaviest product from a cat cracker. DO is also called slurry il, clarified oil, bottoms, and FCC residue. Depending on the refinery ocation and market availability, DO is typically blended into No. 6 fuel, old as a carbon black feedstock (CBFS), or even recycled to extinction.
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DO is the lowest priced product and the goal is to reduce its yield, The DO's yield largely depends on the quality of the feedstock and he conversion level. Naphthenic and aromatic feedstocks tend to yield more bottoms than paraffinic feedstocks. If the conversion is in the ow- to mid-70's, increasing catalyst-to-oil ratio or using a catalyst with an active matrix can reduce slurry yield. Raising conversion educes bottoms yield. If the conversion is in the 80's, little more can be done to reduce bottoms yields.
Decant Oil Quality
DO properties vary greatly depending on the feedstock quality and operating conditions. Selling the DO as CBFS yields a better return than selling it as utter stock. To meet the CBFS specification, DO must have a minimum Bureau of Mines Correlation Index (BMCI) of 120 and a low sh content (Table 6-2). Aromaticity and ash content are the two most mportant properties of CBFS. Table 6-2 Typical Carbon Black Feedstock Specifications
Property
Gravity, °API Asphaltenes, wt% Viscosity, SUS @ 210°F (98.9°C) ulfur, wt% Ash, wt% Sodium, ppm otassium, ppm lash, °F BMCI*
Specification
3.0, maximum 5.0, maximum 80, maximum 4.0, maximum 0.05, maximum 15, maximum 2, maximum 200, (93.3°C) minimum 120, minimum
Bureau of Mines Correlation Index (BMCI) BMCI = (87,552/T) + [473.7 x (141.5/131.6 + °AP1 Gravity)] - 456.8
Where: = Mid boiling point, °R
or Example:
T = 7!0°F (376.7°C) = 7JO°F + 460 = 1I70°R "API = 1.0 MCI = 23.9
00
Fluid Catalytic Cracking Handbook
BMC! is a function of gravity and midpoint temperature. To achieve BMCI of 120, the DO's °API gravity should not exceed 2.0. °API ravity is a rough indication of aromaticity; the lower the gravity, the igher the aromaticity. The ash content of the DO is affected by the reactor cyclone's erformance and catalyst physical properties. To meet the CBFS' ash equirement (maximum of 0.05 wt%), DO product may need to be iltered for the removal of the catalyst fines.
Coke
In a cat cracker, a portion of the feed, mostly from secondary racking and polymerization reactions, is deposited on the catalyst as oke. Coke formation is a necessary byproduct of the FCC operation; he heat released from burning coke in the regenerator supplies the eat for the reaction. The coke structure and the chemistry of its formation are difficult o define. However, coke in FCC comes from at least four sources: • Catalytic coke is a byproduct of the cracking of FCC feed to lighter products. Its yield is a function of conversion, catalyst type, and hydrocarbon/catalyst residence time in the reactor. • Contaminant coke is produced by catalytic activity of metals such as nickel, vanadium, and by deactivation of the catalyst caused by organic nitrogen. « Feed residue coke is the small portion of the (non-residue) feed that is directly deposited on the catalyst. This coke comes from the very heavy fraction of the feed and its yield is predicted by the Conradson or Ramsbottom carbon tests. • Catalyst circulation coke is a "hydrogen-rich" coke from the reactor-stripper. Efficiency of catalyst stripping and catalyst pore size distribution affect the amount of hydrocarbons carried over into the regenerator.
A proposed equation [1] to express coke yield is: Coke yield, wt% = g(Z,, . . . ZN) x (C/O)n x (WHSV)n~s x [e(ABc/RTrx)]
Where:
(Z,, . , .) = function of feed quality, hydrocarbon partial pressure, catalyst type, , etc.
Products and Economics
C/O WHS V
AEC R TRX
201
= 0.65 = cat-to-oil ratio = weight of hourly space velocity, weight of total feed/hr divided by weight of catalyst inventory in reaction zone, hr-1 = activation energy ~ 2,500 Btu/lb-mole (5828 J/G - mole) = gas constant, 1.987 Btu/lb-mole-°R (8.314 J/G- mole °K) = reactor temperature, °R
The coke yield of a given cat cracker is essentially constant. The FCC produces enough coke to satisfy the heat balance. However, a more important term is delta coke. Delta coke is the difference between the coke on the spent catalyst and the coke on the regenerated catalyst. At a given reactor temperature and constant CO2/CO ratio, delta coke controls the regenerator temperature. Reducing delta coke will lower the regenerator temperature. Many benefits are associated with a lower regenerator temperature. The esulting higher cat/oil ratio improves product selectivity and/or provides the flexibility to process heavier feeds. Many factors influence delta coke, including quality of the FCC eedstock, design of the feed/catalyst injection system, riser design, operating conditions, and catalyst type. The following is a brief discussion of these factors: * Feedstock quality. The quality of the FCC feedstock impacts the concentration of coke on the catalyst entering the regenerator. A "heavier" feed containing a higher concentration of coker gas oil will directionally increase the delta coke as compared with a "lighter," resid-free feedstock. * Feed/catalyst injection. A well-designed injection system provides a rapid and uniform vaporization of the liquid feed. This will lower delta coke by minimizing non-catalytic coke deposition as well as reducing the deposits of heavy material on the catalyst. • Riser design. A properly designed riser will help reduce delta coke by reducing the back-mixing of already "coked-up" catalyst with fresh feed. The back-mixing causes unwanted secondary reactions. • Cat/oil ratio. An increase in the cat/oil ratio reduces delta coke by spreading out some coke-producing feed components over more catalyst particles and, thus, lowering the concentration of coke on each particle.
02
Fluid Catalytic Cracking Handbook
« Reactor temperature. An increase in the reactor temperature will also reduce delta coke by favoring cracking reactions over hydrogen transfer reactions. Hydrogen transfer reactions produce more coke than cracking reactions. * Catalyst activity. An increase in catalyst activity will increase delta coke. As catalyst activity increases so does the number of adjacent sites, which increases the tendency for hydrogen transfer reactions to occur. Hydrogen transfer reactions are bimolecular and require adjacent active sites.
FCC ECONOMICS
The cat cracker's operational philosophy is dictated by refinery conomics. Economics of a refinery are divided into internal and xternal economics. Internal economics largely depends on the cost of raw crude and he FCC unit's yields. The cost of crude can outweigh the benefit from he cat cracker yields. Refiners who operate their units by a kind of ntuition may drive for more throughput, but this may not be the most rofitable approach. External economics are factors that are generally forced upon the efineries. Refiners prefer not to have their operations dictated by xternal economics. However, they may have to meet particular equirements such as those for reformulated gasoline. To maximize the FCCU's profit, the unit must be operated against ll its mechanical and operating constraints. Generally speaking, the ncremental profit of increasing feed is more than the incremental rofit from increasing conversion. The general target is to maximize asoline yield while maintaining the minimum octane that meets lending requirements. Because of the high cost of new units and the importance of the CC on refinery profitability, improvements should be made to the xisting units to maximize their performance. These performance ndices are: * * * *
Improving product selectivity Enhancing operating flexibility Increasing unit capacity Improving unit reliability
Products and Economics
203
* Reducing operating costs * Meeting product specifications » Reducing emissions
Product selectivity simply means producing more liquid products and ess coke and gas. Depending on the unit's objectives and constraints, elow are some of the steps that directionally improve product selectivity; » Feed injection. An improved feed injection system provides optimum atomization and distribution of the feed for rapid mixing and complete vaporization. The benefits of improved feed injection are reduced coke deposition, reduced dry gas yield, and improved gasoline yield. * Riser termination. Good riser termination devices, such as closed cyclones, minimize the vapor and catalyst holdup time in the reactor vessel. This reduces unnecessary thermal cracking and nonselective catalytic re-cracking of the reactor product. The benefits are a reduction in dry gas and a subsequent improvement in conversion, gasoline octane, and flexibility for processing marginal feeds. * Reactor vapor quench. LCO, naphtha, or other quench streams can be used to quench reactor vapors to minimize thermal cracking. * Reactor stripper. Operational and hardware changes to the stripper improve its performance by removing the entrained and adsorbed hydrocarbons. The benefits are lower delta coke and more liquid products. * Air and spent catalyst distribution. Modifications to the air and spent catalyst distributors permit uniform dispersion of air and spent catalyst into the regenerator. Improvements are lower carbon on the catalyst and less catalyst sintering. The benefits are a cleaner and higher-activity catalyst, which results in more liquid products and less coke and gas.
xamples of increasing operating flexibility are: * Processing residue or "purchased" feedstocks. Sometimes, the option of processing supplemental feed or other components, such as atmospheric residue, vacuum residue, and lube oil extract, is a means of increasing the yields of higher-value products and reducing the costs of raw material by purchasing less expensive feedstocks.
204
Fluid Catalytic Cracking Handbook
* ZSM-5 additive. Seasonal or regular use of ZSM-5 catalyst will center-crack the low-octane paraffin fraction of the FCC gasoline. The results are increases in propylene, butylene, and octane—all at the expense of FCC gasoline yield. * Catalyst cooler(s). Installing a catalyst cooler(s) is a way to control and vary regenerator heat removal and thus allow processing of a poor quality feedstock to achieve increased product selectivity, * Feed segregation. Split feed injection involves charging a portion of the same feed to a different point in the riser. This is another tool for increasing light olefins and boosting gasoline octane.
An example of increasing FCCU capacity is oxygen enrichment: * Oxygen enrichment. In a cat cracker, which is either air blower or regenerator velocity limited, enrichment of the regenerator air can increase capacity or conversion provided there is good air/ catalyst distribution and that the extra oxygen does not just burn CO to CO2.
In recent years, numerous mechanical improvements have been mplemented to increase the run length and minimize maintenance work during turnarounds. Examples include: * Expansion joints. Improvement in bellows metallurgy to Alloy 800H or Alloy 625 has reduced the failures caused by stress corrosion cracking induced by polythionic acid. Additionally, placing fiber packing in the bellow-to-sleeve annulus, instead of purging with steam, has reduced bellows cracking. Reliability has also increased with the use of dual ply bellows. * Slide or plug valves. Cast-vibrating of the refractory lining and stern/guide modifications have minimized stress cracking and erosion. « Air distributors. Improvements in the metallurgy, refractory lining of the outside branches, and better air nozzle design, combined with reducing L/D of the branch piping, have reduced thermal stresses, particularly during start-ups and upset conditions. * Cyclones. Changes in refractory anchor and material, the hanging system, longer L/D, and using more welds in the anchors have improved cyclone performance.
Products and Economics
208
UMMARY
Improving FCC unit profitability requires operating the unit against s many constraints as possible. Additionally, selective modifications f the unit's components will increase reliability, flexibility, and roduct selectivity, and reduce emissions.
REFERENCES
1. Venuto, P. B., and Habib, E. T., Jr., Fluid Catalytic Cracking with Zeolite Catalysts. New York: Marcel Dekker, 1979. 2. Lee, S. L., de Wind, M., Desal, P. H., Johnson, C., and Asim, M. Y., "Aromatics Reduction and Cetane Improvement of Diesel Fuels," Akzo Catalyst Chemical Seminar, Dallas, Texas, October 12, 1993. 3. Keyworth, D. A., Reid, T. A., Kreider, K. A., Tatsu, C. A., and Zoller, J. R., "Controlling Benzene Yield from the FCCU," presented at NPRA Annual Meeting, San Antonio, Texas, March 21-23, 1993. 4. Keyworth, D. A., Reid, T., Asim, M., and Gilman, R., "Offsetting the Cost of Lower Sulfur in Gasoline," presented at NPRA Annual Meeting, New Orleans, La., March 22-24, 1992. 5. Reid, T. A., "The Effect of ZSM-5 in FCC Catalyst," presented at World Conference on Refinery Processing and Reformulated Gasolines, San Antonio, Texas, March 23-25, 1993. 6. Engelhard Corporation, "Maximizing Light Cycle Yield," The Catalyst Report, TI-814. 7. Engelhard Corporation, "Prediction of FCCU Gasoline Octane and Light Cycle Crude Oil Cetane Index," The Catalyst Report, TI-769. 8. Engelhard Corporation, "Controlling Contaminant Sodium Improves FCC Octane and Activity," The Catalyst Report, TI-811. 9. Engelhard Corporation, "Catalyst Matrix Properties Can Improve FCC Octane," The Catalyst Report, TI-770. 0. Pine, L. A., Maher, P. J., and Wachter, W. A., "Prediction of Cracking Catalyst Behavior by a Zeolite Unit Cell Size," Journal of Catalysis, No. 85, 1984, pp. 466–476.
CHAPTER 7
Project Management and Hardware Design
Since 1942, when the first FCC unit came onstream, numerous rocess and mechanical changes have been introduced. These changes mproved the unit's reliability, allowed it to process heavier feedstocks, o operate at higher temperatures, and to shift the conversion to more aluable products. However, incorporating these changes in an existing unit is a major roject, usually more complicated than building a new unit. The two ritical components of a successful mechanical upgrade (or erection of a ew unit) are effective project management and proper design standards. This chapter addresses project management aspects of a revamp. It lso provides design guidelines that can be used by a refiner in electing the revamp components. The original driving force for a roject is often a particular mechanical problem or a process bottleeck. The ultimate objective of a revamp should be a safe, reliable, nd profitable operation.
PROJECT MANAGEMENT ASPECTS OF AN FCC REVAMP
The modifications/upgrades to the reactor and regenerator circuit are made for a number of reasons: equipment failure, technology changes, nd/or changes in processing conditions. The primary reasons for upgradng the unit are improving the unit's reliability, increasing the quantity nd quality of valuable products, and enhancing operating flexibility. The revamp (or erection of a new unit) requires successful execution f each phase of the project: • Pre-project • Process design
206
Project Management and Hardware Design
20?
* Detailed engineering « Pre-construction * Construction » Commissioning/start-up
Pre-project
In the pre-project phase, a refiner must take many steps "in-house" efore embarking upon a mechanical upgrade of an FCC unit. This is articularly true if the scope includes the use of new technology. ncluded in these pre-project activities are: * * * * «
Identifying the unit's mechanical and process constraints Identifying the unit's operational goals Optimizing the current unit's performance Obtaining a series of validated test runs Producing a "statement of requirement" or "revamp objectives" document « Selecting an engineering contractor
In many cases, a refiner decides to revamp a cat cracker and employ new technology without first identifying the unit's mechanical and rocess limitations. Sometimes money is spent to relieve a constraint nd the unit hits another constraint almost immediately. Failure to erform a proper constraint analysis of the existing operation can result n focusing on the wrong issues for the revamp. In addition, the evamp goals must match the refinery's overall objectives. The refiner should identify economic opportunities internally before pproaching a technology licensor. For example, what is the primary onsideration: more conversion, higher throughput, or both? At times, refiner may prefer to do the work internally, as opposed to hiring xternal resources, but all possible options should be explored. It may often be more economical to purchase the desired product rom another refiner than to produce it internally. The "market place" an be a less expensive source of incremental supply than the refiner's wn in-house production capabilities. Prior to a mechanical upgrade, the refiner must ensure that, given xisting mechanical limitations, the unit's performance has reached ts full potential with catalyst and operational changes. It is much asier to determine the effects of the mechanical upgrade with a well-
208
Fluid Catalytic Cracking Handbook
operated unit. Use of more cost-effective changes could achieve the same return as expensive revamp options when an optimized base case is determined. Any project yield improvements should be based on conducting a series of operating test runs. The test runs should reflect "typical'* operating modes. The results should be material/heat balanced. Another test run should be performed just prior to the revamp. A comparison of the results, pre- and post-revamp, should reflect no major changes in the catalyst reformulation. The revamp objectives, constraints, and requirements must be clearly stated in a statement of requirement document transmitted to the engineering contractor. The document should be sufficiently detailed and require minimum interpretation so as to avoid oversights and unnecessary site visits. Selection of a competent engineering contractor to perform process design and detail engineering is a key element in the overall success of a project. Important factors to consider when choosing a qualified contractor are: • • • •
Successful experience in FCC technology and revamps Related experience held by key members of the project team Current and projected workloads Biases and preferences as they relate to proven technologies and suppliers • The strength and chemistry of project team members • Range of services expected from the contractor (e.g., front-end engineering, detailed engineering, complete E.P.C. though start-up) • Engineering rate, mark-up, and unit cost of a "change order"
Process Design
Most companies have their own technology for the pre-design phase. For the purposes of this book, this phase will be referred to as frontend engineering design (FEED). FEED finalizes the process design basis so that the detailed engineering phase can commence. In most cases, FEED is performed by an engineering contractor, but sometimes it is prepared internally by the refiner. The FEED package must be sufficiently completed so that another engineering contractor can finish the detailed engineering with minimum rework.
Project Management and Hardware Design
209
In a revamp or construction of a new unit, which involves a techology upgrade, the engineering contractor commonly supplies a set f product yield projections. Refiners normally use these yield preictions as the basis when conducting an economic evaluation and erformance guarantee. It is essential that the refiner review these rojects carefully to ensure that they agree with the theory and approach xpressed by the licensor and that similar yield shifts have been bserved by other refiners installing similar technologies. In other words, the refiner should independently check the validity of projected ield improvements. During the FEED phase of the project, the engineering contractor an be asked to prepare two cost estimates. The initial cost estimate s usually prepared during the very early stages. The accuracy of this stimate is usually plus or minus 40% to 50%. This is a factored stimate of equipment and terms of reference. The second cost estimate s prepared at, or near, the completion of the FEED package. The ccuracy of this cost estimate is normally plus or minus 20%. This stimate is usually the basis for obtaining funding for the detailed ngineering stage. The format of the cost estimate is just as important as the content, The format can make a difference when proving whether or not the ontent is accurate. Therefore, the refiner should require that the ontractor present cost estimates in a format that is easy to understand nd analyze. In addition, the refiner's cost engineer should independently eview the cost estimate to ensure its accuracy and applicability, and lso to determine the contingency amounts that the owner should maintain in his funding plans. The FEED package typically consists of the following documents; • • • • • • « • • •
Project scope of work and design basis Process flow diagrams (PFD) Feedstock and product rates/properties Utility load data Operating philosophy, start-up, and shutdown procedures List of equipment, materials of construction, and piping classes Piping and instrumentation diagrams (P&ID), tie-in, and line list Instrument index, control valve, and flow element data sheets Electrical load, preliminary instrument, and electrical cable routing Preliminary plot plan and piping planning drawings
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Fluid Catalytic Cracking Handbook
• Specifications and standards • Cost estimate • Project schedule
Detailed Engineering
In the detailed engineering stage, the mechanical design of various components is finalized so that the equipment can be procured from the qualified vendors and the field contractor can install it. In preparing construction-issue drawings, the designer should pay special attention to avoiding field interference and allowing sufficient clearance for safety, operability, and maintainability. To ensure project-related safety, health, and environmental issues have been identified and resolved, the refiner should have in effect a process safety program that confirms the project complies with OSHA requirements. Procurement of materials in a timely fashion is a necessary part of detailed engineering. Successful procurement requires: • » • • • •
Early involvement of the procurement team Identification of long-lead and critical items Identification of "approved" vendors Identification of appropriate specification standards Competitive bid evaluation based on quality, availability, and price Establishment of a quality control program to cover fabrication inspection « Establishment of an expediting system to avoid unnecessary delays
Pre-construction
Activities performed in the pre-construction or pre-turnaround stage are essential to the success of the project. Some of the key activities are: • • • • •
Finalizing the project strategy plan Determining required staffing Identifying lay-down needs and securing specific areas Performing the detailed constructibility study Identifying additional resources, such as special equipment or special skills
Project Management and Hardware Design
211
* Completing an overall execution schedule * Reviewing the schedule to maximize pre-shutdown work * Maximizing pre-shutdown tasks
Construction
The guidelines for screening the general mechanical contractor and ther associated subcontractors are similar to those for selecting an ngineering contractor. The scope and complexity of the work will argely dictate the choice of the general contractor. Aside from availbility and quality of skilled crafts, the contractor's safety record and he dedication of the front-line supervisor to the worker's safety should e an important factor in choosing a contractor. Early selection of the general contractor is critical. The general ontractor should be brought in at 30% to 40% engineering completion o review the drawings and interface with the engineering contractor. Additionally, early constructibility meetings among the refiner, engineerng contractor, and general mechanical contractor will prove valuable n avoiding delays and reworks.
Pre-commissioning and Start-up
A successful start-up requires having in place a comprehensive plan hat addresses all aspects of commissioning activities. Elements of such plan include: * Preparation of the operating manual and procedures to reflect changes associated with the revamp * Preparation of training manuals for the operator and support groups « Preparation of a field checklist to inspect critical items prior to start-up * Development of a QA/QC certification system to assure that the installation has complied with the agreed standards and specifications
Post-project Review
Shortly after the start-up and before the general contractor leaves he site, a meeting should be held among key members of the project xecution team to obtain and document everyone's feedback on what went right, what went wrong, and what could have been done better.
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Fluid Catalytic Cracking Handbook
A summary of the minutes of this meeting should be sent to the participants and other relevant personnel. Once the operation of the unit has "lined out," it is time to conduct a series of test runs to compare performance and economic benefits of the unit with what was projected as part of the original project ustification. The results can also be used to determine if the unit's performance meets or exceeds the engineering contractor's performance guarantee.
Useful Tips for a Successful Project Execution
A successful project is defined as one that meets its stated objectives (safety, improved reliability, increased liquid yield, reduced maintenance costs, etc.) on or under-budget and is completed on or ahead of schedule. Some of the helpful criteria that ensure a successful project are as follows: • Plan carefully. This will minimize changes » Set the major reviews (PFDs, P&IDS, etc.) early, as opposed to waiting until the basic design is completed. This will minimize the project's cost by lessening rework • Assign dedicated refinery personnel to be stationed in the engineering contractor's office to coordinate project activities and act as a liaison between the refinery and the contractor • Make sure the key people from the operations, maintenance, and engineering departments are kept fully informed and that their comments are reflected early enough in the design phase to minimize costly field rework • Centralize all decision making to avoid project delays
Many aspects of FCC development have been the result of "trial and error." The development of present design standards is as much art as it is science. Consequently, it is appropriate to review some of the key developments that have influenced the current design philosophy behind the FCC reactor and regenerator: « Catalyst quality. The early FCC catalysts were neither very active nor very selective; yield structure contained too much coke at the
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expense of gasoline and other valuable products. Regenerators were operating in partial combustion at temperatures of around 1,100°F (590°C). The introduction of zeolite into the catalyst in the late 1960s brought about a significant impact on the FCC process. The zeolite-based catalyst allowed major yield shifts to light liquid products. * Higher-temperature operation. Advances in catalyst technology, the need to process heavier feedstocks, and the need to maximize the yield of desired products has resulted in operating the regenerator and reactor at higher temperatures. These higher temperatures have had deleterious effects on the mechanical components of the reactor/regenerator. The main drawback of a higher temperature operation is a higher induced stress, resulting in a lower load-carrying capacity of steel. * Refractory quality. Refractory lining was first developed mainly for use in the iron and steel industries. It was not until the refractory manufacturers began developing products specifically designed for FCC applications that tremendous improvements in erosion and insulating properties were realized. * More competitive refining industry. The run length of the early FCC units was very short; the unit was shut down every year or so. The general approach at the time was to make the necessary repairs and replace the damaged internal components. Once the industry became more competitive, the drive was to increase the unit's ran length, improve its reliability, and maximize the quantity and quality of desired products.
The evolution and improvement of the above-mentioned topics set he background for providing FCC design parameters. The following ections present the latest commercially-proven process and mechanical esign recommendations for FCC reactor-regenerator components. These design guidelines, though not universally agreed upon by very FCC "expert," can be useful to the refiner in ensuring that the mechanical upgrade of a unit will be safe, reliable, and profitable. The components of the reactor-regenerator circuit in which process nd mechanical design recommendations are provided are as follows: * Feed injection system * Riser and riser termination * Spent catalyst stripper
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* * * * *
Fluid Catalytic Cracking Handbook
Standpipe system Air and spent catalyst distributors Reactor and regenerator cyclones Expansion joints Refractory
Feed Injection System
Any mechanical revamp to improve the unit yields should always begin with installing an efficient feed and catalyst distribution system. This is the single most-important component of the FCC unit. An efficient feed and catalyst injection system maximizes gasoline yield and conversion at the expense of lower gas, coke, and decant oil and allows downstream technology to perform at its full potential. Ideally, a well-designed feed and catalyst injection system will achieve the following objectives: » Distribute the feed and regenerated catalyst throughout the crosssection of the riser to ensure that all feed components are subjected to the same cracking severity » Atomize the feed uniformly and instantaneously * Avoid re-contacting of the "spent catalyst" with the fresh feed * Produce proper oil droplet size to penetrate through the catalyst over the 360° cross-sectional area of the riser * Avoid erosion of the riser wall and attrition of the catalyst * Perform without plugging or erosion
Process Design Considerations for Feed Nozzles
Table 7-1 contains a summary of the process and mechanical design criteria commonly used in specifying high-efficiency feed nozzles. The mechanical design of any feed nozzle should be robust and easy to maintain. Its long-term mechanical reliability is critical in achieving he expected benefits of the upgrade. The following mechanical problems are often encountered: erosion of the nozzle tip(s), erosion of the riser wall, and blockage of the nozzles.
Catalyst Lift Zone Design Considerations
To maximize the benefits of feed nozzles, the regenerated catalyst must be distributed evenly throughout the cross-section of the riser.
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Table 7-1 Process and Mechanical Design Criteria for FCC Feed Nozzles
njectors
ressure drop
Nozzle exit velocity Dispersion media and rate
Orientation and location
eed nozzle type nsert material Nozzle tip
Multi-nozzles, typically 4 or 6 nozzles per riser located at the periphery of the riser and projected upward. 40 psi to 60 psi (280 to 420 kpa) at the design feed rate. 150 ft/sec to 300 ft/sec (45 to 100 m/sec) Steam, 1 wt% to 3 wt% of feed rate for conventional gas oil; 4 wt% to 7 wt% for residue feedstocks Radial; 3 to 4 riser diameters above the point where the regenerated catalyst enters the riser Readily retractable 304 H stainless steel Hard-surfaced inside and out
This requires pre-aceelerating the catalyst to the feed zone. Steam r fuel gas is often used to lift the catalyst to the feed injection. Figure -1 shows the design criteria of using steam as a lift media to deliver "dense" supply of catalyst to the feed nozzles,
Riser and Riser Termination
In most of today's FCC operations, the desired reactions take place n the riser. In recent years, a number of refiners have modified the CC unit to eliminate, or severely reduce, post-riser cracking. Quick eparation of catalyst from the hydrocarbon vapors at the end of the ser is extremely important in increasing the yield of the desired roduct. The post-riser reactions produce more gas and coke versus ess gasoline and distillate. Presently, there are a number of commercially roven riser disengaging systems offered by the FCC licenser designed o minimize the post-riser cracking of the hydrocarbon vapors. The process and mechanical guidelines used in designing most of he new or revamped units are summarized in Table 7-2. (text continued on page 218)
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Fluid Catalytic Cracking Handbook
Hydrocarbon Feed
Dispersion Steam Figure 7-1. Schematic of a typical feed nozzle.
Table 7-2 Process and Mechanical Design Guidelines for FCC Risers
Hydrocarbon residence time
Vapor Velocity
Geometry
Termination
Configuration Material
1 second to 3 seconds based on the riser outlet conditions. Depending on the degree of catalyst back-mixing in the riser, the catalyst residence time is usually 1.5 to 3.5 times longer than the hydrocarbons. 20 ft/sec (6 m/s) minimum (without oil feed), 65 ft/sec to 85 ft/sec (20 to 25 m/s) at the design feed rate. Vertical: to simulate plug flow and to minimize catalyst back-mixing Riser-cyclone separator attached to another separation device to minimize re-cracking of hydrocarbon vapors. External or internal. Carbon steel, "cold wall" as opposed to "hot wall,"
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To Reactor or Cyclone
T 3 to 5 Riser Daimeters Disf ISteam (Typical for Multiple Nozzles) Superficial
velocity 0.3 - 0.4 ft/sec
Steam or fuel gas
Drain Figure 7-2. Schematic of a typical catalyst lift system.
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Fluid Catalytic Cracking Handbook
text continued from page 215)
Spent Catalyst Stripper
A properly designed stripper minimizes the quantity of entrained and adsorbed hydrocarbons that are carried over to the regenerator with he spent catalyst. This goal should be accomplished by the use of
Table 7-3 Reactor-Stripper Process and Mechanical Design Criteria
Catalyst flux Stripping steam rate Stripping steam superficial velocity Catalyst residence time Steam quality
500-700 lb/min/ft2 (40 to 55 kg/sec/m2) 2-5 lb/1,000 Ib of circulating catalyst 0.5-0.75 ft/sec (.15-.25 m/sec) 1-2 minutes Superheated ~ 100°F (55°C)
Steam Distributor(s)
Minimum of two—upper and lower Concentric rings or pipe grid Minimum of one nozzle per ft2 of crosssectional area of the stripper
Nozzles
Orientation Exit velocity Pressure drop
L/D
Pointing downward 125-150 ft/sec (38-45 m/sec) Minimum of 1 psi (7 Kp) or 30% of the bed height Minimum of 5, or long enough to expand "vena contracta"
Material of Construction
Stripper shell
Distributors
Baffles Nozzles
Carbon steel, "cold wall" with 4 in. (10 cm) medium weight refractory lining Carbon steel, top distributor externally lined with 1 in. (~2 cm) thick erosion-resistant refractory Carbon steel Carbon steel, Schedule 80 minimum
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tripping steam. The major drawbacks of allowing the hydrogenich hydrocarbons into the regenerator are losses of liquid products, hroughput, and catalyst activity. Although proper design will greatly enhance stripper performance, is also very important to note that the performance of the stripper s also largely influenced by the type of feedstocks, catalyst, and perating conditions. The key process parameters for designing a tripper are listed in Table 7-3 (also see Figure 7-3). Catalyst flux is defined as catalyst circulation rate divided by the full" cross-sectional area of the stripper. For efficient stripping, it is esirable to minimize the catalyst flux to reduce the carryover of ydrogen-rich hydrocarbons into the regenerator. Up to a certain point, stripping efficiency is proportionate to increasing he stripping steam rate. However, excess stripping steam overloads
rifice
Dual Diameter Nozzle
Figure 7-3.
Schematic of a stripping steam distributor.
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Fluid Catalytic Cracking Handbook
the reactor cyclones, main column, and sour water treating system. Therefore, stripping steam should often be varied to determine the optimum rate. The optimum stripping steam rate usually corresponds to a value in which there will be no reduction in the regenerator bed temperature. Catalyst residence time in the stripper is determined by catalyst circulation rate and the amount of catalyst in the stripper. This amount usually corresponds to the quantity of the catalyst from the centerline of a "normal" bed level to the centerline of the lower steam distributor. A higher catalyst residence time, though it increases hydrothermal deactivation of the catalyst, will improve stripping efficiency. It is important to note that, depending on the stripper pressure and emperature, a certain fraction of stripping steam is carried with the spent catalyst into the regenerator. Example 7-1 shows how to determine this amount. Example 7-1
Calculate the amount of entrained stripping steam into the regenerator from a reactor-stripper with the following conditions:
Catalyst Catalyst Stripper Stripper Catalyst
skeletal density flowing density operating pressure operating temperature circulation rate
= = = = =
150 lb/ft3 (2400 kg/m3) 35 lb/ft3 (560 kg/m3) 25 psig (173 Kp) 980°F (525°C) 40 short ton/min = 4,800,000 Ib/hr (2200 mt/hr)
Solution:
Volume of entrained steam = 1/35 - 1/150 = 0.0219 ft3 of steam/lb of circulating catalyst (0.0014 mVkg) P
M P + 14.7 ~ 10.73* t + 460
Where:
p = Gas or vapor density, lb/ft3 M = Molecular weight P = Pressure, lb/in.2 gauge = Temperature, °F
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. . 18 25 + 14.7 Steam density = x 10.73 980 + 460
= 0.0462 Ib of steam/ft3 of steam (.74 kg/m 3 )
Entrained steam = (0.0219 ft3 of steam/lbs of catalyst) x (0.0462 Ibs of eam/ft3 of steam) x 4,800,000 Ib/hr = 4,858 Ib/hr (2204 kg/hr)
tandpipe System
Proper standpipe design is one of the most important factors in btaining good circulation. The standpipe provides the necessary head ressure required to circulate the catalyst. A standpipe assembly is ypically comprised of three major parts: hopper, standpipe, and slide r plug valve. The function and design of each part is described below.
Hopper Design
A catalyst hopper (see Figure 7-4) provides sufficient time for the nitial deaeration of the catalyst. Proper catalyst deaeration should
Debris Guard
Figure 7-4. Schematic of a typical catalyst hopper.
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Fluid Catalytic Cracking Handbook
maximize catalyst density and maintain the catalyst in a "fluidized" tate. Table 7-4 contains the key process parameters used in designing tandpipe hoppers.
Standpipe
The standpipe provides the necessary head pressure required to achieve proper catalyst circulation. Standpipes are sized to operate in he fluidized region for a wide variation of catalyst flow. Maximum catalyst circulation rates are realized at higher head pressures. The higher head pressures can only be achieved when the catalyst is luidized. Table 7-5 contains typical process and mechanical design criteria for standpipes.
Slide or Plug Valve
The slide or plug valve regulates the flow of catalyst between the eactor and regenerator. The slide valve also provides a positive seal against reversal of the hydrocarbons into the regenerator or hot flue gas into the reactor. Table 7-6 summarizes typical process and mechanical parameters for designing slide valves. The formula to calculate catalyst circulation rate through a slide valve is: W = Ap x Cd x 2,400x ^AP^p
Where:
W A5 Cd AP
= Catalyst circulation rate, Ib/hr (kg/hr) = Port or orifice opening, in.2 (m2) = Discharge coefficient of = 0.85 = Valve pressure drop, psi (pascals) = Density of catalyst in the standpipe = lb/ft3 (kg/m3)
Table 7-4 Process Design Considerations for Standpipe Hoppers
Hopper entrance diameter Angle of cone Desired catalyst density Catalyst velocity
2.25 times the standpipe diameter 35°–45° off the vertical 40–45 lb/ft3 (640-720 kg/m3) 0.5–1.0 ft/sec (.15-.3 m/sec)
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Table 7-5 Process and Mechanical Design Criteria for Catalyst Standpipes
atalyst flux
100-300 lb/sec/ft2 (500–1,450 kg/sec/m2)
atalyst velocity
2-6 ft/sec (.6-2 m/sec), target for 4 ft/sec (1.3 m/sec)
Desired density
40–45 lb/ft3 (650-800 kg/m3)
Geometry
Vertical or sloped at maximum angle of 45° (off vertical)
Material
Carbon steel, "cold wall" with 5 in. (12 cm) thick heavy weight, erosion-resistant refractory lining
upplemental aeration
Every 5-8 ft (1.5-2.5 m) along the standpipe; use rotameters to regulate aeration flow
Example 7-2
To illustrate the use of the above equation, determine the catalyst irculation rate from the following information: Slide valve DP
=5
psi (35 K p )
Slide valve opening
=
40% corresponding to a port opening of 200 in. 2 (1,290 cm2)
Catalyst density
=
35 lb/ft3 (560 kg/m3)
herefore: W = 200x0.85x2,400x V35x^ = 5,397,333 Ib/hr (560kg/hr) = 45 short tons/min (41 mt/min)
Air and Spent Catalyst Distributor
The main purpose of the regenerator is to produce a clean catalyst, minimize afterburn, and reduce localized sintering of the catalyst. For fficient catalyst regeneration it is very important that the air and the pent catalyst are evenly distributed. Although, in recent years, the esign of air distributors has improved significantly, the same cannot
24
Fluid Catalytic Cracking Handbook Table 7-6 Process and Mechanical Design guidelines for Slide Valves
perating pressure drop
Minimum 1.5 psi (10 Kp), maximum 10 psi (70 KP)
% opening @ design circulation
40%–60%
Material
Shell: carbon steel with 4-5 in. (10-12 cm) thick heavy weight, single-layer, cast-vibrated refractory with needles. Internals: 304H stainless steel for temperature >1,200°F (650°C) and Grade H, \\% chrome for <1,200°F. Internal components exposed to catalyst should be refractory-lined for erosion resistance. Sliding surfaces should be hard-faced, minimum thickness \ in. (3 mm).
onnet design
Sloped bonnet (30° minimum) for self-draining of catalyst.
urge
Purgeless design of stuffing box. Guides: slotted, hard-surfaced, and supplied with purge connections (normally closed). Nitrogen is the preferred choice of purge gas.
ctuator type
Electro-hydraulic for fast response and accurate control.
ctuator response time
A maximum of 3 seconds
e said for spent catalyst distributors. This is particularly true in the ase of side-by-side FCC units. Most side-by-side units suffer from neven distribution of the spent catalyst. A well-designed air distributor system has the following characteristics: • Distributes the air uniformly to the spent catalyst • Mechanically withstands a wide range of operating conditions, including start-up, shutdown, normal operation, and upset conditions • Provides reliability with minimum maintenance
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Two types of air distributors have been used in the past. In some f the original side-by-side units, both the air and the catalyst flowed hrough the distributor. In virtually all the air distributors designed oday, only air flows through the distributor. Flat grid plate, dome, pipe, ring, and flat pipe grid are the five ypical configurations of air distributors presently being used. The most ommon types are ring and flat pipe grid distributors. Overall, pipe rid is preferred over air ring mainly due to a more uniform coverage nd a lower discharge velocity, which tends to minimize catalyst ttrition. In addition, the pipe grid maintains the same coverage of the ross-sectional area regardless of the air rate. Rings obtain their overage from jet penetration and the coverage will be reduced at air ates less than design value due to lower velocity. The three primary factors affecting the mechanical performance of he air distributor system are erosion, thermal expansion, and mechanical ntegrity of supports. The distributor's design should reflect the erosive ature of high catalyst/air velocities, thermal expansion for various perating conditions, and corresponding supports to minimize thermal xpansion loads. The process and mechanical design considerations of n air distributor are shown in Table 7-7 (also see Figure 7-5),
Reactor and Regenerator Cyclone Separators
A cyclone separator is an economical device for removing parculate solids from a fluid system. The induced centrifugal force (see igure 7-6) is tangentially imparted on the wall of the cyclone cylinder. his force, with the density difference between the fluid and solid, ncreases the relative settling velocity. Cyclone separators are extremely important to the successful operaon of the cat cracker. Their performance can impact several factors, ncluding the additional cost of fresh catalyst, extra turnaround mainenance costs, allowable limits on emission of the particulates, incremental energy recovery in the wet gas compressor, and hot gas expander. Designing an "optimum" set of cyclones requires a balance between he desired collection efficiency, pressure drop, space limitations, and nstallation cost. Cyclone process and mechanical design recommendaons are shown in Table 7-8. (text continued on page 230)
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Fluid Catalytic Cracking Handbook
Table 7-7 Process and Mechanical Design Criteria for Air Distributors
Recommended type
Pipe grid distributor
Nozzle exit velocity
100-200 ft/sec (30–60 m/sec)
ressure drop
1.5–2.0 psi (10–15 Kp) at design air rate; 10%–30% of the bed static head at minimum air rate for downward-pointing nozzles
Material
304H stainless steel, externally lined with 1 in (2.5 cm) thick erosion-resistant refractory
Branch pipe
L/D ratio of less than 10 to minimize the support requirement and vibration
Branch arm connection
Continuous pipe through the main header and slotted opening Forged fittings instead of miters for supporting the headers; the forged fittings minimize failures due to stress cracking
ittings
Nozzles
ype and orientation
Dual diameter nozzles with orifice in the back of nozzle; downward at 45°
ength
Minimum of 4 in. (10 cm)
/D
5/1 to 6/1
ocation of first nozzles
8–12 in (20-30 cm) from the edge of the slot in the branch arm
Drain size
in. (~1 cm) spaced evenly
he pressure drop of the nozzle's orifice can be calculated from the equation below: Po
2 x gc x 144 1C
here: Vo po gc C
= Velocity of air through the orifice, ft/sec = Density of air, lb/ft3 = Gravitational constant, 32.2 ft/sec2 = Discharge coefficient = 0.85
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Orifice
Figure 7-5. Typical layout of a pipe grid distributor.
227
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Fluid Catalytic Cracking Handbook
Outlet Tube
Catalyst/Vapor Barrel
Cone
Dustbin
Dipleg
Figure 7-6.
Schematic of a typical cyclone.
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Table 7-8 Process and Mechanical Design Guidelines for Reactor and Regenerator Cyclones
Vapor Velocities at Design Feed Rate
Cyclone Type
Reactor, single-stage
Reactor or regenerator, primary or first- stage
Reactor, secondary or second-stage
Regenerator, secondary or second-stage
Minimum cyclone velocity
Inlet ft/sec (m/sec)
Outlet ft/sec (m/sec)
65-75 (20-23) 60–70 (18-21) 75–85 (23-26) 80-85 (-25) 25–35 (8–10)
100-120 (30-35) 65-75 (20-23) 100-110 (30-33) 120-140 (36-40)
Minimum overall collection efficiency = 99.995% ingle-stage or primary dipleg mass flux = 100-125 Ib/ft2/sec (500–600 kg/m2/sec)
Parameters
Cyclone length/inside diameter Cyclone inlet height/ width
Reactor cyclones
Regenerator cyclones
Dimensional Specifications
Single-Stage
Primary
Secondary
5.0
3.5–4.5
4.5-5.5
2.2-2.5
2.2–2.5
2.2-2.5
Material
Chrome-rnoly alloy lined with 1 inch-thick erosion-resistant refractory. 304H stainless steel, lined with 1 inch-thick erosion-resistant refractory Chrome-moly alloy, internal or external depending on cyclone geometry. Carbon steel, "cold wall" design to avoid hightemperature stress cracking.
enetration of the gas outlet tube into each cyclone should be at least 80% of the cyclone nlet duct height. he projected vortex (see Figure 7-6) should be a minimum of 15" (40 cm) above the dustowl outlet. rickle valves should be partially shrouded.
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Fluid Catalytic Cracking Handbook
text continued from page 225)
Expansion Joint
Efforts should be made to eliminate the use of expansion joints in process piping. However, if needed, the expansion joints are used o mitigate the pipe stresses caused by large thermal movements. Table 7-9 lists the recommended mechanical design criteria for expansion joints,
Refractory
In the FCC unit, refractory is used extensively to resist heat, erosion, corrosion, or any combination of the three. The insulating properties protect the reactor-regenerator components from the elevated temperature. The erosion properties protect the equipment from a highvelocity catalyst. Refractory also provides chemically resistant protecion against corrosion reactions. To ensure a high quality refractory application, the refiner must require the following:
Table 7-9 Mechanical Design Recommendation for Expansion Joints
Shell's material
Carbon steel, "cold shell design," castvibrated 5 inch (12 cm) thick refractory lining
Bellow's material
Inconel 625
Purge requirement
Packed bellows, no purge
Configuration of bellows
Two-ply bellows with pop-out indicator for detecting leakage; each bellows should be capable of maintaining the full pressure
Packing material
Ceramic fiber blanket
Minimum bellows temperature
400°F (205°C) to minimize condensation and subsequent acid attack
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Table 7-10 Recommended Refractory Specifications for Various FCC Components
Catalyst carrier lines (riser, standpipe, wye, slide valves, etc.)
Reactor, regenerator, and stripper shells
Reactor and regenerator cyclones
Cyclone diplegs
Air distributors
Combustion air preheater
4-5 inch (~10 cm) thick, cast-vibrated, single-layer, heavyweight (140-160 lb/ft3) (2,250–2,550 kg/m3), reinforced with stainless steel fibers. The refractory is supported by footed "wavey V~" type anchors, 304 stainless steel, and spaced on 8-12 inch (20-30 cm) centers. The anchor's length should be 1 inch (-2 cm) shorter than the thickness of the refractory lining, 4 inch (10 cm) thick, "gunned," single-layer, medium weight (80-100 lb/ft3) (1,300–1,600 kg/m3), with steel fibers, supported by footed "wavey V" type anchors with plastic tips. Internally lined with 1 inch (2.5 cm) thick, erosion-resistant refractory attached by 304 stainless steel hexmetal anchors with every hex on every other row welded independently. The refractory should be cut out flush with the top of hex. The hexmetal should be discernible after the refractory installation. Primary diplegs: internally lined with 1 inch (2.5 cm) thick, erosion-resistant lining. Secondary diplegs: lined internally with 1 inch refractory on the top 5 feet (1.5 m) from the dust-bowl. Externally, diplegs should be lined where the spent catalyst returns and other known turbulent or wear areas. Externally lined with 1 inch (2.5 cm) thick, erosion-resistant refractory supported by 304H stainless steel hexmetal anchors or other individually unitized anchors. The refractory should be cut out flush with the top of hex. The hexmetal should be discernible after the refractory installation. 4-5 inch (10–12 cm) thick, medium-weight, gunned, with needles, and thermal shockresistant refractory.
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Fluid Catalytic Cracking Handbook
* A mill test report of each pallet of refractory * Job-site qualification of all refractory installers before the actual work is performed * Production test samples to ensure that the required physical properties are met
To ensure longer service life for the refractory application, a thorough ryout procedure should be implemented and practiced. Slow thermal ryout is required to minimize and eliminate refractory spalling. General refractory selection guidelines for various components of he reactor-regenerator circuit are listed in Table 7–10.
SUMMARY
Because the FCC is one of the most efficient conversion processes, very effort must be made to ensure that its operation is optimized. This may require revamping the existing unit to incorporate the latest roven technologies. Carrying out a successful revamp requires proper dentification of the unit's mechanical and process limitations. Failure o perform a proper constraint analysis of the existing operation can esult in focusing on the wrong issues for the revamp. The process nd mechanical design guidelines detailed in this chapter will be eneficial for the proper selection of the revamp components of the eactor-regenerator.
. Murphy, J. R., Air Products-HRI, "The Development of Feed and Air Distribution Systems in Fluid Catalytic Cracking," presented at the 1984 Akzo Chemicals Symposium, Amsterdam, The Netherlands . Tenney, E. D., Marsulex Environmental Technologies, "FCC Cyclone Problems and How They Can Be Over Come with Current Designs" presented at the Grace-Davison FCC Technology Conference, Toledo, Spain, June 3-5, 1992. . Cabrera, C. A., Hemler C. L., and Davis, S. P., UOP LLC, "Improve Refinery Economics Via Enhanced FCC Operations," presented at Katalistik's 8th Annual Symposium, Budapest, Hungary, June 1–4, 1987. . Kalen, B., and O'Broin, A. E., Emtrol Corporation, "A Unique Solution— Cyclone Support in a High-Temperature Environment," presented at Katalistik's 8th Annual FCC Symposium, Budapest, Hungary, June 1-4, 1987,
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Ratterman, M., "An Approach to the Design and Analysis of Data from the Standpipe System on FCC Units," Gulf Research and Development, Pittsburgh, Pennsylvania, October 1983. Wrench, R. E., and Glasgow, P. E., The M.W. Kellogg Company, "FCC Hardware Options," Paper No. 125C, presented at the AIChE National Meeting, Los Angeles, California, November 17–22, 1991.
CHAPTER 8
Troubleshooting
The cat cracker plays a key role in the overall profitability of the efinery. It must operate reliably and efficiently. It must also operate afely and comply with federal, state, and local environmental requirements. A typical FCC unit circulates tons of catalyst per minute, processes various types of feedstock and uses hundreds of control oops, any of which can make operation difficult. Proper troublehooting will ensure that the unit operates at maximum reliability and efficiency while complying with environmental concerns. Troubleshooting deals with identifying and solving problems. Problems can be immediate or long term and can be associated with off-spec products, poor efficiency, process improvements, capacity increases, or potential shutdown items. Problems can be related to management, operation, hardware and equipment, or process issues. Solutions can nclude improved operating procedures and training, preventative maintenance, or installation of new equipment or controls. This chapter outlines fundamental steps toward effective troublehooting. It provides a practical and systematic approach to developing a solution. General guidelines are provided for identifying problems and determining a diagnosis. In particular, the following FCC-related problem areas are addressed in detail: • • • • • • • • •
Catalyst Circulation Catalyst Loss Coking/Fouling Flow Reversal High Regenerator Temperature Afterburn Hydrogen Blistering Hot Gas Expander Products Quality and Quantity
234
Troubleshooting
235
This chapter is written with the unit process engineer in mind. No matter where the problem originates, the unit process engineer will e the point person for solving it.
A successful troubleshooting assignment will require someone to: * * • •
Be a good listener Gather historical background Evaluate "common" and "uncommon" causes of problems Examine goals and constraints to verify the applicability of the present operation
Management, engineering, and operations departments perceive problems ifferently. Frequently, there is someone familiar with the operation who most likely knows the symptoms and possibly can offer a solution o the problem, but for various reasons, people who are in a position o implement the solution may not have thought to ask him or her. ypically, those closest to the problem are the unit operator and maintenance foreman and they will offer the most valuable input. All our operating shifts should be consulted. Do not draw any conclusions efore gathering all applicable facts. Examine similar problems that have occurred previously in the ystem to determine how they were diagnosed and solved. Review the perating and maintenance records. Compare the performance of the ormally operating unit to the current problematic operation. Make ure that all the unit trends are current, including catalyst data and eat and weight balance data. Note any changes that might relate to he problem. Reliable historical data always helps to identify and iagnose problems. Begin by listing all potential causes or combinations of causes, using "brainstorming" approach. Then, systematically rule out each cause. Do not eliminate uncommon causes too quickly—if it were an easy roblem it would have already been taken care of it. Additionally, nsure that limits outlined by process and equipment documentation re consistent with the actual operation of the unit. Most FCC problems are due to changes in the feedstock, catalyst, perating variables, and/or mechanical equipment. As previously stated,
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Fluid Catalytic Cracking Handbook
he solution can take the form of improving yields, avoiding shutdowns, or increasing unit reliability.
CATALYST CIRCULATION
Catalyst circulation is like blood circulation to the human body. Without "proper" catalyst circulation, the unit is dead. Troubleshooting circulation problems requires a good understanding of the pressure balance around the reactor-regenerator circuit and the factors affecting catalyst fluidization. The fundamentals of fluidization and catalyst circulation are discussed in Chapter 5. Catalyst circulation is largely influenced by the physical layout of he unit and the fluidization properties of the catalyst. Some cat crackers circulate with ease regardless of the catalyst's physical properties. However, in other designs, the unit can experience circulaion difficulties with minor changes in catalyst properties. There are two types of circulation problems in the FCC unit that equire troubleshooting. The first occurs when the apparent maximum circulation has been reached, and the second occurs when the unit is experiencing erratic circulation.
Limited Circulation
Often, the feed rate or the conversion cannot be increased because he unit is circulation-limited. Effective troubleshooting of this limitaion requires having in place a methodology such as the one shown n Figure 8-1. To begin troubleshooting, start by conducting a single-gauge pressure survey of the reactor-regenerator and main column circuits (Chapter 5). The survey should provide the foundation for pinpointing the cause(s) and identifying solutions. Higher catalyst circulation usually requires opening the regenerated and spent catalyst slide (or plug) valves. Higher circulation increases he pressure drop in the riser and in the reactor cyclones, lowering he differential pressure across the slide valves. This causes the valves o open further, until the unit finds a new balance. To ensure proper control and to protect from reversal of hydrocarbons to the regenerator or the back-flow of hot flue gas to the (text continued on page 240)
I
O
o o
o
.c
CO
5
I
O
< ^> CO
0
S
2 a.
CO
ke sure instrument dings are correct
Verify Aeration Gas flow to maximize pressure build-up
Improper placement of the Aeration taps
Use Rotameters instead of Restriction Oriffces(RO's)
i
Restriction Orifices are either plugged or improperly sized
Low Catalyst density in the Standpipe
Troubleshooting catalyst circulation.
Check if the Catalyst properties have changed Figured-IB.
Too much, too little or no Aeration Gases
-fluidization of the talyst in the Standpipe
Insufficient pressure build-up in the Standpipe
Low Pressure Upstream of the Slide Valve
r
r
v
\ f
• Adjust the Pumparound rates • Add Top or Side P/A
\^
^/
^.
High delta P across the Main Fractionator
V
J
\(
Refer to 'Coking/Fouling' Troubleshooting Section
\.
J
High Delta P across the Reactor Ovhd. vapor line
Figure8-1 C. Troubleshooting catalyst circulation.
• Add Fins to the Trim Coolers -19 flns per inch * Water Wash the Condensers • Reduce No. of tube passes on the water side * Check pressure drop between Fin-Fans and Trim Coolers
)r
High delta P across the Overhead Condensers
High Pressure Downstream of the Slide Valve
V
\f
V
• Increase Fluffing Gas or Steam to the base of Riser « Replace the Curved section of the Riser
J
High Delta P across the Riser
40
Fluid Catalytic Cracking Handbook
text continued from page 236)
eactor, slide valves are typically operated with a 25% to 60% opening nd a pressure differential of 2 psi to 6 psi (15–40 Kp). Any increase n the throughput or the conversion will either increase the opening, educe the differential, or both.
Causes of Insufficient
Circulation
A lower pressure differential across the slide valve and/or a higher han "normal" slide valve opening are common evidence of a unit's eaching its circulation limitation. The causes of low differential are • Insufficient pressure buildup upstream of the slide valve * Excessive pressure drop downstream of the slide valve
Catalyst circulation can also be limited by mechanical problems with he slide/plug valves. They may have limited travel or will not open ully. This is indicated by the lack of response when an adjustment is made. The differential pressure and the flow will not change. Troublehooting the valves is a function of the valve design and the vendor will supply troubleshooting information,
Low Upstream Pressure Insufficient pressure upstream of the slide valve can be caused by: • Partial plugging in the standpipe * Too little, too much, or no aeration gas either with the catalyst entering the standpipe or along the standpipe
In a properly designed standpipe, the flow of the catalyst develops smooth and uniform static head over the entire length of the standpipe, provided the catalyst entering the standpipe is properly fluidized. This buildup of pressure head provides the necessary driving force for atalyst circulation. However, as the catalyst travels down the standpipe, it can lose fluidity due to compression of interstitial gas being arried with the catalyst. This is particularly true in a long standpipe nd in low-pressure regenerators operating at low pressure. To retain fluidity of the catalyst and to maintain catalyst densities n the 35 to 45 lb/ft3 (560–720 kg/m3) range (the fluid range), many standpipes require external aeration gas to be injected into the down-flowing
Troubleshooting
241
atalyst. The correct quantity of aeration medium and the correct ocation of aeration taps are essential in achieving optimum catalyst ensity (see Figure 8-2). External aeration is not ordinarily needed in hort standpipes (less than 20 feet) (6 m) because sufficient gas is sually drawn into the standpipe to keep the catalyst fully fluidized.
Troubleshooting Upstream Pressure
Effective troubleshooting of the circulation system requires a methdology similar to the procedures outlined in Figure 8–1. Some of the ey steps are as follows: • Obtain a pressure/density profile upstream and downstream of the slide valves. • Verify any changes in catalyst physical properties. • Ensure that the correct amount of aeration gas is injected along the standpipes. One procedure is to vary the aeration flow until the maximum slide valve differential is observed.
Restriction orifices with upstream pressure regulators are frequently mployed to distribute aeration gas into the standpipes. The orifices ..Slip Joint Rotameter or Orifice/Pressure Regulator
Steam or Nitrogen
Figure 8-2.
Typical standpipe aeration.
42
Fluid Catalytic Cracking Handbook
re sized for critical flow so that the constant flow of aeration gas is roportional to upstream pressure regardless of changes in the downtream pressure. Once the unit is running well, it is often assumed that the aeration ystem is sized properly, but changes in the catalyst physical properties nd/or catalyst circulation rate may require a different purge rate. It hould be noted that aeration rate is directly proportional to catalyst irculation rate. Trends of the E-cat properties can indicate changes n the particle size distribution, which may require changes in the eration rate. Restriction orifices could be oversized, undersized, or lugged with catalyst, resulting in over-aeration, under-aeration, or no eration. All these phenomena cause low pressure buildup and low lide valve differential.
High Downstream Pressure
Sometimes insufficient differential across the regenerated catalyst lide valve is not due to inadequate pressure buildup upstream of the alve, but rather due to an increase in pressure downstream of the slide alve. Possible causes of this increased backpressure are an excessive ressure drop in the "Y" or "J-bend" section, riser, reactor cyclones, eactor overhead vapor line, main fractionator, and/or the main fracionator overhead condensing/cooling system. The pressure drop in the "Y" or "J-bend" section could be from mproper fluidization or a flaw in the mechanical design. There are ften fluffing gas distributors in the bottom of the "Y" or along the J-bend" that are designed to promote uniform delivery of the catalyst nto the feed nozzles. Mechanical damage to these distributors or too ttle or too much fluffing gas affect the catalyst density, causing ressure head downstream of the slide valve. The riser pressure drop is related mainly to the catalyst circulation ate and the slip factor. Catalyst circulation rate is largely a function f the oil feed rate, the reactor temperature, and the feed temperature. ncreasing the feed rate, reactor temperature, or lowering the feed emperature will increase the pressure drop across the riser. Slip factor is defined as the ratio of catalyst residence time in the iser to the hydrocarbon vapor residence time. Some of the factors ffecting the slip factor are circulation rate, riser diameter/geometry, nd riser velocity.
Troubleshooting
243
High pressure in the riser could also be due to insufficient fluidizaon gas in the base of the riser. Fluffing gas will vary the catalyst ensity; more fluffing gas lowers the density in the system and the ackpressure on the slide valve. The pressure drop across the reactor cyclones, reactor vapor line, main fractionator, and main column overhead condensing/cooling ystem can be too high. The pressure drop is primarily a function of apor velocity. Any plugging can increase the pressure drop,
roubleshooting Downstream Pressure Use the same procedure as for upstream pressure.
Erratic Circulation
Erratic circulation occurs when the catalyst is not developing a mooth and uniform static head over the entire length of the standpipe. When this happens, the catalyst packs and bridges across the standpipe. ymptoms of erratic circulation include: • • • • • • • •
Severe vibration and movement of the standpipes A noise similar to train chugging Sudden loss of the pressure above the slide valve Fluctuation in the slide valve delta P Ragged reactor temperature and/or stripper level control Pressure swings in the regenerator and the gas plant Cycling of the slide valve Other instrumentation problems
Causes of Erratic Circulation Several factors contribute to erratic circulation. Included are: • A foreign object, such as a piece of refractory, partially obstructing the flow of catalyst in the standpipe. • Improper aeration—either too much or too little. • Changing catalyst properties. This can be due to changes in the fresh catalyst's physical properties and/or malfunctioning of the cyclones.
44
Fluid Catalytic Cracking Handbook
roubleshooting Erratic Circulation To troubleshoot erratic circulation, one must: » Verify that all the instrument readings are "telling the truth" « Verify all the aeration taps are open * Make sure that neither too much nor too little aeration gas is being applied * Verify aeration gas is not wet * Verify that the fresh catalyst properties have not changed * Verify any recent design changes in the standpipes and/or catalyst hopper * Check recycling of the regenerator fines and/or the slurry recycle
Figure 8-3 shows a step-by-step approach to troubleshooting rratic circulation.
CATALYST LOSSES
Catalyst losses will have adverse effects on the unit operation, the nvironment, and operating cost. Catalyst losses appear as excessive
Problem:
Catalyst is not developing a smooth and uniform static head over the entire length of the Standpipe
Ir
Symptom:
Evidence:
Figure 8-3A.
Catalyst packs and bridges across the Standpipe
Severe vibrations and movement of the Standpipes Fluctuation in the Slide Valve delta P A chugging noise similar to "Train" noise Ragged Reactor temperature and/or Stripper level control Sudden loss of pressure above the Slide Valve Pressure swings in the Regenerator and Gas Plant Troubleshooting erratic catalyst circulation.
Troubleshooting
245
ause:
r Improper Aeration Aeration is not adequate to maintain fluidity due to compression
Solution: > Make sure that neither too much nor too little Aeration Gas is being used
Figure 8-3B.
Catalyst is too coarse: Catalyst fines content is too low
Recycling of Regenerator fines and/or Slurry recycle Consider another Catalyst with a different particle size distribution Check Catalyst Iron Content Check properties of recent Catalyst purchase
A foreign object has partially restricted flow of Catalyst in Standpipe
• Use Gamma Ray to confirm plugging • Use high pressure Steam or Nitrogen to dislodge material
Troubleshooting erratic catalyst circulation.
Cause: Catalyst is defluidized to its bulk density; Low Regenerator Pressure
Wet Aeration Air or Steam
Solution: • Increase Regenerator pressure
Figure 8-3C.
Ensure Aeration medium is dry
Recent mechanical revisions to the Unit
Verify the design or\ Hopper and/or Standpipe Plugged vent line on external Hopper Verify amount and orientation of Aeration injection
Troubleshooting erratic catalyst circulation.
4i
Fluid Catalytic Cracking Handbook
arryover to the main fractionator or losses from the regenerator. Evidences of catalyst losses are: * An increase in the ash and BS&W content of the slurry oil * An increase in the recovery of catalyst fines from the electrostatic precipitator or the tertiary separator » An increase in the opacity of the precipitator stack gases * A decrease in the 0 to 40 microns fraction of the equilibrium catalyst or an increase in average particle size * A gradual loss of the catalyst level in the reactor stripper and/or in the regenerator
Causes of Catalyst Losses Common causes of catalyst losses include: * * * *
Changes Changes Changes Changes
in in in in
catalyst properties operating conditions the mechanical condition of the unit operating practice
Changes in the fresh catalyst's physical properties may contribute o catalyst losses. The losses could be due to the fresh catalyst's being soft." "Softness" is evidenced by the quality of the catalyst binder nd the large amount of 0–40 microns. It will increase the attrition endency of the catalyst and thus its losses. Changes in operating parameters also affect catalyst losses. Exmples are: * An increase and/or decrease in catalyst loading to the cyclones * Overloading the cyclones, even at a constant/or higher efficiency, will result in higher catalyst losses * An increase in the feed atomizing and/or stripping steam, causing catalyst attrition and generating fines * An addition of a large amount of steam to the regenerator, particularly to the torch oil nozzles, again causing catalyst attrition
Often, the main cause of catalyst losses is a change in the mechanical ondition of the unit. Examples are: * Trickle valves are either stuck "closed" or "open," possibly due to the hinges being warped or bound. "Warpage" could be due to
Troubleshooting
• • • • • •
• »
247
exposure to high temperatures. Erosion could be due to excessive gas leakage into the diplegs. Trickle valves have fallen off due to inadequate bracing and/or high superficial gas velocity in the regenerator. Holes have formed in the diplegs because of high cyclone velocity or external impingement. Spalled coke or refractory is lodged in the diplegs. This can be caused by improper curing or inadequate refractory supports. Cracks can form in the internal plenum, possibly due to thermal stresses. The dipleg diameter is either too small or too large. Improperly designed, eroded, or even missing restriction orifices used for steam purge or aeration nozzles could cause catalyst attrition. Catalyst attrition is also caused by broken air and stripping steam distributors. Low catalyst level in the regenerator could uncover the diplegs and allow backflow. High catalyst level can prevent the primary cyclones from draining or prevent the trickle valves from operating properly.
Troubleshooting Catalyst Losses
To stop excessive catalyst losses, it should be identified whether the oss is from the reactor or the regenerator. In either case, the following eneral guidelines should be helpful in troubleshooting catalyst losses: « Verify the catalyst bed levels in the stripper and regenerator vessels. • Conduct a single-gauge pressure survey of the reactor-regenerator circuit. Using the results, determine the catalyst density profile. • Plot the physical properties of the equilibrium catalyst. The plotted properties will include particle size distribution and apparent bulk density. The graph confirms any changes in catalyst properties. • Have the lab analyze the "lost" catalyst for particle size distribution. The analysis will provide clues as to the sources and causes of the losses. • Compare the cyclone loading with the design. If the vapor velocity into the reactor cyclones is low, consider adding supplemental steam to the riser. If the mass flow rate is high, consider increasing the feed preheat temperature to reduce catalyst circulation.
48
Fluid Catalytic Cracking Handbook
• Confirm that the restriction orifices used for instrument purges are in proper working condition and that the orifices are not missing. • Consider switching to a harder catalyst. For a short-term solution, if the losses are from the reactor side, consider recycling slurry to the riser. If the catalyst losses are from the regenerator, consider recycling catalyst fines to the unit.
igure 8–4 is a summary of the above discussions.
Nearly every cat cracker experiences some degree of coking/fouling. Coke has been found on the reactor walls, dome, cyclones, overhead apor line, and the slurry bottoms pumparound circuit. Coking and ouling always occur, but they become a problem when they impact hroughput or efficiency.
Evidence of Coking/Fouling Coking/fouling in the reactor and the main column can be detected by: * Cavitation and/or loss of the main column bottoms pumps * Fouling and subsequent loss of heat transfer coefficient in the bottoms pumparound exchange * High pressure drop across the reactor overhead vapor line * Excessive catalyst carryover to the main column
Causes of Coking/Fouling Coke forms in the reactor and main column circuit because of: * * * *
Changes Changes Changes Changes
in in in in
operating parameters catalyst properties feedstock properties mechanical condition of the equipment
Changes in Operating Parameters
The operating conditions of the unit, particularly during startups and eed interruptions, will have a large influence on the formation of coke. oke normally grows wherever there is a cold spot in the reactor ystem. When the temperature of the metal surfaces in the reactor
« O
0
O
5 n
Troubleshooting
249
Problem: Excessive Catalyst loss can cause Unit Shutdown and possible State or Federal Environmental fines
* Increase in Ash and BS&W content of Slurry Bottoms Product * Increase in flue gas Opacity * Loss of Reactor/Regenerator levels * increase in recovery of fines from Electrostatic Precipitator(ESP) * Increase in Regenerator pressure * Decrease in 0–40 microns fractions of E-Cat * Increase in 80+ microns fractions of E-Cat * Change in Catalyst Average Particle Size (APS)
Evidence:
Figure 8-4A.
Changes in Fresh Catalyst properties Fresh Catalyst is "too soft" - has low concentration of I fines
L
• Analyze make-up for PSD & Attrition 1 Plot E-Cat properties - Look for Abnormal Peaks ' Analyze Fines for PSD Consider using a "harder" Catalyst
Troubleshooting catalyst losses.
Changes in Operating conditions j j j i
j
' Increase in Catalyst circulation rate 1 Increase in Cyclone loading 1 Decrease in Cyclone Inlet velocity 1 Increase in Cyclone Gas Outlet velocity 1 Increase in use of Atomizing Steam 1 Increase in Reactor or Regenerator bed levels 1 Water in Steam x ^ Verify Cyclone loading ' Check for any missing RO's Verify accuracy of the Regen. & Rx Catalyst levels: raise or lower Bed level Recycle Slurry to Riser or recovered fines to Regen.
Figure 8-4B.
Changes in Mechanical conditions
• Trickle Valves are either stuck closed or opened 1 Trickle Valves are either warped or eroded ' Trickle Valves have fallen off Holes in the Oiplegs
Cracks in the Plenum 1
Diplegs diameter either "too small" or "too large" Holes in Catalyst Cooler/ Steam Generator "Chunk" of Coke has fallen into Dipleg
Bumping the Regenerator
Troubleshooting catalyst losses.
50
Fluid Catalytic Cracking Handbook
walls and/or the vapor line falls below the dew point of the vapors, ondensation occurs. Condensation and subsequent coke buildup are ue to cooling effects at the surface. A high fractionator bottoms level, a low riser temperature, and a igh residence time in the reactor dome/vapor line are additional perating factors that increase coke buildup. If the main column level ises above the vapor line inlet nozzle, "donut" shaped coke can form t the nozzle entrance. A low reactor temperature may not fully vaporize the feed; unvaorized feed droplets will aggregate to form coke around the feed ozzles on the reactor walls and/or the transfer line. A long residence ime in the reactor and transfer line also accelerate coke buildup. Insufficient bottoms pumparound to the main column heat-transfer one can also form coke.
Changes in Catalyst Properties
Certain catalyst properties appear to increase coke formation. Catalysts with high rare earth content tend to promote hydrogen transfer reacons. Hydrogen transfer reactions are bimolecular reactions that can roduce multi-ring aromatics.
Changes in Feedstock Properties
The quality of the FCC feed also impacts coke buildup in the reactor nternals and vapor line and fouling/coking of the main column circuit. The asphaltene or the resid content of the feed, if not converted in he riser, can contribute to this coking.
Changes in Mechanical Condition of the Equipment
Damaged or partially plugged feed nozzles can contribute to coke ormation due to poor feed atomization. Damaged shed-trays in the bottom section of the main cloumn can ause coke formation due to non-uniform contact between upflowing apors and downflowing liquid.
Troubleshooting Steps
The following are some of the steps that can be taken to minimize oking/fouling:
Troubleshooting
251
« Avoid dead spots. Coke grows wherever there is a cold spot in the system. Use "dry" dome steam to purge hydrocarbons from the stagnant area above the cyclones. Dead spots cause thermal cracking, * Minimize heat losses from the reactor plenum and the transfer line. Heat loss will cause condensation of heavy components of the reaction products. Insulate as much of the system as possible; when insulating flanges, verify that the studs are adequate for the higher temperature. * Improve the feed/catalyst mixing system and maintain a high conversion. A properly designed feed/catalyst injection system, combined with operating at a high conversion, will crack out highboiling feeds that otherwise could be the precursors for the formation of coke. * Follow proper start-up procedures. Introduce feed to the riser only when the reactor system is adequately heated up. Local cold spots cause coke to build up in the reactor cyclones, the plenum chamber, or the vapor line. * Keep the tube velocity in the bottoms pumparound exchanger(s) greater than 7 ft/sec. Putting the bundles in parallel for more heat recovery may lead to low velocity. * Hold the main column shed tray's liquid temperature under 700°F and minimize the level and residence time of the hot liquid. Ensure adequate wash to shed decks to minimize coking in the bottom of the main column. Some paraffinic feeds may require a lower temperature. * Utilize a continuous-cycle oil flush into the inlet of the bottoms exchanger. This keeps the asphaltenes in solution and increases tube velocity. * Verify that no fresh feed is entering the main column. Feed can enter the main column through emergency bypasses or through the feed surge tank vent line.
igure 8-5 is a summary of the above discussions.
LOW REVERSAL
A stable pressure differential must be maintained across the slide alves. The direction of catalyst flow must always be from the regenerator (text continued on page 254)
Figure 8-5A.
r temperature ence Time in the d Main Column ms temperature s Pumparound Rate nger tube wall e J
in the Reactor heat-up olumn Bottoms
Troubleshooting coking/fouling.
Cracked Feedstock
High Mole weight Asphaltene & Resins, precipitate and bind to process equip. High Levels of
Damaged or pa plugged Feed N Loss of the She Feed leaking th Bottoms exchan Feed Diversion
Changes in Mecha Conditions of the Equipment
Higher pressure drop across the Reactor Overhead vapor line
Changes in feedstock Properties
Fouling and loss of Heat Transfer in Bottoms Exchanger
High Level of RareEarth in the Catalyst Low Catalyst Micro Activity Test (MAT)
Cavitation and/or loss of Main Column Bottoms Pumps
in Operating s
:
Unscheduled Unit interruptions loss of profit and higher maintenance costs
* Increase Bottoms traffic * Inject a continuous Cycle Oil flush into inlet Bottoms PA Exchangers * Install duplex filters upstream of Bottoms Pumps * Install high efficiency feed nozzles * Use 1" or larger tube diameter * Keep C7 insolubles in Slurry System less than 5% wt * Use U-tube for Bottoms Exch. * Draw more Bottoms Product * Have a spare Bottoms Exchanger bundle
Figure 8-5B. Troubleshooting coking/fouling.
comes to Coking, Residence Time is e as increasing temp, by 25°F
ber:
erly insulate RX Overhead piping Main Column Inlet nozzle the tube velocity > 7 ft/sec Main Column Bottoms erature < 700°F a "dry" Dome Steam System
Recommendations:
54
Fluid Catalytic Cracking Handbook
ext continued from page 251)
o the reactor and from the reactor-stripper back to the regenerator. A egative differential pressure across the regenerated catalyst slide valve an allow fresh feed and oil-soaked catalyst to backflow from the riser nto the regenerator. This flow reversal can result in uncontrolled urning in the regenerator and potentially damage regenerator internals, osting a refiner several million dollars in production loss and mainenance expense. Similarly, a negative pressure differential across the spent catalyst lide valve can allow hot flue gas to backflow to the reactor and he main fractionator, severely damaging the mechanical integrity of hese vessels. Some of the main causes of loss of pressure differential across the lide valves are as follows: • Loss of the air blower or the wet gas compressor • Presence of water in the feed • High catalyst circulation rates resulting in excessive slide valve opening and low differential • Loss of regenerator or stripper bed levels • Failure of the reactor temperature controller and reactor-stripper level controller • Bypass open around a shutdown valve Troubleshooting flow reversal is outlined in Figure 8-6.
Reversal Prevention Philosophy
The FCC process is very complex and many scenarios can upset perations. If the upset condition is not corrected or controlled, each cenario could lead to a reversal. Table 8-1 contains a cause/effect hutdown matrix indicating scenarios in which a shutdown (reversal) ould take place. In most cases, a unit shutdown is not necessary if dequate warning (low alarms before low/low shutdowns) is provided. he operating staff must be trained to respond to these warnings. The shutdown system will have adequate interlocks to prevent nadvertent trips. The system must include two-out-of-three voting or ackup instruments. The operators must trust the system for it to emain in service.
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Fluid Catalytic Cracking Handbook
Reactor Temperature Controller
Regenerated Catalyst Slide Valve
Low Select
Valve Differential
Pressure Controller
Figure 8-6B. Troubleshooting flow reversal.
Slide valves will have an independent low differential pressure verride controller to prevent the reaction temperature controllers from pening the slide valves to the point where low differential pressure ould allow feed back to the regenerator.
HIGH REGENERATOR TEMPERATURE
The regenerator "dense phase" temperature (see Figure 8-7) can be dversely affected by the feed quality, the condition of the catalyst, he operating variables, and the mechanical conditions of the unit: * Feedstock: A higher fraction of 1,050 °F+ (565°C+) in the feed. For many refiners, the regenerator temperature is the limit on addition of heavy ends to the feed. High-efficiency feed nozzles can be used. Naphtha quench can be added. * Catalyst: A higher level of rare earth or an increase in the matrix content. Switch to a more coke-selective catalyst. * Operating variables: Reduced stripping steam or atomizing steam; higher preheat or riser temperature. Restore the steam flows; this is not the right place to solve a sour water problem.
Close Close
Close
Open Open
Open
Close Close
eed
high n
Close
Open
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Process
Close
Close
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Close
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Close
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Open
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Spent Regan Feed to Slurry HCO Catalyst Emergency Riser Recycle Recycle Slide Valve Steam
Close
Open
Closed
Process
Close
Riser Emergency Steam
Regan Catalyst Slide Valve
low
sure
lide
lide rential
ue tial
ue low sure
Effect:
Table 8-1 A Cause and Effect Shutdown Matrix
ndations
V
Low Stripping Steam Low Dispersion Steam High Preheat Temperature High Reactor Temperature Low Cat/Oil Ratio
Operating Conditions
Install high efficiency Feed Nozzles Lower Preheat Temperature Inject Naphtha Quench to Riser increase Stripping and Dispersion Steam Switch to a Coke Selective Catalyst
• High Level of Rare Earth • High Level of MaWx ) Activity v J
Catalyst
Mechanical Conditions • Damaged Strippin Steam distributor • Damaged Feed Nozzles • Damaged Air or Catalyst Distribut
Figure 8-7, Troubleshooting high regenerator temperature.
•High Fraction of 1050°F+ Material
'Feedstock ^ '
Low Catalyst/Oil ratio Poor Yields Loss of Catalyst Activity Equipment Wear V ,
Troubleshooting
25i
* Mechanical conditions: Damaged stripping steam distributor and/ or damaged feed nozzles; stripper internal damage. Increase carrying gas injection below the feed nozzle. Increase stripping and dispersion steam. Correct the problem during the next turnaround,
NCREASE IN AFTERBURN
It is important that combustion of the coke in the spent catalyst ccur in the dense bed of catalyst. Without the catalyst bed to absorb his heat of combustion, the dilute phase and flue gas temperatures ncrease rapidly. This phenomenon is called afterburning. It is critical hat spent catalyst and combustion/lift air are being introduced into he regenerator as evenly as possible across the catalyst bed. It is also mportant to note that vertical mixing is much faster than lateral mixing. The magnitude of afterburn in the regenerator largely depends on he operating conditions of the unit and the effectiveness of the contact etween the combustion air and the spent catalyst. The geometry of he regenerator and the distribution of the spent catalyst also impact he level of afterburning. Generally speaking, regenerators operating n partial combustion do not experience the same level of afterburning s compared with full-burn regenerators. * Operating variables: Low catalyst residence time in the regenerator, low regenerator dense bed temperature, low regenerator pressure, insufficient stripping steam, and insufficient CO promoter, Increase regenerator bed level. Increase promoter flow. Increase feed preheat, increase reactor temperature, or change the catalyst formulation to increase the regenerator temperature. Regenerator pressure can be increased to push the burn down into the catalyst bed. * Mechanical conditions: Damaged or improperly designed air and/ or spent catalyst distributors. Partially damaged stripping steam distributors that, depending on the style of spent catalyst distributor, could allow flashing of the entrained hydrocarbon vapors into the regenerator dilute phase. Corrective actions include: — Design of mechanically robust spent catalyst and air distributors. — Selective plugging/addition of air distributor jets to match the spent catalyst distribution — Selective rotation of the regenerator primary cyclone dipleg termination
60
Fluid Catalytic Cracking Handbook
HYDROGEN BLISTERING
Hydrogen blistering is the result of cyanide-induced corrosion in he light ends section of the unit, usually the top of the absorber. The implified chemical reactions are as follows: Fe + 2HS~ -» FeS + S~2 + 2H FeS + 6CN- -> F e C N - 4 + S-2
Normally, a layer of FeS scale, produced in the first reaction, protects he interior of the pipe or vessel. However, in the second reaction, yanide removes the FeS protective scale exposing more free iron to eact with H2S and releasing more hydrogen. Without protection, this ycle continues until blistering, cracking, and eventual total corrosion of equipment occurs.
Causes of Higher Levels of Cyanide The primary sources of higher cyanide production are: * * * »
Higher levels of nitrogen in the feed Higher reactor temperature Operating in partial combustion mode Higher matrix activity of the catalyst
Steps to Control Hydrogen Blistering
The best way to minimize hydrogen blistering is to control the orrosion rate. Both corrosion and hydrogen blistering rates can be ignificantly reduced by implementing the following steps (also see Figure 8-9): 1 . Install a water-wash system for dilution and removal of cyanide from the unit. Cascading wash water from the high-pressure zone back to the main column overhead or to the first-stage wet gas compressor outlet is attractive, but it is better to use overhead water and pump it from low pressure to high pressure. 2. Add polysulfide solution to neutralize the cyanide. Air can be injected into the main column overhead to make polysulfide,
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•A lower Main Column top temperature •Partial Combustion •Higher Reactor Temp.
•Increase in Matrix Activity
crease in ed Nitrogen
•Lack of stress relief •Poor material & workmanship • Inadequate Resid Time to Separate G Lean OH, & Water
/Mechanical Condit
Figure 8-9. Troubleshooting hydrogen blistering.
nitor Ammonia & Chloride in the Overhead Water, keep Ammonia Sulfide <5,000 ppm Steam Condensate as Water Wash at a Rate of 1-2 gpm/1000 bbl of Fresh Feed Ammonium Polysulfide Solution (especially if HCN > 25 ppm) to 10–20 ppm Residual HCN ke sure the Wash Water is injected uniformly into the Gas Stream "Feed Forward" Water Wash Scheme instead of "Reverse Cascade" tall & monitor Hydrogen probes in the key areas ct a Dehazer Additive into the High Pressure Separator (HPS) tall Coalescer in the Lean Oil and Sponge Oil Streams ss relieve carbon steel
Operating Conditions
Catalyst Properties
Hydrogen Blistering, Cyanide Attack
Feed Properties
blem:
Troubleshooting
2S3
3. Install and monitor corrosion and hydrogen probes in key areas, 4. Inject a filming amine to provide a protective barrier that will prevent cyanide from contacting the iron sulfide layer.
Power recovery trains recover energy from the flue gas. The FCC tarts to resemble a large jet engine; air is compressed into a comustion zone and expanded across a turbine. Power recovery increases he efficiency of the unit but adds one more mechanical device to an lready long list. Since they are too big to bypass, power trains need o be as reliable as the rest of the unit. The main concerns in the design and operation of a power recovery ystem are catalyst fines and temperature. Catalyst fines will lead to erious blade wear, deposits, power loss, and rotor vibration. Deposit ccurs most frequently where flue gas velocities are at maximum evels, such as blade outer diameter.
Causes of Blade Wear, Power Loss, and Rotor Vibration 1. Increase in catalyst loading to the regenerator cyclones and thirdstage separator 2. Increase in flue gas rate 3. Increase in fresh catalyst addition rate 4. Catalyst that is too hard or too soft 5. Sodium and vanadium in the catalyst, resulting in the formation of low-melting eutectic, which makes the catalyst very sticky even at high temperatures
Troubleshooting Steps 1. Regular monitoring of rotor blade conditions by visual inspection, photographs, and/or video recording. A port is usually installed for this. 2. Continuous monitoring of rotor casing vibration, bearing temperatures, and the expander inlet/outlet temperatures. Problems can be either instantaneous or slow-growing. Instantaneous problems occur during startup, upset, and shutdown, and are easy to note. Slow-growing problems can creep up and are almost
64
Fluid Catalytic Cracking Handbook
invisible while everything is running well. Compare the readings month-to-month to spot trends. 3. Continuous monitoring of the third-stage separator performance. If catalyst is showing up downstream, consider using more than the "standard" 3% flue gas underflow. The blowcase needs more attention than it usually gets. 4. On-line cleaning—inject walnut hulls into the inlet of the expander weekly. 5. Thermal shocking—reduce feed in 20% increments while maintaining maximum air rate to the regenerator. Cool the expander inlet temperature to around 1,000°F (540°C) and hold for at least one hour. This is not a procedure that the expander vendor supports, but it is practiced by many refiners. Figure 8-10 provides an outline of the above discussion.
The performance of the unit is a function of feed quality, catalyst properties, operating variables, and the mechanical condition of the unit. The indicators used to measure the unit's performance are: * * • *
Conversion Dry gas yield Gasoline quality Light olefin yield
Observing a Low Conversion
"True" conversion is affected by feed quality, catalyst, operating variables, and mechanical conditions (Figure 8–11A).
Feedstock Quality The feed properties that lower conversion are: • Increase in residue (1,050°F+) (565°C+) content • Increase in feed impurities such as nickel, vanadium, sodium, or nitrogen « Increase in naphthene and aromatic fractions of the feed
n
Troubleshooting
Problem:
Evidence:
265
* Loss of Revenue * Unscheduled Shutdown * Personal/Equipment Hazards
. * * * *
High level of vibration Loss or gain of efficiency Changes in Control Valve position Excessive blade wear Catalyst deposits on shroud Sulfide/Chloride corrosion on blades disc roots
Figure 8–1OA. Troubleshooting the hot gas expander.
Feedstock
* Increase in Catalyst loading to Cyclones »Increase in Catalyst addition •Increase in Flue Gas flow
O
O * On-line cleaning * Regular injection of "abrasive" walnut hulls * Check stm cond. and flow •Contact OEM
•Too "soft" or too "hard" ' Excessive 0-40 Micron tract 'Wet Steam
'Thermal shocking of the Expander If "hulls" fail 'Contact OEM UseSpht. Stmfor blanket cooling
High levels of Sodium and Vanadium form eutectics, making the Catalyst sticky
' Regular visual, photographic and video monitoring of the Rotor Blade conditions, esp. after Upset ' Check flows & press in Disc cavity
* Regular monit. of casing vib., bearing temp. and Expander Inlet/Outlet temps. * Instrument . underflow line
Figure 8-1 OB. Troubleshooting the hot gas expander.
atalyst Properties
Both the fresh and E-cat properties have a large influence on onversion. The prime properties are: • Decrease in microactivity (MAT) « Decrease in the surface area * Increase in the CRC
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I.ower Gasoline Yield
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Changes in Mechanical Conditio
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Lower Gasoline Octa Figure 8-11D
Reduce Gasoline Olefins Figure 8-11E
High Dry Gas Yield
Loss of Revenue Off-Spec Products
Figure 8-11 A. Troubleshooting desired product quantity and quality.
w Fee^ erties k for changes in gen, Nl, V, °F API gravity, K or, Rl & 650°F
e Contaminants s Paraffins e Aromatics of run for Feed
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^
er Conversion
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Troubleshooting
267
The decreases in microactivity and surface area are strong functions f thermal deactivation in the regenerator and the presence of metals n the feed.
Operating Variables The following operating parameters lower conversion: • • • •
Decrease Decrease Decrease Decrease
in in in in
the the the the
reactor temperature catalyst-to-oil ratio atomizing steam fresh catalyst addition rate
Mechanical Conditions
Damaged or plugged feed nozzle(s) and/or damaged stripping steam istributor(s) are the common causes of mechanical failures that ffect "true" conversion. Note that the "apparent" conversion, as iscussed in Chapter 5, is affected by the distillation cut point and main column operations.
roubleshooting Steps • Trend the feedstock properties; look for changes in the K factor, 1,050°F+ (565°C+), aniline point, refractive index, and °API gravity. The feed endpoint may have been increased to fill the unit. The conversion penalty may be a small price to pay for the increased capacity, but the penalty can be minimized. Verify that the refinery LP reflects current data on yields and product quality. • Plot properties of the fresh and equilibrium catalysts; ensure that the catalyst vendor is meeting the agreed quality control specifications. Verify that the catalyst vendor has the latest data on feed properties, unit condition, and target products. Verify the fresh makeup rate. Check for recent temperature excursions in the regenerator or afterburning problems. • Trend the reactor temperature, cat-to-oil ratio, and atomizing steam rate. Verify the accuracy of the reactor temperature thermocouple and atomizing steam flow meter. • Perform a single-gauge pressure survey around the feed nozzles. Calculate the hydrogen content of the spent catalyst. Conduct a
268
Fluid Catalytic Cracking Handbook
gamma ray scan test to verify the mechanical condition of the stripping steam distributor.
Observing a High Dry Gas Yield
Dry gas yield is affected by everything that affects conversion Figure 8-11B). Changes to increase conversion can increase the dry gas yield. High gas yield shows up as higher speed on the compressor (if centrifugal). In many cases, lower molecular weight (due to higher hydrogen content) can have the same effect.
Feedstock Quality The feed parameters that increase the dry gas yield are: * Increase in nickel and vanadium content * Increase in naphthene, olefin, and aromatic concentration, which is indicated by an increase in the refractive index and decreases in aniline point and K factor
Catalyst Properties The E-cat properties that increase dry gas yield are: * Increase in the level of nickel, vanadium, and sodium * Decrease in E-cat activity, surface area, fresh catalyst activity, and rare earth content * Increase in the gas and coke factors of the E-cat
Operating Variables Operating parameters that increase dry gas yield are: * Increase in the reactor temperature * Increase in the regenerator temperature « Decrease in the atomizing steam * Increase in slurry or HCO recycle
Mechanical Conditions Mechanical conditions that can increase dry gas yield are: » A failing reactor temperature thermocouple * Partially plugged or damaged feed nozzles
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roubleshooting Steps The following steps should be carried out: • Track changes in feed metals content, trend the aniline point, and refractive index. • Trend changes in catalyst activity, surface area, rare earth, and metals content. Consider adding/increasing metals inhibitor. • Trend changes in the molecular weight of the gas at the firststage suction. Verify that overhead cooling and wash systems are in order. • Verify the position of the wet gas compressor spillback. Determine if the compressor turbine needs water washing. Trend the level of inert gases in the dry gas. « Calibrate the reactor temperature controller. Conduct a pressure survey around the feed nozzle piping to verify its mechanical integrity. • If no significant problems are found other than feedstock changes, verify that the refinery LP team has current data on unit yields and product quality with this feedstock. The result of troubleshooting may be that increasing dry gas may be a necessary price for changes in the feed.
Observing a Lower Gasoline Yield
The FCC "true" gasoline yield largely depends on changes in feed uality, catalyst properties, operating variables, and mechanical conitions (Figure 8–11C).
eedstock Quality
Paraffinic feedstocks produce the most gasoline yield (but the owest octane). The common indicators of any increase in feed parffinicity are: • • • •
Increase in the K factor Increase in the aniline point Increase in the nickel-to-vanadium ratio Decrease in the fraction of "cracked" material
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Catalyst Properties The fresh catalyst properties that increase gasoline yield are: « Increase in the rare earth content • Increase in the zeolite content • Increase in the unit cell size
Operating Conditions The operating parameters that increase gasoline yield are: • Decrease in the feed preheat temperature and subsequent increase in the catalyst-to-oil ratio • Decrease in the carbon content of the E-cat if the carbon is greater than 0.1 wt% • Increase in the reactor temperature if overcracking is not occurring • Decrease in the ZSM-5 additive—a shift in FCC gasoline at the expense of LPG
Mechanical Conditions • Deterioration of the feed nozzles » Erroneous stripper level
Troubleshooting Steps • Trend the feed °API gravity, K factor, and aniline point. Verify any changes in paraffin content of the feed. • Plot the catalyst's unit cell size, rare earth, and activity. Check if there is any fluctuation in catalyst properties. • Verify the gasoline end point, vapor pressure, and LCO distillation to ensure minimum undercutting of LCO.
Observing a Low Gasoline Octane
In general, any parameter that increases the gasoline yield will also ecrease its octane. One reason is that the high-octane components in he gasoline tend to be denser than the low-octane components. Therefore, any change that produces more gasoline will result in a ower octane. Again, feedstock, catalyst, operating variables, and mechanical conditions play important roles in affecting gasoline octane Figure 8-1 ID).
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Solutions:
Higher Feed K Factor Increase in the Catalyst Rare Earth content Decrease in the Catalyst Matrix Activity Larger Zeolite Unit Cell Size Add ZSM-5 Additive Higher Cat/Oil ratio Higher Mix Zone temperature
Split Feed injection Riser Quench Figure 8-11E. Troubleshooting high gasoline olefins.
eedstock Quality Gasoline octane is increased by: • Increases in the refractive index • Decrease in the K factor and aniline point • Increase in the bromine number
Catalyst Properties
The fresh catalyst's chemical properties also influence the FCC asoline octane. Gasoline octane is increased by: • Decrease in rare earth and unit cell size • Decrease in sodium content • Increase in matrix activity
Operating Conditions
A number of operating variables can change the octane value. The actors that increase octane are: • Increase in the reactor temperature. In general, one research octane number increase per 17°F (10°C) increase in the reactor temperature. • Decrease in the catalyst-to-oil ratio (by increasing thermal reactions).
Troubleshooting
• • • • •
275
Increase in coke content of the regenerated catalyst. Increase in the regenerator temperature. Increase in the naphtha quench or HCO recycle. Decrease in the gasoline end point. Decrease in the gasoline vapor pressure.
Mechanical Conditions
The main mechanical conditions that affect octane are the type and ondition of the feed nozzles. Low-efficiency feed nozzles actually ncrease the gasoline octane due to promotion of thermal reactions in he mix zone. High-efficiency feed nozzles improve feed/catalyst mixing and increase the gasoline yield, but decrease gasoline octane.
roubleshooting Steps • Plot the feed refractive index, °API gravity, and aniline point. Determine any shift in the amount of cracked gas oil in the feed. • Track the unit cell size, matrix activity, and rare earth content of the catalyst. • Determine if coke on the catalyst has changed. « Verify accuracy of the reactor temperature. • Check for changes in the gasoline end point and vapor pressure. • Check the conditions of the feed nozzles and amount of atomizing steam.
Gasoline Vapor Pressure/Light Olefln Yield
Reformulated gasoline specifications require lower vapor pressure n the blended gasoline. It also requires maximum feed to the alkylaon unit. This puts more pressure on the gas plant, particularly the ebutanizer. Floating the tower pressure is often the best way to meet oth constraints.
This chapter highlights the common problems, symptoms, and robable causes that may be encountered in troubleshooting FCC units. n addition, a systematic approach is outlined to provide solutions and orrective action. The suggested solutions are necessarily generic but pply to a wide variety of units.
CHAPTER 9
Debottlertecking and Optimization
Troubleshooting, optimization, and debottlenecking are three steps n a continuous process. There is some overlap and gray area among them. Troubleshooting refers to the solution of short-term problems. The ssignment is usually initiated by operations or maintenance. The olution usually involves something that can be done online. Troublehooting was discussed in Chapter 8. Optimization refers to maximizing feed rate and/or conversion with he existing equipment while reaching as many constraints as possible. can be the response to changes in the feed quality, ambient conitions, or the market demands. Although it is not discussed separately ere, it is the incentive for most debottlenecking projects. Debottlenecking often refers to hardware changes, small or large. It is irected at the bottlenecks identified during optimization. It includes rojects that cannot be completed online, such as installing new nternals in a vessel. Debottlenecking is the main focus of this chapter.
NTRODUCTION
Most FCC units are big profit makers. Therefore, they are operated o several constraints. Debottlenecking is the effort to locate and vercome these constraints. The profitability of an FCC operation is maximized when the unit is "pushed" simultaneously against multiple onstraints. Debottlenecking means finding the constraint or combinaon of constraints that cost the refinery lost opportunities and arriving t the right fix. A properly configured advanced process control (APC) system could llow for on-line, continuous optimal unit operation and push the FCC perations to multiple constraints simultaneously. The main purpose of debottlenecking is to increase the refinery's rofit margin. In the FCC, this usually means: 276
Debottlenecking and Optimization
27?
• Raising the feed rate » Processing lower quality feedstocks • Reducing dry gas and coke yields, therefore, increasing total liquid products
As with troubleshooting, a proper debottlenecking exercise must onsider the effects of feedstock, catalyst, operating conditions, mechanical ardware, environmental issues, and the ability of the rest of the efinery to handle the additional feed/product rates and quality.
APPROACH TO DEBOTTLENECKING
Debottlenecking requires a comprehensive test run to determine the peration's present status. Elements of a test run include: « • • • • • •
Overall and component material balance Reactor/regenerator heat balance Hydrogen balance Sulfur balance Reactor/regenerator pressure survey Utility balance Evaluation of the interaction among feed quality, catalyst properties, and operating conditions • Main fractionator and gas plant modeling
If the object of debottlenecking is to run heavier feeds, multiple test uns may be needed with heavy feed added in stages. The next step is to identify the incremental value of: » • • •
Fresh feed rate Each FCC product Octane and cetane numbers Other product quality issues (sulfur, slurry ash level, etc.)
With this information, the constraints on operation can be identified nd the value of addressing them can be evaluated.
mproving FCC Profitability through Proven Technologies
Once the performance of the FCC unit is optimized through the use f new catalyst and operating practices, the unit's profitability can be urther improved by installing proven hardware technologies. The urpose of these technology upgrades is to enhance product selectivity
78
Fluid Catalytic Cracking Handbook
nd unit reliability. Since the 1980s, mechanical upgrade of FCC units as proceeded at a fast pace. New feed/catalyst injection systems and limination of post riser reactions have been the forefront of these mechanical upgrades.
Apparent Operating Constraints
The unit operating philosophy and its apparent operating limits often ictate unit constraints. For example, limitations on the main column ottoms temperature, the flue gas excess oxygen, and the slide valve elta P often constrain the unit feed rate and/or conversion. Unfortunately, ome of these limits may no longer be applicable and should be rexamined. Some of them may have resulted from one bad experience nd should not have become part of the operating procedure.
Debottlenecking
The remainder of this chapter contains suggested ways of addressing onstraints in the following areas of the FCC unit: • Feed preheat section • Reactor-regenerator section • Main fractionator and gas plant
Included are discussions regarding the feed/catalyst system, instrumentation, and off-sites. It should be noted that a change in one system sually affects others.
eed Circuit Hydraulics
Figure 1-5 shows a typical feed preheat configuration. A hydraulic mitation usually manifests itself when increasing fresh feed rate and/ r installing high efficiency feed injection nozzles.
Typical Feed Preheat Section
The hydraulic pinch points in the feed preheat system are identified ith a single-gauge pressure survey. The bottlenecks are often related to: • Feed pumps • Fresh feed control valve
Debottlenecking and Optimization
* • » »
279
Piping Preheat exchangers Preheat furnace Feed nozzles
The feed pump will be re-rated for the new conditions. With higher iscosity and higher gravity, the pump driver may need work. If he system is not adequate, heavier feed can be piped through a eparate circuit in parallel with the existing circuit, preferably on flow atio control. If the pump is the bottleneck, before changing it, consider: * Installing a larger impeller. • (Turbine:) Increasing turbine speed. Evaluate the steam level and consider adding an exhaust condenser. * (Motor:) Changing to a variable speed drive (VSD). VSD's make startup easier and most can support 10% overspeed. • Changing the driver. * Adding pumps in parallel. • Adding a booster pump downstream.
As shown in Example 9–1, increasing the pump impeller size from 3 inches to 13.5 inches increases the flow by 3.8%, discharge pressure y 7.8%, and horsepower by 12%. Increasing the turbine speed from ,300 rpm to 3,400 rpm increases the flow by 3%, the discharge ressure by 6.1%, and the horsepower by 9.4%. Example 9-1 Q,, h,, bhp,, d,, n, = Initial Capacity, head, brake horsepower, diameter, and speed Q2, h2, bhp2, d,,, n2 = New Capacity, head, brake horsepower, diameter, and speed
Diameter Change Only Speed Change Only Diameter & Speed Change
Q2 = Q^d,)
2
Q2 = Q^yn,) 2
= h^cyd,) _
hp2 = bhPl(d2/d,)3
H2 = h^/n,)
Q2 = Q,(d2/d, x n/n,) 2
Bhp2 = bhp^n/n,) 3
h2= h^d/d, x n2/n})2 bhp2= bhp.Cd/d, x n/n,) 3
80
Fluid Catalytic Cracking Handbook
Given: d1 =13 in. d? = 13.5 in.
n} = 3,300 rpm n2 = 3,400 rpm
Flow Increase 3.8% (impeller only) 3.0% (speed only) 7% (impeller and speed) Head Increase 7.8% (impeller only) 6.1% (speed only) 14.5% (impeller and speed) Horsepower Increase 12.0% (impeller only) 9.4% (speed only) 22.5% (impeller and speed)
New internals in the control valve or a larger control valve can be he cheapest option if no piping needs to be changed. If the pressure drop in the feed piping is excessive, consider increasng the line size or installing a parallel line. Check the existing flange atings if any changes are made in the pump or piping, or if the emperature is changed significantly. If diluent is being added to the feed, evaluate the optimum point or minimum pressure drop and maximum heat recovery. The preheat furnace can be a bottleneck. The first consideration is hat it may not be needed in the new operation. With the increase in he FCC rate, the pressure drop will increase. Consider: • Using the furnace bypass. • Verifying the position of the inlet balancing valves. When balancing a heater, operators tend to pinch the valves. At least one of the valves should be wide open. • Decoking the heater. Consider hydraulic cleaning. • Increase the number of tube passes. Changing from a two-pass to a four-pass arrangement can reduce the pressure drop by over 75% (see Example 9-2). • Adding diluent downstream.
Debottlenecking and Optimization Example 9-2 Changing Piping in Furnace from Two-Pass to Four-Pass Case I: Two-Pass Furnace 50,000 BPD total charge (25,000 BPD to each pass) °API gravity of feed = 25 Furnace outlet temp. = 500°F Furnace tube diameter (I.D.) = 4.5 in. AP 100 = 0.0216 x
Where:
Pi(K) f p Q d
= Pressure drop (psi) per = Friction factor = 0.017 = Flowing density = 47.4 = Actual flow rate = 864 = Tube inside diameter =
100 feet of pipe lb/ft3 GPM 4.5 in.
AP!00 = 7.0 psi
Assuming a total 700 ft of equivalent pipe in the furnace, the total ressure drop is 49 psi Case II: Switching to Four-Pass AP 1(X! = 1.9 psi
Assuming a total 500 ft of equivalent pipe in the furnace, the total ressure drop is 9.4 psi
aving in pressure drop = 49.0 - 9.5 = 39.5 psi or an 81% reduction
This section addresses the following:
* Mechanical limitations • Riser termination device « Feed and catalyst injection system
281
82
Fluid Catalytic Cracking Handbook
* Spent catalyst stripper « Slide valves • Regeneration
Mechanical Limitations
Mechanical limitations include the design temperature and pressure f the reactor and the regenerator.
Dehottlenecking the Reactor Pressure/Temperature
The FCC reactor pressure is usually controlled at the suction of the wet gas compressor. The reactor pressure is the wet gas compressor sucion pressure plus pressure drop through the main fractionator system. Reactor temperature is usually directly controlled by adjusting the lide valve openings or changing the pressure differential between the egenerator and reactor. Mechanical design conditions of the reactor ystems can limit operating at more severe conditions. To debottleneck hese limitations: * The reactor vessel can be rerated based on actual metal thickness and corrosion history at the new operating temperature. • An external cyclone can be used to unload the vessel. • Internal lining can be added. * A reactor quench system can be used. • Split feed injection can be considered. * The riser and the reactor can be replaced with a cold-wall design.
Debottlenecking the Regenerator Pressure/Temperature
The regenerator is already a cold-wall vessel; re-rating is not often ractical. High regenerator temperature typically requires installing ither catalyst coolers, operating with partial combustion, or injecting quench stream into the riser.
Riser Termination Device (RTD)
Post-riser hydrocarbon residence time leads to thermal cracking and on-selective catalytic reactions. These reactions lead to degradation f valuable products, producing dry-gas and coke at the expense of
Debottlenecking and Optimization
283
asoline and LPG. Improvements in FCC catalyst have eliminated any ncentive for these reactions. Thermal reactions are a function of time and temperature; yields are roportional to (time)*(exp[~E/RT1). Figure 9-1 shows the typical effects f vapor residence time and temperature on dilute phase cracking. or example, at 5 seconds residence time, the dry-gas yield increases % when the reactor temperature increases from 960°F to 980°F. ncreasing the residence time to 10 seconds increases the dry gas yield nother 8%. Since the mid-1980s, FCC technology licensors and a number of il companies have employed a number of RTD's to reduce nonelective post-riser cracking reactions. Two general approaches have een used to reduce post riser cracking. The most widely used approach direct connection of the cyclones to the riser and on to the reactor apor line. The second approach is quenching the reactor vapors ownstream of the riser-cyclones (rough-cut cyclones). RTD's separate the catalyst and the oil vapor immediately at the end f the riser. The cyclone vapor usually discharges directly to the econd-stage cyclones and then to the reactor vapor line. The catalyst directly discharged into the stripper. The "reactor" is simply a vessel or holding the cyclones. Technologies are offered by:
0
10
20
30
40
50
Residence time, sec. Figure 9-1.
Liquid loss from thermal cracking.
60
84
• • » • •
Fluid Catalytic Cracking Handbook
ABB Lummus Global (Lummus) Exxon Research & Engineering (ER&E) Kellogg Brown & Root (KBR)Mobil Oil UOP Stone & Webster Engineering Corporation (SWEC)/IFP
BB Lummus' DCC Features
ABB Lummus's RTD consists of a two-stage reactor cyclone system see Figure 9-2). The riser cyclones (the first stage) are hard-piped to he riser. Attached to the end of each riser cyclone dipleg is a conentional trickle valve as shown in Figure 9-3. Each trickle valve has small opening to prevent catalyst defluidization, which can be a roblem, especially during start-ups. At the vapor outlet of the first-stage cyclones, an opening allows ntry of stripping steam/vapors and reactor dome steam. This opening s sized to allow the second stage cyclones to be operated at a negative ressure relevant to the reactor housing pressure. Attached to the end of the upper reactor cyclone diplegs are horizontal, ounter-weighted flapper valves (Figure 9-4). These valves provide a ght seal between discharging catalyst and upflowing vapors in the eactor housing. ER&E's RTD offering is principally similar to the Lummus design. n the ER&E design, the riser cyclones are not hard-piped to the riser. However, the outlet of the riser-cyclones are directly connected to the nlet of the upper cyclones.
BR Closed Cyclone System
In the KBR system, as with the ABB Lummus design, the riser yclones are hard-piped to the riser. The diplegs of both the riser yclone and the upper reactor cyclone are often sealed with atalyst. This minimizes the carry-under of reactor vapors into the eactor housing and maximizes the collection efficiency of the ser cyclones. No trickle or flapper valves are used on the first stage. The riser yclone diplegs terminate with a splash plate (Figure 9-4A). The upper eactor cyclone diplegs use conventional trickle valves. Sealing the pper reactor cyclone diplegs with about two feet of catalyst provides
Debottlenecking and Optimization
285
Stripper Gas
(D
be
P1>P2>P3
Figure 9-2. Lummus direct-coupled cyclone design.
nsurance in case the trickle valves stick open. In this design, the riser yclones operate at a positive pressure and sealing the diplegs minimizes carry-under of reactor vapors into the reactor housing. The catalyst must be fluidized to provide an effective seal for the iplegs. Fluidization is critical; without it, the diplegs cannot discharge he catalyst and will plug, with possible massive carry-over to the main actionator. To ensure this uniform fluidization, the system uses an dditional steam distributor.
86
Fluid Catalytic Cracking Handbook
Cyclone Dipleg*
Restraint Figure 9–3. Typical trickle valve.
Cyclone Dipleg Pivot
Counterweight
\ Figure 94.
Typical flapper valve.
In KBR closed cyclone technology, each set of riser and upper eactor cyclones is connected via the use of a "slip joint" conduit. The tripper steam and hydrocarbons, as well as dome steam, exit he reactor housing by entering through this conduit as shown in Figure 9-5.
Debottlenecking and Optimization
287
Cyclone Dipleg
05
16 o Braces (as required)
Splash Plate Figure 9-4A.
Typical splash plate.
UOP VSS System
UOP's current RTD offering is the vortex separation system (VSS), s shown in Figure 9-6. VSS is for FCC units having an internal riser nd a similar design (VDS) is for external risers. The catalyst-vapor mixture travels up the riser through the chamber and exits through everal arms. These arms generate a centrifugal flow pattern that eparates the catalyst from the vapor inside the chamber. The catalyst ccumulates in a dense phase at the base of chamber, where it is "pre~ tripped" prior to flowing into the reactor stripper. The stripped ydrocarbon vapors are fully contained in the chamber and exit with he rest of the riser effluent vapors to the secondary cyclones. The reactor vapors leave the VSS through an outlet pipe. Secondary yclones are directly connected to this outlet pipe through an expansion
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Fluid Catalytic Cracking Handbook
Product Dome
Riser
Splash Plate
Catalst Level
Figure 9-5. KBR closed cyclone system.
oint. The VSS outlet pipe contains several vent pipes in which the eactor dome steam and a portion of the stripping steam/hydroarbon vapors leave the reactor through these vent pipes.
tone & Webster Engineering Corporation (SWEC)
SWEC offers a reactor quench system rather than a closed cyclone ystem. Their typical RTD is an external, rough-cut cyclone (see Figure -7). The vapors from the rough-cut cyclone enter the reactor vessel.
Debottlenecking and Optimization
Expansion Joint
Flapper Valve
Spent Catalyst to Stripper
Figure 9-6.
UOP vortex separation system.
289
90
Fluid Catalytic Cracking Handbook
LCO Quench
Pre-Stripping Steam
Figure 9-7.
To Catalyst Stripper
SWEC external cyclone with quench.
Debottlenecking and Optimization
291
The recovered catalyst enters the reactor via an external dipleg. Aside rom external rough-cut cyclones, SWEC also offers riser-cyclones, eferred to as LD2 (Linear Disengaging Device), intended to separate atalyst from reactor vapors quicker than conventional cyclones (see Figure 9–7A). LCO quench is injected into the vapors leaving the rough-cut yclone. The reactor temperature is usually reduced to 930°F where hermal cracking is minimal. This design often requires a pre-stripping ing at the outlet of each dipleg to ensure steady catalyst discharge rom the external dipleg.
To Main Column
Upper Cyclones Vapor (Catalyst)
Stripper
igure 9-7A. Webster).
SWEC LD2 riser termination device (courtesy of Stone and
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Fluid Catalytic Cracking Handbook
Feed Nozzles Important features of a feed injection system include: • • • • • • •
Fine atomization of feed High velocity coverage of riser cross-section Intimate mixing of catalyst and oil Rapid heat transfer from catalyst to oil Instantaneous vaporization of feed Minimum catalyst back-mixing Maximum catalytic reactions while minimizing thermal reactions
A good feed injection system will produce: • Small droplet size • Efficient mixing of oil and catalyst • Complete riser coverage
The feed injection system has come a long way. The early designs eatured open pipes with no consideration for feed vaporization or atalyst/vapor mixing. Currently, FCC technology licensors offer many versions of feed injection systems. Figure 9-8 is a typical modern feed nozzle. In general, these nozzles incorporate some of the following design features:
Oil Inlet
Diverging Dual Slot
Target Bolt
Figure 9-8. SWEC feed nozzle.
Debottlenecking and Optimization
2§3
* Steam is used to disperse and atomize the oil/residue feed • The spray pattern of the oil/steam leaving the nozzle tips tends to be flat (fan spray) • The assembly includes multiple nozzles in a radial pattern * The nozzles are designed for a "medium" oil-side pressure drop, generally in the order of 50 psi
Some of the general criteria for choosing feed injection technology nclude: • Total installed cost * Dispersion steam and/or lift steam/gas requirements, including flow rate, temperature, and pressure * Oil pressure requirement • Proven track-record of operational reliability
The choice of the feed injection system should be based on the endor's experience in similar units with similar feeds and on his yield rojection and/or performance guarantee. However, it may be difficult o substantiate the guarantee when other changes are being made in he unit.
Spent Catalyst Stripper
Spent catalyst from the reactor/cyclones discharges into the stripper. Stripping steam displaces hydrocarbon vapors entrained with the atalyst and removes volatile hydrocarbons from the catalyst. As part of optimizing the unit, the stripping steam rate should be djusted up or down by 5%. The regenerator temperature and/or CO2/ CO ratio will be the main indicator of insufficient stripping. The test ends when there is no significant response in the regenerator temperature. In the past several years, more attentions have been given to improvng mechanical performance of the reactor stripper. Proprietary stripper designs are being offered by the FCC technology licensers in attempts o improve the catalyst/steam contact.
Debottlenecking Catalyst Circulation
Any attempt to increase the unit feed rate will generally require an ncrease in catalyst circulation. The unit pressure balance and catalyst irculation were covered in Chapter 8.
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Fluid Catalytic Cracking Handbook
The following should be considered when debottlenecking: * * * *
Differential pressure alarm/shutdown Increasing slide valve size Standpipes Catalyst selection
Differential Pressure Alarm/Shutdown
Differential pressure shutdowns are a critical part of the unit's safety ystem. No attempt to lower the setting on the shutdown should be made without adequate consideration. On the other hand, pressure is ost across the slide valves and costs money. Multiple, independent differential pressure alarm/shutdown switches can be installed with two out of three voting. This can satisfy the afety requirement, increase the comfort factor, and gain valuable pressure drop. Radial feed nozzles also minimize the possibility of a reversal. New valve actuators can operate more quickly and reliably, also increasing he safety factor. The test run may indicate that the slide valve is open too far. Most operators prefer to keep the valve in the 40% to 60% range. They lose major comfort zone if the valves open more than this. A larger valve or larger port can be installed in the existing valve.
Standpipes
If the unit pressure balance indicates that either the pressure gain n the Standpipes is inadequate or the delta P across the slide valves s erratic, standpipe aeration and instrumentation should be examined. Redesigning the aeration systems or replacing the Standpipes can gain valuable pressure drop. Proper instrumentation can include independent eration flow to each tap, flow indicators/controllers on each, and ifferential pressure indicators between the taps. Beyond the Standpipes, the available delta P across the valve is ffected by the pressure drop in other circuits. For the regenerated atalyst slide valve, downstream pressure is affected by: * Feed injection system » Riser
Debottleneeking and Optimization
2S5
• Reactor cyclones • Reactor vapor line » Main fractionator and overhead system
The regenerated catalyst slide valve upstream pressure is increased by: • Increasing the regenerator bed level « Increasing the regenerator pressure • Increasing the 0 to 40 micron content of the circulating catalyst
Debottlenecking Combustion Air
Many FCC units are constrained by the air blower, particularly uring the summer months. Air blowers are commonly designed to eliver a given volume of air. However, the heat balance demands a iven weight of air (oxygen). Therefore, the amount (by weight) of ir pumped by an air blower decreases with: • Increasing air blower inlet temperature • Increasing ambient relative humidity • Decreasing suction pressure
As the air rate is increased, low-cost items that can be implemented o increase the flow of air/oxygen into the regenerator include: • Ensuring the air blower suction filters are clean • Ensuring the pressure drop in the suction piping is not excessive • Ensuring the pressure in the air blower discharge piping system, particularly across the check valve and air preheater, is not excessive
o deliver more air consider: • Lowering the regenerator pressure • Lowering the regenerator catalyst bed level
Evaluate the trade-off between the air blower capacity and wet gas ompressor capacity. Spare horsepower at one can be used to unload he other. Consider: • Cooling the inlet air through the use of a chiller or suction water spray • Using portable air blowers during the hottest months • Oxygen injection • A bypass around the air heater
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Fluid Catalytic Cracking Handbook
Other more capital-intensive modifications include installing a dedicated air blower for the spent catalyst riser. The spent catalyst riser often requires a higher back-pressure than the main air blower to deliver the catalyst into the regenerator. Therefore, less total combusion air would be available if one common blower is used to transfer pent catalyst and provide combustion air to the air distributors. Taking higher-pressure services off the main air blower can allow t to run out on the curve and deliver more air. The main air blower an also be upgraded to provide added capacity. This includes reducing eal clearance, increasing the flow passing area, and increasing the wheel tip diameter. The original equipment manufacturer (OEM) can be contacted for feasibility of this upgrade.
Regeneration
Regenerator designs have changed since most units were built. If he unit test run indicates high CRC or if the catalyst will benefit from a lower CRC, the regenerator internals should be reviewed. If the data ndicates wide temperature differences across the bed or afterburning, or if the unit has had some excursions, it should be examined. The regenerator review will include spent catalyst distribution, air distribution, and cyclones. If the test run with heavy feed indicates a emperature limitation, catalyst coolers, partial combustion, or riser quench should be considered.
FLUE GAS SYSTEM
The FCC is usually constrained by environmental permits. If the unit undergoes significant expansion, it may lose "grandfather" protection. The environmental limits include the amount of coke burned in the egenerator and emission rates of particulates, SOx, NOX, and gasoline ulfur. Increasing the feed rate or running heavier crude can increase ll of these emissions. The technology for control is discussed in Chapter 10.
FCC CATALYST
The FCC catalyst's physical and chemical properties dictate how much feed can be processed. Chemical properties, such as rare earth
Debottlenecking and Optimization
2§7
nd unit cell size (UCS), affect the unit heat balance and wet gas ompressor loading. Physical properties, such as particle size distriution and density, can limit catalyst circulation. Consider reformulating the catalyst—custom formulations are availble. Increasing rare-earth content can reduce the wet gas rate. Catalyst s usually selected for properties other than its ability to flow. Howver, if it does not flow, it is not going to work well. Catalyst physical roperties should be compared with those of catalysts that have circuated well. Evaluate the economics of using metal passivation additives nd other catalyst enhancing additives.
Debottlenecking Main Column and Gas Plant
Debotflenecking usually results in more feed. Both the main fractionator nd the gas plant must be able to recover the incremental product. The main fractionator can be limited by several factors including: • Heat removal limitations » Tray flooding • Fouling and coking Heat removal can be limited by several factors including: • • • • •
Fixed reboiling duties in the gas plant Lack of heat exchanger in the pumparound circuits Jet or liquid flooding in one or more sections of the main fractionator High bottoms temperature leading to fouling or high LCO endpoint Overhead condensing capacity
Moving heat up the tower improves fractionation by increasing the apor-liquid traffic. This is limited by flooding constraints and excesive temperature in the bottom. One method of maximizing the LCO end point is to control the main ractionator bottoms temperature independent of the bottoms pumpround. Bottoms quench ("pool quench") involves taking a slipstream om the slurry pumparound directly back to the bottom of the tower, hereby bypassing the wash section (see Figure 9-9). This controls the ottoms temperature independent of the pumparound system. Slurry s kept below coking temperature, usually about 690°F, while increasng the main column flash zone temperature. This will maximize the CO endpoint and still protect the tower.
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Fluid Catalytic Cracking Handbook
HCO
Reactor Vapors
P/A
Pool Quench
TC J-^
Prod
Slurry —> Recycle
Figure 9-9.
Pool quench to main column bottoms.
If the main fractionator bottoms temperature is limited to 690°F, dding a "pool quench" can provide additional LCO product recovery. Assuming there is no penalty for the bottoms product quality and there s available cooling capacity in the upper section of the fractionator, his incremental LCO yield is valuable. If flooding occurs in the main fractionator, increasing the bottoms umparound rate reduces vapor loading, but can have a negative affect n fractionation. Normally, the economic incentive is to maximize the fresh feed rate nd/or conversion, sacrificing the bottoms cut-point and rate. Increasng conversion by 1.5% (through increasing the riser top temperature y 10°F), provides an incremental profit even though LCO is lost o bottoms. Either high-capacity packing and/or high-efficiency, high-capacity trays an be installed. Trays in the bottoms wash-section can be replaced with rid or packing. The packing has greater capacity at lower pressure drop. The typical packed column has one or more beds, each consistng of packing, a support plate, a hold-down support plate, and a quid distributor.
Debottlenecking and Optimization
299
In a packed column, liquid and vapor flow counter-currently and eparation between the liquid and vapor phases takes place connuously. In contrast, in a column with trays, separation occurs in tages. In a packed column, vapor does not bubble through the liquid s in the columns with trays. For this reason, and due to the absence f the vapor-flow orifices, packed columns operate at a much lower ressure drop. In addition, because liquid and vapor contact in a acked column is less agitated than in a trayed column, packed olumns are less likely to foam. Satisfactory operation must be between the upper and lower limits or both liquid and vapor flow rates. At liquid rates below 0.5 GPM er square foot of packing cross-section, liquid distribution is not niform enough to ensure thorough wetting. At liquid rates between 5 GPM and 70 GPM per square foot of packing, the column is onsidered liquid-loaded and becomes very sensitive to additional quid or vapor flow. An adequate vapor rate produces a pressure drop greater than 0.1 nch of liquid per foot of packing. Flooding occurs when the pressure rop exceeds 1.3 to 2.5 inches of liquid per foot of packing. At high apor rates, the liquid cannot flow down the column. The liquid distributor is the most important internal structure of a acked column. The distributor strongly influences packing efficiency. must spread the liquid uniformly, resist plugging/fouling, provide ree space for gas flow, and allow operating flexibility. Packed columns can flood prematurely. Some of the reasons include: • Fouling (caused by precipitation, lodgment of loose material and debris damaged packing) • Foaming • Improper feed introduction • Restricted liquid outlet
In addition to changing to packing or high-efficiency trays, the tower an be unloaded by: • Removing more heat from the pumparound returns, either by generating steam or adding coolers. This can decouple the fractionator from the reboilers in the gas concentration unit. • Reviewing the LCO product system. If some or all of the LCO is being hydrotreated, that portion can bypass the stripper if it is
300
• •
• «
Fluid Catalytic Cracking Handbook
direct-fed to the other unit through pressure vessels. Stripping is difficult to justify and sends wet feed to the unit. Changing the control system so stripping steam flow is proportional to LCO stripper product. Reviewing the overhead water wash: most overhead condensers are washed continually to minimize fouling. Since multiple bundles are common, solenoids and a PLC can be used to wash one bundle at a time, for approximately ten minutes each. This can lower the pressure drop and increase the available cooling with minimal impact. Advanced instrumentation can be used. If the rich oil is being returned from the secondary absorber, consider different processing.
Dehottlenecking Wet Gas Compressor
A portion of liquid from the overhead receiver is refluxed back to he tower and the remainder is pumped on to the gas plant. The vapor rom the receiver goes to the wet gas compressor. The pressure of he reactor/main fractionator system is usually controlled at the compressor suction. Improving overhead cooling will increase the wet gas compressor capacity. Excessive pressure drop or limited cooling in the overhead system decreases the capacity. This can result from: • • » «
Inadequate surface area Uneven distribution of hydrocarbon vapors and/or cooling water. Corrosion and salt deposition Limited water flow rate from elevated water coolers (consider adding a booster pump at grade) • Rapid fouling caused by water outlet temperature above 125°F
The wet gas compressor is always run to a limit, therefore, increasing he available flow will always benefit the unit. The flow can be ncreased by: • Parallel cooling of the overhead vapor. The pressure drop across overhead cooling systems ranges from 2.0 psi to more than 10.0 psi; 5.0 psi is typical. • On-line solvent or water wash to minimize blade fouling on both the compressor and turbine.
Debottlenecking and Optimization
301
• Closing the spillback valves. • Removing external streams. If gas comes from another unit or vents from a column in the gas concentration unit, consider routing it to the interstage rather than the suction. The refinery needs to evaluate if external streams are worth recovering or whether they can be routed elsewhere. • Installing an advanced surge control system. • Verifying that the flow rates of corrosion inhibitor and antifoulant are adequate for the new operating conditions.
mproving Performance of Absorber and Stripper Columns
The objective of the primary absorber/stripping towers is to maximize recovery of C3 and heavier components while rejecting C2 and ghter to fuel. C3 is first absorbed and then C, and lighter components re stripped. Although maximizing C3-C4 recovery for alkylate feed very profitable, lower recoveries are often accepted to maximize the CC conversion and/or feed rate. Propane/propylene recovery can be enhanced by: • Increasing the gas plant pressure. A 10 psi increase in absorber pressure increases C3 recovery by 2% (Figure 9-10). However, this can reduce the wet gas compressor capacity. Fractionation efficiency decreases as the column pressure increases. • Reducing the operating temperature. Consider adding an intercooler on the absorber. Minimize lean oil temperature. Consider the use of a chiller. Each 10°F reduction in lean oil temperature will increase C3's recovery about 0.8% (Figure 9–11). • Increasing lean oil rate. This rate is often limited by the debutanizer hydraulic and reboiling/cooling capacity. A 50% increase in lean oil/off-gas ratio increases C3's recovery about 2%. • Removing water from the lean oil. Installation of water draws and/ or a coalescer can improve recovery. Water can become trapped in the tower and cause poor tray efficiencies, foaming, and premature flooding. • Minimizing over-stripping. Over-stripping can start a wheel with the absorber. A 10% cut in stripping rate can increase C3's recovery by 0.8% (see Figure 9-12). (text continued on page 304)
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Fluid Catalytic Cracking Handbook
96
10
20
30
40
50
Delta System Pressure, psig Figure 9-10.
C3 recovery vs. system pressure.
93.0
91.0
60
65
70
75
80
Lean Oil Temperature, °F Figure 9–11. C3 recovery vs. lean oil, °F.
85
90
-
flg
3-
1
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text continued from page 301)
Debottlenecking Debutanizer Operation
As the gasoline Reid vapor pressure (RVP) is reduced, the operation of the debutanizer becomes more critical. The allowable vapor pressure n gasoline makes it difficult to prevent heavy ends in the alkylation eed. This can limit the production of gasoline without sacrificing alkylation. This limitation is often from insufficient overhead cooling and rebelling: * Optimum debutanizer feed preheat temperature can optimize column loading. Increasing preheat temperature reduces reboiler duty and loading in the stripping section of the tower. Decreasing preheat temperature decreases overhead condensing duty and loading in the rectifying section. Adding an exchanger on the stripper bottoms can make this a controllable variable. * Delta P indicators should be installed on both the top and bottom section. * Optimize the operating pressure to balance reboiling, condensing, and loading. Consider floating pressure control. With tightening vapor pressure specifications, the debutanizer is an excellent candidate for this type of control. Floating pressure will unload the tower and provide better separation. * If slurry pumparound is the heat medium, consider HCO pumparound to minimize fouling. * Revamp the tower internals with high-capacity trays or packing. * If the receiver vent is in continuous service, route it back to the wet gas compressor interstage rather than to the suction. Consider adding a chiller on the vent gas.
NSTRUMENTATION
Additional analyzers should be considered. Temperature and pressure re no longer adequate to control distillation columns to tight specifcations. Consider chromatographs on the overhead streams. One hromatograph with multiple sample streams can be adequate for most ervices. Ensure that qualified service is available locally.
Debottlenecking and Optimization
305
If the unit does not have a distributed control system (DCS), a ebottlenecking project is the right time to justify it. If it does have a DCS, advanced control projects should be justified. A DCS: • Will provide better control of the unit and stay closer to constraints. Operating closer to constraints is what optimization and debottlenecking are all about. • Has trending and reporting ability. Data can be dumped to a spreadsheet program and variables plotted against one another. • Is a valuable troubleshooting tool. • With a host computer allows moving on to advanced control and rnulti-variable control. The unit is sensitive to day/night temperature swings and the multi-variable control can track ambient changes.
Many case histories are available on converting to a DCS on the un or during a turnaround. Upgrading will pay off in the long ran.
Tankage/Blending
Significant debottlenecking in the FCC will affect the tankfarm and lending system. They will handle increased product yields and changes n the quality. Blending needs maximum warning about changes in asoline components.
team/BFW
Adding a catalyst cooler may back a boiler down, or it may require more BFW and a home for the steam. New feed nozzles may require more steam. A cogeneration unit can be an attractive option.
our Water/Amine/Sulfur
Plant
Running heavier crude to the FCC will convert more of the sulfur n the refinery crude to H2S.
elief System
Increasing the wet gas compressor capacity and increasing duties hrough the gas plant can impact the flare system.
06
Fluid Catalytic Cracking Handbook
uel System
Both offgas rate and composition may change. Verify that increased hydrogen content will not impact any heaters. Depending on the header design, it could be a problem if it all goes o the same branch of the header. With the nation's projected natural gas shortfall and projected price ncreases, this may be a good time to consider gas export.
UMMARY
Cat cracking has been, and will continue to be, a big "money maker" for the refining industry. It is unlikely that any new cat rackers will be built (especially in the U.S.) in the near future. herefore, emphasis will be placed on finding ways to improve the perational reliability and profitability of the existing FCC's. Performance of an FCC unit is often maximized when the unit is perated against multiple constraints simultaneously. It is essential that he specified constraints allow for minimum "comfort zones." An perator-friendly advanced control program, coupled with proper election of catalyst formulation, would allow optimizing the performance of the unit on a daily basis. This chapter provided several cost recommendations that, once mplemented, would provide cost-effective added value to the operaon of the FCC. Examples of such items include tips on debottleecking the air blower, wet gas compressor, and catalyst circulation. his chapter also discussed the latest technologies regarding the riser ermination devices, as well as feed injection systems. Prior to implementing any new technologies, it is critical that the objectives and the mitations of the unit are clearly defined to ensure the expected enefits of the new technology are realized. The selected technology must match the mechanical limitations of given cat cracker. All the technologies discussed in this chapter have een commercially proven, therefore the choice must include the total nstalled costs, as well as the projected benefits to the refinery.
CHAPTER 10
Emerging Trends in Fluidized Catalytic Cracking
Although the demand for transportation motor fuels in North America projected to be limited, economic growth in other parts of the world ill require crude oil-based fuels. The Far East, Latin America, and e former Soviet Union are areas where there will be substantial emand for transportation fuels. The collapse of communism, the rivatization of state-owned oil companies, and the global awareness f "environmentally clean fuels" will cause this growth. In the coming years, the refining industry will be experiencing major hallenges. In the United States, refiners are faced with excess refining apacity, projected slow growth, and high capital and operating costs comply with environmental health and safety regulations. The oil ndustry in general, and the refining industry in particular, are techologically sophisticated. They have a long history of innovations and roven track records in responding to challenge. It is likely that the reliable crude oil supply will not diminish any me soon. Petroleum-derived fuels will remain the primary source of ansportation energy for well into the twenty-first century. Producers nd refiners have been, and will be, environmentally responsible. The xisting infrastructure of advanced product distribution systems can ompete with alternative fuels readily. Future fuels will be competitive, oth economically and environmentally. New global market conditions ill dictate closure of inefficient facilities and investment in new chnology. Larger and more efficient operations will survive and will cus on the "niche market." In the U.S., the crude processing capacity is expected to increase odestly, at a projected rate of 0.5 percent per year. No new refinery
307
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Fluid Catalytic Cracking Handbook
s expected to be built in the U.S. The production of lighter, higher value products is expected to continue. Residual fuel will continue o decline. The demand for gasoline is projected to be stable with xcess octane. Optimum performance and reliability of FCC units will play an mportant role in the competitiveness and survival of refineries. The FCC has proven to be a versatile process, changing to meet the needs nd demands of refiners. As one of the most efficient conversion processes in the refinery, it will continue to play a key role in meeting uture reformulated fuel demands. This chapter discusses: • « • •
Evolution of reformulated fuel and its impact on FCC operations Resid upgrading through the FCC Gaseous emissions from the FCC Emerging developments in catalyst, process, and hardware
REFORMULATED FUELS
The passage of the Clean Air Act Amendment (CAAA) on November 15, 1990, started a process for regulating the composition and quality f gasoline and diesel fuels sold in the United States. The CAAA's ntent was to improve the nation's air quality by reducing ozone and other air pollutants. Title II of the CAAA requires the manufacture nd sale of "cleaner" fuels in order to reduce evaporative and comustible emission of: 1. Volatile organic compounds (VOCs) 2. Nitrogen oxides (NOX) 3. Toxins including benzene, formaldehyde, acetaldehyde, 1,3 butadiene, and polycyclic organic material (POM).
VOC Emissions
VOCs can be emitted from the fuel system and from the exhaust system. Fuel system: Evaporative emissions of gasoline are mainly due to he presence of butane and the low-boiling light olefins (C4 and C5). Reducing gasoline vapor pressure and removing these olefins can limit he amount of evaporative emissions. Light olefins are photo-chemically eactive; removing them will improve ozone.
Emerging Trends in Fluidized Catalytic Cracking
309
Exhaust system: The engine operating mode controls the tailpipe missions of hydrocarbons (HC) and carbon monoxide (CO). Over 0% of HC and CO emissions are generated during cold-start and warm-up due to incomplete combustion. Fuel vaporization and fuel/ ir mixing are important factors in achieving thorough combustion of he hydrocarbons. Gasoline can be modified to vaporize quickly. This is accomplished by: • Decreasing the end point or 90% boiling point « Reducing the aromatic content * Adding oxygenates
Since a gasoline engine burns vaporized fuel, the heavy end of the uel contributes to its partial vaporization in a cold engine. Reducing he 90%-point or the 50%-point temperature will reduce HC emissions n the engine exhaust. Aromatic levels and carbon content of the gasoline also have a gnificant effect on the tailpipe emissions of HC and CO. Because f their high heat of vaporization and high boiling point (see Figure 0-1), aromatics do not vaporize readily. This is an incentive to minimize aromatics.
2.5 c •3 CO
1
X
I
'
1
i
.
\
\
^Others P JP"p
.P
•« -Op..
1
0.5
0
-
\
Aromatics *A XV^ y •*»^ Vx "*^*.^__
.
1.0
I
\ Alcohols
8-3.1.5
<s
^x
2.0
•8o o* si o
!
___
A""**A"At». c r\
p. p.. op
_:
""*«*._
P
P
t 1 1 f i 100 200 300 Boiling point, °F
-
-
1 400
X = alcohols, A = aromatics, P = paraffins, O = olefins and E = ethers. Figure 10-1. Heat of vaporization versus boiling point [16].
10
Fluid Catalytic Cracking Handbook
Oxygenates reduce CO emissions by "enleaning" the fuel-to-air mixture. Enleanment of the fuel with oxygenates has the most impact n CO emissions. However, oxygenates, particularly ethers, are often sed as a "substitute" and can replace aromatics in achieving octane pecifications. Reducing aromatics further reduces CO and HC emissions,
NOX Emission
Direct exposure to NOX can cause respiratory problems. VOCs and NOX are catalyzed by sunlight to form ground level ozone, often eferred to as smog. NOX can be generated from either stationary ources or mobile sources. In 1997, the EPA changed the ambient tandard for ozone from 0.12 ppm to 0.08 ppm and the applicable test eriod was increased from 1 to 8 hours. The main stationary sources of NOX are gas turbines, fired heaters, ower generation plants, and, of course, the FCC. The amount of NOX roduced is a function of residence time and combustion temperature. ombustion temperature is influenced by fuel composition. The main mobile source of NOX is the combustion of a fuel in an nternal combustion engine. Because aromatics have the highest comustion temperature among hydrocarbon types (see Figure 10-2), they
i
i
u- 4,200 -
£ I 4,000~-o
Aromatics Oiefins u
1 3,800
1 ig 3,600
'*~
Others x* Alcohols /x
j 0
""""""•* A«^,,Aii_ • ~ * " " "" • "o™*" * *"»O.
*
i
100 200 Boiling point, °F
300
Theoretical flame temperatures assuming adiabatic and stoichiometric air.
igure 10-2. Flame temperature of aromatics, olefins, paraffins, and cohol [16].
Emerging Trends in Fluidized Catalytic Cracking
311
end to produce higher amounts of NOX in the exhaust gases than lefins and paraffins. One might think that because oxygenates have lower combustion emperatures, they will generate less NO X . However, because the nleanment effect raises combustion temperature, oxygenates actually ncrease NOX emissions. Consequently, some compromise may be eeded with respect to oxygen content of fuels and its effect on HC, CO, and NOX emissions.
Benzene Emission
Benzene is a known carcinogen. The U.S. Environmental Protecon Agency (EPA) has identified benzene as a toxic air pollutant TAP). Benzene is present in automotive evaporation, refueling vapors, nd exhaust. The control of benzene emissions from the fuel system includes miting the amount of benzene in the fuel and vapor recovery at uel stations. Exhaust benzene is a function of aromatics and benzene content. xhaust benzene emission is calculated by: EXB = 1.884 + 0.949 x BZ + 0.113 x (A - BZ)
Where:
XB = Exhaust benzene, milligram/mile Z = Benzene, vol% = Aromatics, vol%
or example, the exhaust benzene for a gasoline having 30 vol% romatics and I vol% benzene is = 6.11 milligram/mile.
CAAA Regulations
As of November 1, 1992, all gasoline sold in the 39 CO nonttainment areas contained 2.7 wt% oxygen during the winter months. Beginning January 1, 1995, regulations mandated that gasoline sold n the nine worst ozone non-attainment areas contain at least 2.0 wt% xygen and not more than 1 vol% benzene and 25 vol% total aromatics. Other cities that have had mobile source emission problems can "optn" voluntarily to the use of reformulated fuels.
12
Fluid Catalytic Cracking Handbook
From 1995 through 2000, a 15% reduction in VOCs and other air oxins was expected in the non-attainment areas. Beginning in the year 000, an additional reduction of 25% is required unless the EPA etermines it too costly or not feasible. Even in that case, the reduction will not be less than 20%. Oxygen was added as oxygenated hydrocarbon components: methyl ert-butyl ether (MTBE), tert-amyl methyl ether (TAME), ethyl tertutyl ether (ETBE), di-isopropyl ether (DIPE), ethanol, methanol, and ertiary butyl alcohol (TBA). The properties of oxygenates, as they elate to gasoline blending, are shown in Table 10–1. The key points of the regulations require: * The certification of fuels: Each refiner, blender, or importer of gasoline must ensure that per-gallon emissions levels of VOCs, NOX, CO, and toxic air pollutants do not exceed the gasoline sold in 1990. • Effective October 1993, highway diesel fuel was limited to a maximum sulfur content of 0.05 wt% (500 ppm) and a minimum cetane rating of 40. « Engine additives were required in all gasoline to prevent deposits in engines and fuel supply systems. » Lead and lead additives were eliminated by the end of 1995.
Blending octane (R + M)/2 Blending RVP, psi Boiling point, °F Density @ 60°F, Ib/gal Water solubility, wt% Max. concentration, vol% ax. oxygen, wt%
Table 10-t Oxygenates Properties MTBE
ETBE TAME TBA
Ethanol
Methanol
110
111
105
100
115
108
8 131 6.2
4 161 6.2
1 187 6.4
181 6.6
18 173 6.6
31+ 148 6.6
1.4 15.0
0.6 12.7
12.4 16.1
10
9.7
2.7
2.0
2.0
3.7
3.7
3.7
Emerging Trends in Fluidized Catalytic Cracking
313
The standards for conventional or non-RFG gasoline are shown in able 10-2. Oil companies can choose to comply with the requirements n either a per-gallon baseline or the company's 1990 baseline. If there s an incremental volume of fuel above the 1990 production rate, the aseline will be adjusted using the industry's baseline data (see Table 10-3). The industry's baseline gasoline is an average of properties of ll the U.S. gasoline marketed in 1990. The Simple Model RFG required the addition of oxygenates and it mited the amount of benzene, sulfur, olefins, and T9Q. The RVP was lso lowered for six months during the summer period. Given these equirements, companies can choose to comply on a per-gallon basis Table 10-4) or adopt the 1990 industry average basis (Table 10-5), Starting January 1998, the EPA's Complex Model went into effect. he Complex Model provides a set of equations that predict VOC, NOX, and toxic emissions, using eight gasoline properties. These roperties are RVP, oxygen, aromatics, benzene, olefins, sulfur, E200,
Table 10-2 Conventional Gasoline Standards Specifications
Properties
xhaust benzene, mg/mile ulfur, ppmw—yearly average Olefins, vol%—yearly average , °F—yearly average
Maximum Maximum Maximum Maximum
100% 125% 125% 125%
Table 10-3 U.S. Industry 1990 Baseline for Non-RFG Gasoline
Aromatics, vol% Olefins, vol% Benzene, vol% Sulfur, ppmw Exhaust benzene, mg/mile T_ °F
28.6 10.8 1.6 338 6.45 332
of of of of
baseline baseline baseline baseline
14
Fluid Catalytic Cracking Handbook Table 10-4 RFC Simple Model per Gallon Standards
VP VOC Control Region 1 (south) VOC Control Region 2 (north) Oxygen content, wt% oxics reduction enzene, vol% ulfur, ppmw—yearly average Olefins, vol%
1,2 psi, maximum 8,1 psi, maximum 2.0-2.7 15.0%, minimum 1.00%, maximum 100% baseline, maximum 100% baseline, maximum 100% baseline, maximum
Table 10-5 RFG Simple Model Average Gasoline Standards (Phase I)
Oxygen content, wt%
oxics reduction enzene, vol%
ulfur, ppmw yearly average Olefins, vol% yearly average 90, °F, — yearly average
VOC—Control Region 1 (south) Standard: 7.1 psi, max. Per-gallon: 7.4 psi, max VOC—Control Region 2 (north) Standard: 8.0 psi, max. Per-gallon: 8.3 psi, max. Standard: 2.1-2.7 Per-gallon: 1.5-2.7 16.5%, min. Standard: 0.95, maximum Per-gallon: 1.30, maximum 100% baseline, maximum 100% baseline, maximum 100% baseline, maximum
nd E300. E200 and E300 are the percent of gasoline evaporated at 00°F and 300°F, respectively. The Complex Model contains the following: • Seven exhaust emission equations for VOCs and NOX, and five for toxins (benzene, butadiene, formaldehyde, acetaldehyde, and polycyclic organic material (POM).
Emerging Trends in Fluidized Catalytic Cracking
315
» Four non-exhaust emission equations for VOCs (diurnal, hot soak, running loss, and refueling emissions). * Four corresponding non-exhaust emission equations for benzene.
hese non-linear equations can be embedded into the refinery's linear rogram (LP) to achieve compliance and optimize the gasoline blend. The key FCC gasoline components that influence RFG are: • Sulfur * Benzene and aromatics • Olefins
In the year 2000, Phase II of reformulation begins. The Phase II andards are shown in Table 10-5A. Compliance with the standards determined using the Complex Model. The regulations for reformulated gasoline were published in the federal register on February 6, 1994.
ulfur
Sulfur in gasoline contributes to the SOX air quality problem and eactivates the catalyst in the catalytic converter. Emissions from a oisoned converter contain higher levels of VOC, NOX, and CO. As ated earlier, VOC and NOX are catalyzed by sunlight to form smog. Table 10-5A Complex Model Phase II Per Gallon Standards (After Year 2000) VOCX emissions performance reduction
VOC control Region 1 VOC control Region 2
> 27.5 > 25.9
Toxic emissions reduction (%)
NOX emissions reduction (%) Gasoline designated as VOC controlled Gasoline not designated as VOC controlled
> 5.5 > 0.0
Oxygen (wt%)
> 2.0
Benzene (vol%)
> 1.0
16
Fluid Catalytic Cracking Handbook
he sulfur compounds in FCC gasoline consist of Cj-C4 mercaptan nd various thiophenes. The California Air Resources Board (CARB) set an average sulfur specication of 40 ppm for 1996, with a maximum of 80 ppm. The CAAA's omplex Model also addresses sulfur issues in its set of equations. For the U.S., the new EPA rules will limit sulfur in gasoline to 30 pm, phased between 2004 and 2006. The automobile industry has made a strong case for lower sulfur because of its effect on the atalytic converter. The converter has the same catalyst as the refinery eformer and it is poisoned just as easily by sulfur. Refiners can address the sulfur issue in stages, but decisions should e made that will leave the door open for further reductions. If ydrotreating is selected, the design can include oversized reactors, onnections for a spare compressor, or connections for adding amine crubbing inside the recycle loop. Some process or catalyst changes an buy time, some can solve the problem. Reducing gasoline sulfur specifications is not limited to the U.S. n Canada, gasoline sulfur will be reduced in two phases. The first hase is a reduction from the existing level of 360 ppm to 150 by 002. Beginning in 2005, the gasoline sulfur level will be reduced to 0 ppm. In Europe, beginning in the year 2000, the gasoline sulfur in he 15 countries in the European Union is reduced to 150 ppm. By he year 2005, the gasoline sulfur will be lowered to 50 ppm. In Japan, ypical gasoline sulfur is 35 ppm, which is below the current requirement of 100 ppm. Other countries are expected to follow the lead to nact regulations to reduce gasoline levels. FCC gasoline is by far the largest sulfur contributor (up to 90%) n the gasoline pool. Typical sulfur content of the FCC gasoline ranges om 150 ppm to 3,500 ppm. The amount of the FCC gasoline in the nished gasoline blend normally ranges from 35% to 45%. Controlling asoline pool sulfur requires reducing the sulfur content of FCC asoline. Several options are available: • • • • •
FCC feed hydrotreating Gasoline end point reduction FCC gasoline hydrotreating Catalyst additives Bio-catalytic desulfurization
Emerging Trends in Fluidized Catalytic Cracking
317
For nonhydrotreated feed, the gasoline sulfur level is about 10% of he feed sulfur level. For hydrotreated feed, it is about 5%. For xample, if the sulfur content of a nonhydrotreated feed is 1.0% 10,000 pprn), the sulfur in FCC gasoline will be 1,000 ppm. Assuming 0% desulfurization, feed to the FCC unit will contain 0.2% (2000 pm) sulfur, resulting in FCC gasoline containing 100 ppm sulfur. ulfur compounds that survive the hydrotreater are in the heavy raction and tend to end up in the LCO, decanted oil, and coke. Hydrotreating or moderate pressure hydrocracking of the FCC feed rovides many benefits. Besides reducing sulfur in the FCC gasoline, CC feed hydrotreating reduces NOX and SOX emissions from the egenerator flue gas, increases conversion, increases gasoline yield, and educes catalyst consumption. Often around 95% desulfurization is equired to achieve the desired gasoline sulfur. However, a number of efiners cannot justify the high capital cost of FCC feed hydrotreating. Reducing the gasoline end point can significantly decrease the FCC asoline sulfur (Figure 10-3). As much as 50% of the sulfur can be ontained in the last 10 vol%. With a high sulfur crude mix, this end oint reduction may not be sufficient to meet sulfur specifications. The isposition of this high-sulfur, high-aromatic gasoline can be a problem. One option is to combine this heavy fraction with the LCO stream and esulfurize it in the diesel hydrotreater. After hydrotreating, the heavy asoline can be separated and sent to the gasoline pool. This may require onverting the stripper into a fractionator or adding a fractionator. Selective hydrogenation (HDS) of the FCC gasoline can be a posive choice for meeting the required sulfur levels. However, deep HDS f the FCC gasoline can saturate olefins and cause octane loss. The light CC gasoline is rich in olefin while the heavy FCC gasoline is rich in romatics and sulfur. The choice of a proper catalyst and operating onditions is important in maximizing sulfur reduction and minimizing ctane loss. A number of commercial processes are proven in this service. Caustic extraction can remove mercaptan sulfur in light fractions, ut not higher carbon number mercaptans or other types of sulfur molecules that are in the FCC gasoline. Catalyst additives can reduce FCC gasoline sulfur by about 15%, hey work by converting mercaptan, thiophene, etc., to H7S. A seconary benefit of the additives is an approximate 10% reduction in the CO sulfur.
Figure 10-3,
400
420
430
Gasoline sulfur versus its end point
Gasoline End Point, Deg. F
410
440
Emerging Trends in Fluidized Catalytic Cracking
319
Biocatalytic desulfurization (BDS) technology employs the concept f enzymatic removal of sulfur compounds without changing the undamental structure of hydrocarbons. This process uses bacteria to electively remove sulfur from gasoline, diesel, and diesel blend tocks. In 1997, the Coordinating Research Council (CRC) commisioned a study to assess the applicability and feasibility of biodesulfuriation technology for the removal of sulfur from gasoline. It is xpected that this process will be ready for commercialization within he next four to six years. Energy BioSystems Corp, The Woodlands. exas is developing a process to use bacteria for the removal of sulfur om diesel and diesel blend stocks. The first commercial BDS unit s expected to be operational in 2001. The choice for removal of sulfur from gasoline and diesel requires xamining the whole refinery and understanding its total impact. For ome refiners, post-treating, or pre-treating and post-treating may be he right decision. Factors to consider include available capital, availability f tankage/utilities, and the reliability of newly installed technology,
Aromatics and Benzene
Tailpipe emissions of HC and CO are affected by the levels of heavy romatics in gasoline. Like sulfur, the heavy aromatics are in the back nd of the boiling range (Figure 10-4). As with sulfur, reduction of nd point directly controls the concentration of heavy aromatics in nished gasoline. The benzene content of FCC gasoline is typically in the range of .6 vol% to 1.3 vol%. CAAA's Simple Model requires RFC to have maximum of 1 vol% benzene. In California, the basic requirement also 1 vol%; however, if refiners are to comply with averaging rovisions, the maximum is 0.8 vol%. Operationally, the benzene ontent of FCC gasoline can be reduced by reducing catalyst-oil ontact time and catalyst-to-oil ratio. Lower reactor temperature, lower ates of hydrogen transfer, and an "octane catalyst" will also reduce enzene levels. Most of the benzene in the gasoline pool comes from the reformer nit (reformate). To reduce the reformate's benzene, one must modify he feedstock quality and/or operating conditions. Benzene's precursors n the reformer feed (C5 and C6) can be prefractionated and sent to n isomerization unit. The reformer operating pressure can be reduced
ui
O
8 ti %'|OA '
Emerging Trends in Fluidized Catalytic Cracking
321
o reduce benzene and aromatics. Post-fractionation of the re formate tream is another option but it requires the installation of a reformate plitter. The light aromatics (benzene, toluene, and mix xylene) in the light" reformate can be extracted for petrochemical feedstock. Toluene an be converted to benzene through the hydrodealkylation process. Benzene can be saturated to cyclohexane and eventually isomerized. A combination of the benzene saturation system and paraffin isomeriation will enable the refiner to control benzene while improving the asoline octane pool.
Olefins
In 1990, U.S. gasoline contained about 10 vol% olefins, the majority f which emanate from FCC gasoline. FCC gasoline has 25 vol% to 35 ol% olefins. Of these olefins, C5–C7 olefins account for about 85% of he total pool. For non-RFG gasoline, as with sulfur, the regulation allows he maximum olefin content to be 125% of the 1990 baseline values. Light olefins, particularly tertiary olefins, are very reactive in formng ozone and also increase gasoline pool RVP. However, the future end of most FCC operations is projected to produce more olefin feed, ut little will reach the gasoline pool. This is because olefins, particuarly the C4 and C5 olefins, can either be "alkylated" and/or "etherified," r used for petrochemical feedstock. Commercial alkylation is the reaction of isobutane with C3 through C5 olefins in the presence of either sulfuric acid or hydrofluoric acid see Example 10-1). Etherification is the reaction of a tertiary olefin with an alcohol or water in the presence of an acidic catalyst (see xample 10-2). Example 10-1 Alkylation of Propytene and Butylene
ropylene Alkylation CH 3 — CH = CH2 + CH3 — CH — CH2 -> CH 3 — CH — CH — CH 2 —CH 3 CH3
PROPYLENE
ypical Yield:
+
ISOBUTANE
CH3
CH3
-> DIMETHYLPENTANE
22
Fluid Catalytic Cracking Handbook
.0 Volume of propylene + 1.3 volume of isobutane -> 1.80 volume f alky late.
Butyiene Alk\ lation CH3 CH3 — C = CH2 + CH3 — CH — CH2 -> CH3 — CH — CH, — CH — CH3 CH,
PROPYLENE +
CH3
CH,
ISOBUTANE
-»
CH3
2,2,4 TRIMETHYLPENTANE
ypical Yield:
.0 volume of butylene +1.2 volume of isobutane —> 1.70 volume f alkylate.
Example 10-2 Etherification of Isobutviene
CH3 — C — CH3 + CH3 — OH CH?
-> CH3 — C — O — CH3 CH3
ISOBUTYLENE + METHANOL -> METHYL TERTIARY BUTYL ETHER (MTBE)
ypical Yield:
.0 volume of isobutane + 0.43 volume of methanol —> 1.27 volume f MTBE.
There are etherification processes, such as MTBE and TAME, aimed t producing ethers from C5, C6, and C7 tertiary olefins. Both alkylate and ether have excellent properties as gasoline blending omponents. They have a low RVP, a high road octane, no aromatics, nd virtually zero sulfur. The emphasis on alkylation and etherification will continue in both the U.S. and the rest of the world.
Emerging Trends in Fluidized Catalytic Cracking
323
A conventional FCC unit can be an "olefin machine" with proper perating conditions and hardware. Catalysts with a low unit cell size and high silica/alumina ratio favor olefins. Additionally, the addition of ZSM, with its lower acid site density and very high framework silica-alumina atio, converts C?+ gasoline into olefins. A high reactor temperature nd elimination of the post-riser residence time will also produce more lefins. Mechanical modification of the FCC riser for "millisecond" racking has shown potential for maximizing olefin yield.
Challenges Facing RFG
RFG is a cost-effective fuel that improves air quality and is a mechanism through which the refining industry can be competitive. The Complex Model is most likely here to stay. The concentration of asoline sulfur must be reduced and the gasoline RVP will most likely e limited to about 7 psi. Nevertheless, in the years to come, numerous ssues regarding RFG will be facing refiners. Most are regulatory, olitical, and bureaucratic issues. Following are some of these issues: « Public perception of RFG regarding health effects of ethers, price increase, and engine performance complaints * EPA's ethanol mandate and the subsequent stay of that mandate by federal court * Complexity of testing, distribution, storage, handling, and blending facilities * Record-keeping and development of a uniform certification program, * Intel-changeability of MTBE to ETBE * Interpreting the baseline * The future of opt-in areas: the continual decline in air quality where RFG is not sold * Antidumping, credits, and trading * The program length of oxygenated fuels for CO nonattainment areas * The definition of "domestic supply"
RESIDUAL FLUIDIZED CATALYTIC CRACKING (RFCC)
Deterioration in the worldwide crude oil supply (Table 10-6), ontinual decline in the demand for heavy fuel oil, and recent mechancal and catalyst advances have provided economic incentives to
24
Fluid Catalytic Cracking Handbook Table 10-6 U.S. Crude Characteristics
Year
°APi Gravity
Wt% Sulfur
1983 1984 1985 1986 1987 1988 1989 1990 1991 1992 1993
32.92 32.96 32.46 32.33 32.22 31.93 32.14 31.86 31.64 31.32 31.30
0.88 0.94 0.91 0.96 0.99 1.04 1.06 1. 10 1.13 1.16 1.15
Source: Swain [24]
upgrade the atmospheric and/or vacuum bottoms in the residual fluidzed catalytic cracking (RFCC) unit. Although residue upgrading in he United States is mostly delayed coker based, most new FCC units are either residue crackers or have in-place provisions to process esidue at a later date. This is more pronounced in the new units built n the Far East, Europe, and Australia. The residue from their crude oils is more paraffinic and contains less metals than North Sea or Middle Eastern crude oils, which makes them more suitable for RFCC. An RFCC is distinguished from a conventional vacuum gas oil FCC n the quality of the feedstock. The residue feed has a high coking endency and an elevated concentration of contaminants.
Coking Tendency
Residue feedstocks have a higher coking tendency, which is indicated by higher levels of Conradson carbon and a higher boiling point. The common definition of residue is the fraction of the feed that boils above 1,050°F and Conradson carbon levels greater than 0.5 wt%. The esidual portion of the feed contains hydrogen-deficient asphaltenes and polynuclear compounds. Some of these compounds will lay down on active catalyst sites as coke, reducing catalyst activity and selectivity.
Emerging Trends in Fluidized Catalytic Cracking
325
Feed Contaminants
The residual portion of feedstocks contains a large concentration of ontaminants. The major contaminants, mostly organic in nature, nclude nickel, vanadium, nitrogen, and sulfur. Nickel, vanadium, and odium are deposited quantitatively on the catalyst. This deposition oisons the catalyst permanently, accelerating production of coke and ight gases. Nickel in the feed is deposited on the surface of the catalyst, romoting undesirable dehydrogenation and condensation reactions. These nonselective reactions increase gas and coke production at he expense of gasoline and other valuable liquid products. The eleterious effects of nickel poisoning can be reduced by the use of ntimony passivation. Vanadium in the feed poisons the FCC catalyst when it is deposited n the catalyst as coke by vanadyl porphydrine in the feed. During egeneration, this coke is burned off and vanadium is oxidized to a V*5 oxidation state. The vanadium oxide (V2O5) reacts with water vapor in the regenerator to vanadic acid, H3VO4. Vanadic acid is mobile and it destroys zeolite crystal through acid-catalyzed hydrolysis. Vanadic acid formation is related to the steam and oxygen conentration in the regenerator. Vanadium and sodium neutralize catalyst acid sites and can cause ollapse of the zeolite structure. Figure 10-5 shows the deactivation f the catalyst activity as a function of vanadium concentration. Destruction of the zeolite by vanadium takes place in the regenerator where the combination of oxygen, steam, and high temperature forms anadic acid according to the following equations: 4V + 5 O2 --» 2 V2O5 V2CL + 3 H2O -* 2 VO (OH)3
The produced vanadic acid, VO (OH)3, is mobile. Sodium tends to ccelerate the migration of vanadium into the zeolite. This acid attacks he catalyst, causing collapse of the zeolite pore structure. The presence of increased basic nitrogen compounds, such as pyridines nd quinoline in the FCC feedstock, also attack catalyst acid sites. The esult is a temporary loss of catalyst activity and a subsequent increase
26
Fluid Catalytic Cracking Handbook
68
^>T~^ - ^ T
67
* *
i ^***^***«*^
"
I
C'C^L
^'rC? N „
>7, ^^^^S-^4
? 65 > 1 64
"j-. j Q $^ ^"""P"-^
a
o 63 5 6?
^T*—^^
I
'~~\?°-o . I * ^ sit^O-.
i "''"i. f
h
t
61
~' ~ '"-*.' '"ft "
~~1
60
0
1000
2000
3000
4000
5000
6000
Vanadium, ppm
Figure 10-5. Vanadium deactivation varies with regenerator severity {25],
n coke and gas yields. Additionally, in the regenerator, some of the adsorbed nitrogen is converted to nitrogen oxide (NOX). Although an increase in the sulfur content of the residue feedstock will have a minimal effect on unit yields, the sulfur content of the RFCC products and the flue gas is greater, requiring additional treatng facilities.
Operational Impacts of Residue Feedstocks In the unit, residue feedstocks have the following effects: * Higher delta coke and coke yield, which are associated with residue feedstocks, will result in elevated regenerator temperature and higher combustion air requirements. * Exposure of the catalyst to a variety of feed contaminants and the higher regenerator temperature will reduce both selectivity and activity. * Greater levels of nitrogen and sulfur in the residue feed increase emissions of NOX and SOX from the regenerator.
Emerging Trends in Fluidized Catalytic Cracking
327
Minimizing Detrimental Effects of Processing Residual Feeds
The proper choice of a feed injection system, regenerator, and atalyst are some of the key aspects of successful RFCC operation. An efficient feed injection system produces extremely small droplets hat vaporize quickly. Rapid vaporization minimizes the amount of on-vaporized hydrocarbons that block the active sites. An effective eed nozzle system must instantaneously vaporize and crack asphaltenes nd poly nuclear aromatics to lower boiling entities. The regenerator design, either single-stage or two-stage, should rovide uniform catalyst regeneration, increase flexibility for processing variety of feedstocks, and minimize thermal and hydrothermal eactivation of the catalyst. The catalyst design should be optimized to achieve the followng objectives: » • • • •
Low coke and gas production Efficient bottoms cracking Improved metals resistance Improved thermal and hydrothermal stability An active matrix and a low hydrogen transfer activity to convert the bottoms and minimize delta coke
REDUCING FCC EMISSIONS
The gaseous emissions from the FCC unit are CO, NOX, pariculates, and SOX, All are either locally or nationally regulated. Table 10-7 shows the current allowable limits of the EPA New Source Performance Standards (NSPS) for the emissions of these irborne pollutants. NSPS levels can be triggered by one of the ollowing conditions: • Construction of a new unit • Revamp of the regenerator, provided the modification costs are more than 50% of a comparable regenerator 8 Any capital modification of the unit that increases its emission rates
There is no national requirement limiting NOX emissions from the CC flue gas, but several state and regional agencies have imposed mits on their release. These emissions are directly proportional to
28
Fluid Catalytic Cracking Handbook Table 10-7 EPA's New Source Performance Standards (NSPS) for Gaseous Emissions from the FCC Regenerators
Source
Allowable Limits
Carbon monoxide (CO)*
Less than 500 ppmv in the flue gas
Nitrogen oxides (NOX)
None (local and regional only)
Participates**
A maximum of 1.0 pound of solids in the flue gas per 1,000 pounds of coke burned
Sulfur oxides (SO2 + SO3)*
Exempt if the feed sulfur is less than 0.30 wt% If there is no add-on control such as a wet gas scrubber, 9.8 kilograms of (SO2 + SO3) per 1,000 kilograms of coke burned. This is approximately equal to 500 ppmv. Add-on device: reduce (SO2 + SO3) by at least 90% or no more than 500 ppmv, whichever is less stringent.
Effective January 1984 *Effective June 1973
he quality of FCC stocks, operating conditions, catalyst type, and mechanical condition of the unit. Processing feeds that contain a high oncentration of residue, sulfur, nitrogen, and metals will release a greater amount of SOX, NOX, and particulates. Various technologies re available to reduce flue gas emissions.
Particulates
Electrostatic precipitators (ESP) and wet gas scrubbers (WGS) are widely used to remove particulates from the FCC flue gas. Both can ecover over 80% of filtrable solids. An ESP (Figure 10-6) is typically nstalled downstream of the flue gas heat recovery (prior to atmosheric discharge) to minimize particulate concentration. If both low articulate and low SOX requirements are to be met, a wet gas scrubber uch as Belco's (Figure 10-7) should be considered. If SOX removal
ORMER IER
EN
ROOF
COLLECTING SURFACE RAPPER
HIGH-VOLTAGE SYSTEM SUPPORT INSULATOR
RAPPER INSULATOR
DISCHARGE ELECTRODE RAPPER
Figure 10-6. Typical electrostatic precipitator (ESP),
HOPPER-
DISCHARGE ELECTRODE
INSULATOR COMPARTMENT,
BUS DUCT-
DO
•SIDE
Figure 10-7. Schematic of Belco scrubbing system (courtesy of Belco Corporation).
CULATING PUMP
SODA
Emerging Trends in Fluidized Catalytic Cracking
331
s not a prime objective, an ESP will be less expensive from the tandpoints of both initial capital and operating costs. In some cases, bag house system can be used instead of an ESP,
OX
Three methods are widely used to reduce SOX emissions from the CC flue gas: FCC feed pretreatment Catalyst additives Flue gas desulfurization
Feed hydrotreating or hydrocracking reduces SOX emissions and the ulfur content of FCC products. As discussed earlier in this chapter, many benefits are associated with FCC feed hydrotreating. It is mportant to note that most of the sulfur in a hydrotreated feed is in eavy organic compounds and will be concentrated in the decanted il and coke. Consequently, for a given sulfur in the feed, more SOX will be produced with hydrotreated feed. For refiners having low to moderate levels of SOX in their FCC flue as (less than 1,000 ppm), SOX additives are usually the most ecoomical method of reducing SOX emissions. These additives are njected separately into the regenerator. They capture SO3 in the egenerator (oxidizing atmosphere) and release sulfur as H2S in the reactor educing atmosphere). A reliable on-line SO2 analyzer will ensure that a ufficient quantity of additive is injected. Operating conditions of the egenerator, especially partial versus full combustion and excess oxygen evel, will greatly influence the additive's effectiveness. When processing high-sulfur feeds (greater than 1.0 wt%) or if the equired SOX reduction levels are greater than 80%, other capitalntensive desulfurization technologies must be considered. Several ue gas desulfurization technologies are available. Haldor Topsoe's WSA, United Engineers' Mgo., Exxon, and Belco (Figure 10-7) wet as scrubbing (WGS) are among the most widely used processes to emove SOX. The WGS process removes both SOX and particulates.
CO
The CO levels released from the regenerator flue gas operating ither in complete or partial combustion are normally less than 10
32
Fluid Catalytic Cracking Handbook
pm. For units operating in partial combustion, the flue gas must e sent to a CO boiler. For units operating in complete combustion, he concentration of CO largely depends on the operating conditions f the regenerator (mainly temperature and excess oxygen), the CO romoter level, and the efficiency of the air/spent catalyst distriution system.
NOX
NOX levels in the FCC flue gas typically range from 50-500 ppm. Nitrogen content of the feed, excess oxygen, regenerator residence ime, dense phase temperature, and CO promoter all influence the oncentration of NOX. In the regenerator, most of the NOX is formed as NO, with little N2O or NO2. About 90% of organic nitrogen in the spent catalyst is onverted to inorganic nitrogen, and a very small amount becomes NO. NO can be lowered by reducing excess oxygen and CO promoter. The resent platinum-based promoter oxidizes intermediates such as HCN nd NH3 to NO and decreases the reducing agent such as CO. To reduce nitrogen oxide, thermal and catalytic processes are availble. The thermal process is licensed by Exxon. NH3 or urea is injected nto the flue gas at an elevated temperature (-1600°F, 870°C); NOX s reduced to nitrogen. This process is applicable to FCC units that ave CO boilers. NOX can also be reduced over a catalyst at 500°F o 750°F (260°C to 400°C).
EMERGING DEVELPMENTS IN CATALYSTS, PROCESSES, AND HARDWARE
The FCC process has a long history of innovation and will continue o play a key role in the overall success of the refining industry. The ontinuing developments will primarily be in the areas of catalyst, rocess, and hardware technologies.
Catalyst
Since the mid-1960s, formulation of FCC catalysts has improved teadily. The focus of the research is in the following areas:
Emerging Trends in Fluidized Catalytic Cracking
* * • • * •
333
Improvement in zeolite quality Improvement in the catalyst's binder properties Increase in the quantity and choice of active matrix Customization of catalyst to the unit's objectives and constraints A widespread use of ZSM-5 or similar zeolite Improvements in the developments of catalyst additives for reducing gasoline sulfur and NOX emission
here has also been an ongoing trend to formulate a higher-quality eolite. Higher quality has been reflected in: • Greater silica-to-alumina (SAR) of zeolite. Greater SAR results in a zeolite that is more stable, yields more olefins, improves octane, and increases product selectivity. * Improved crystallinity by producing more uniform zeolite crystals, FCC catalyst manufacturers have greater control over the zeolite acid site distribution. In addition, there is an upward trend in the quantity of zeolite being included in the catalyst.
The selectivity and activity of the catalyst matrix will continue to mprove. The emphasis on bottoms cracking and steady reduction in he reaction residence time demands an increase in the quantity of ctive matrix. The improvements in the catalyst's binder properties will reduce he catalyst attrition rate; thus, lowering the flue gas stack opacity. his improvement allows refiners to use a "harder" catalyst without dversely affecting the catalyst's fluidization properties. Future catalyst formulation will be customized to meet the individual efiner's needs. Catalyst manufacturers will be tailoring catalysts to meet each refiner's requirements. The demand for ZSM-5 additives will increase because of their nherent ability to crack low-octane, straight chain olefins to C3 and C4 olefins and also to isomerize low-octane linear olefins to higher ctane branched olefins. Once ZSM-5's patent has expired, its use hould increase. Further developments in the effectiveness of the FCC gasoline sulfur eduction additives will allow a number of refiners to meet the equired reduction in gasoline sulfur without undertaking costly capital rojects. Additionally, improvements in the CO promoter additives will reduce NOX emissions when the promoter is used. Finally, other
34
Fluid Catalytic Cracking Handbook
ost-effective additives will be developed to not only reduce NOX missions, but also reduce catalyst related fouling in the regenerator ue gas heat recovery system.
Operating Conditions
FCC will still play a dominant role in producing cleaner-burning uels. The inherent flexibility of the process will allow refiners to meet he fuel reformulation requirements. With the anticipated growing emand for alky late and ethers, the FCC operating parameters will e adjusted to maximize production of propylene, isobutylene, and soamylene. The projected trend in operating conditions will be to a igher reactor temperature, a higher catalyst-to-oil ratio, a higher eaction mix temperature, and shorter catalyst contact time.
Technology Development
Since 1942, when the first FCC unit came onstream, new techologies have continuously evolved to maximize performance to meet he ever-changing product requirements and feedstock qualities. Future echnology development will remain dynamic. Examples of the new nd ongoing technologies aimed at enhancing the unit's operational nd mechanical performance, as well as complying with environmental egulations, are: • Reducing sulfur and aromatics in gasoline and distillate. • Minimizing disposal of equilibrium catalyst. • Minimizing catalyst back-mixing in the riser to minimize production of undesirable products. Redesign of the conventional riser for a down-flow of catalyst and vapors could virtually eliminate back-mixing. • Achieving an ultra-short catalyst-hydrocarbon contact time, designed to maximize olefins and gasoline yields while minimizing the bottoms yields. • Eliminating long dilute-phase residence time downstream of the riser to prevent recracking of hydrocarbon vapors in the reactor housing. • Improving feed and catalyst injection systems. • Improving spent catalyst distribution.
Emerging Trends in Fluidized Catalytic Cracking
335
• Improving mechanical reliability of the FCC reactor-regenerator components. • Increasing use of feed segregation to maximize production of light olefins. « Increasing use of riser quench to maximize the reaction mix temperature and to promote maximum vaporization of the feedstock. • Increasing use of catalyst additives to reduce gaseous emissions and to maximize light olefins. These are just some of the many challenges facing FCC operaons today.
SUMMARY
The United States refining industry is undergoing a restructuring hase. Refiners will continue to be under pressure and only the most fficient and profitable operations are going to survive. The survivors will be those who have some niche in the market place, have the ersatility to handle low-cost crude, meet product demand, and conorm to environmental regulations. FCC is one of the cheapest conversion processes. Its inherent exibility can assist a refiner in meeting changing product requirements in spite of the steady decline in feedstock quality. The U.S. Federal RFG program has imposed new challenges for the CC, particularly regarding the sulfur, aromatics, and olefin content of asoline. Various commercially proven technologies, along with evolving echnologies, will be available to comply with these new rules. The use of RFCC will continue to grow, particularly in regions of he world where atmospheric or vacuum residue contains low levels f contaminants. Careful regenerator and feed injection designs are mportant in ensuring a successful operation. Gaseous emissions (CO, NOX, SOX, particulates) have been regulated t local and national levels. The quantity of these emissions is directly elated to the quality of the FCC stocks, operating conditions, catalyst ype, and mechanical conditions of the unit. Processing heavy feeds will release a greater amount of SOX, NOX, and particulates. In conclusion, FCC has had a long history of innovations. New echnological developments will continue to emerge, optimizing its erformance. Its versatility and high degree of efficiency will continue o play a key role in meeting future market demands.
3§
Fluid Catalytic Cracking Handbook
REFERENCES
1. Mauleon, J. L. and Letzsch, W. S., "The Influence of Catalyst on the Resid FCCU Heat Balance," presented at Katalistik's 5th Annual FCC Symposium, Vienna, Austria, May 23-24, 1984. 2. A. W. Peters, G. Yaluris, G. D. Weatherbee, X. Zhao, "Origin and Control of NOX in the FCCU Regenerator," Grace Davison, Columbia, MD. 3. Davis, K., and Ritter, R. E., "FCC Catalyst Design Considerations for Resid Processing—Part 2," Grace Davison Catalagram, No. 78, 1988. 4. Hammershaimb, H. U., and Lomas, D. A., "Application of FCC Technology to Today's Refineries," presented at Katalistiks' 6th Annual FCC Symposium, Munich, Germany, May 22-23, 1985. 5. Kool, J. M., "Commercial Experience with Resid Cracking in Conventional FCC Units," presented at the 1984 Akzo Chemicals Symposium, 6. Hood, R., and Bonilla, J., "Residue Upgrading by Solvent Deasphalting and FCC," presented at the Stone & Webster 5th Annual Meeting, Dallas, Texas, October 12, 1993. 7. Dean, R. R., Kibble, P. W., and Brown, G. W., "Crude Oil Upgrading Utilizing Residual Oil Fluid Catalyst Cracking," presented at Katalistiks' 8th Annual FCC Symposium, Budapest, Hungary, June 1-4, 1987. 8. Johnson, T. E., "Resid FCC Regenerator Design," presented at the M.W. Kellogg Co. Refiing Technology seminar, Houston, Texas, February 9-10, 1995. 9. Letzsch, W., Mauleon, J. L. Jones, G., and Dean, R., "Advanced Residual Fluid Catalytic Cracking," presented at Katalistiks' 4th Annual FCC Symposium, Amsterdam, The Netherlands, May 18-19, 1983. 10. Elvin, F. J., and Krikorian, K. V., "The Key to Residue Cracking," presented at Katalistiks' 4th annual FCC Symposium, Amsterdam, The Netherlands, May 18-19, 1983. 11. Peeples, J. E., "The Clean Air Act, a Brave New World for Fuel Reformulation," Fuel Reformulation, Vol. 3, No. 6, November/December 1993. 2. Dharia, D., Brahn, M., and Letzsch, W., "Technologies for Reducing FCC Emissions," presented at Stone & Webster's 5th annual Refining Seminar, Dallas, Texas October 12, 1983. 3. Yergin, D. and Lindemer, K., "Refining Industry's Future," Fuel Reformulation, Vol. 3, No. 4, July/August 1993. 4. Perino, J. O., "Blending Control Upgrade Projects," Fuel Reformulation, Vol. 3, No. 4, July/August, 1993. 5. Clarke, R. H. and Ritz, G. P., "Method for the Analysis of Complex Mix of Oxygenates in Transportation Fuels," Fuel Reformulation, Vol. 3, No. 4, July/August, 1993.
Emerging Trends in Fluidized Catalytic Cracking
33?
6. Urizelman, G. H., "NOX," Fuel Reformulation, Vol. 1, No. 6, November/ December 1991. 7. Piel, W. J., and Thomas, R. X., "Oxygenates for Reformulated Gasoline," Hydrocarbon Processing, July 1990, pp. 68-73, 8. Hirshfeld, D. S. and Kolb, J., "Minimize the Cost of Producing Reformulated Gasoline," Fuel Reformulation, Vol. 4, No, 2 March/April 1994 9. Unzelman, G. H., "A Sticky Point for Refiners," Fuel Reformulation, Vol. 2, No. 4, July/August 1992. 0. Nocca, J. L., Forestiere, A., and Cosyns, J., "Diversify Process Strategies for Reformulated Gasoline," Fuel Reformulation, Vol. 4, No. 4, September/ October 1994. 1. Desai, P. H., Lee, S. L., Jonker, R. J., De Boer, ML, Vending, J., and Sarli, M. S., "Reduce Sulfur in FCC Gasoline," Fuel Reformulation, Vol. 4, No. 6, November/December 1994. 2. Sarathy, P. R., "Profit from Refinery Olefins," Fuel Reformulation, Vol. 3, No. 5, September/October 1993. 3. Hosteller, R. and Cain, M., BP Oil, private communication, 1995. 4. Reid, T. A., Akzo Nobel, private communication, 1995. 5. Swain, E. J., "U.S. Crude Slate Continues to Get Heavier, Higher in Sulfur," Oil & Gas Journal, January 9, 1995, pp. 37–42. 6. Dougan, T. J., Alkemade, V, Lakhampel, B., and Brock, L. T., "Advances in FCC Vanadium Tolerance," NPRA Annual Meeting, San Antonio, Texas, March 20, 1994, reprinted in Grace Davison Catalagram. 7. Cunic, J. D., Diener, R., and Ellis, E. G., Exxon Research and Engineering, "Scrubbing—Best Demonstrated Technology for FCC Emission Control," presented at NPRA Annual Meeting, San Antonio, Texas, 1990.
APPENDIX 1
Temperature Variation
of Liquid Vis
^HnnHHIHIIIHHIiniUIIIIIIMIHIIIIIIIIIIIIIIIIMIIIIIIIIIIIIIIIIHIIIIIIIIIIIIItllllilllilllUlllllilllllllllllllllMllfllliWIIItlll MnnitMIIIIIIIIIIIIIIIHIMIIIIIIIIIIItlllllllllllllllllllllllllllllllllllllillHieHlllllllillllltllllliSKUMIitHIUitilMfllHIIilllll
Source: U.S. Department of Commerce, adapted from ASTM D-342-39.
338
APPENDIX 2
Correction to Volumetric Average Boiling Point WABP C 80) F V
==• ""^ WABP O 60 3 F VABP
A8TM Diet, 10% - 90 % Slop*
339
APPENDIX 3
TOTAL Correlations
Aromatic Carbon Content: CA = -814.136 + 635.192 x RI(20) - 129.266 x SG + 0.1013 x MW - 0.340 x S - 6.872 x ln(v)
Hydrogen Content: H2 = 52.825 - 14.26 x RI(20) - 21.329 x SG - 0.0024 x MW - 0.052 x S + 0.757 x ln(v)
Molecular Weight: MW = 7.8312 x 10-3 x SG-0-0976 x AP°C1238
Refractive Index @ 20°C: RI(20) = 1 + 0.8447 x SG1-2056 x (VABPoc+273.16r)0557 x
Refractive Index @ 60°C: RK60) = 1 + 0.8156 x SG12392 x (VABP0(: + 273.16)-0.0576 x
ource: Dhulesia, H., "New Correlations Predict FCC Feed Characterization Paramters," Oil & Gas Journal, Jan. 13, 1986, pp. 51-54.
340
APPENDIX 4
n-d-M Correlations v = 2.5 x (RI20OC - 1.4750) - (d2()OC - 0.8510) 05 = (d2()OC - 0.8510) - 1.11 x (RI2fn, - 1.4750) If v is positive: %CA = 430 x v +
If v is negative: %CA = 670 x v +
3660
M
If 03 is positive: %CR = 8 2 0 x G J - 3 x S + 10,000/M 10,600 If 03 is negative: %CR = 1440 x 03 - 3S +
M
%CN = %CR — %CA %C_r = 100— %CRK
Average Number of Aromatic Rings per Molecule (RA): RA = 0.44 + 0.055 x M x v
If v is positive
R^ = 0.44 + 0.080 x M x v
If v is negative
Average Total Number of Rings per Molecule (RT): RT = 1.33 + 0.146 x M x (03 - 0.005 x S)
If 03 is positive
RN = RT — RA
RT = 1.33 + 0.180 x M x (03 - 0.005 x S)
If 05 is negative
Average Number of Napthene Rings per Molecule (RN): R
M
=
RT—RA
ource: ASTM Standard D-3238-80. Copyright ASTM. Used with permission.
341
APPENDIX 5
Estimation of Molecular Weight of Petroleum Ofts from Viscosity Measurements
40 50 60 70 80 90 00 10 20 30 40 50 60 70 80 90
Tabulation of H Function H
334 355 372 386 398 408 416 424 431 437 443 448 453 457 461 465
336 357 374 387 399 409 417 425 432 438 443 449 453 458 462 466
339 359 375 388 400 410 418 425 432 438 444 449 454 458 462 466
341 361 377 390 401 410 419 426 433 439 444 450 454 459 463 466
343 363 378 391 402 411 420 427 433 439 445 450 455 459 463 467
342
345 364 380 392 403 412 420 428 434 440 446 450 455 460 463 467
347 366 381 393 404 413 421 428 435 441 446 451 456 460 464 468
349 368 382 394 405 414 422 429 435 441 447 451 456 460 464 468
352 369 384 395 406 415 423 430 436 442 447 452 456 461 465 468
354 371 385 397 407 415 423 430 437 442 448 452 457 461 465 469
Molecular Weight of Petroleum Oils Viscosity-Molecular Weight Chart LINES OF CONSTANT 210*F (98,89*C) VISCOSITY, cST
500
5
400
300
too
)0
j/
400
500
600
RELATIVE MOLECULAR MASS
Source: ASTM Standard D-2502-92. Copyright ASTM. Used with permission.
343
APPENDIX 6
Kinematic Viscosity to Saybolt Universal Viscosity Equivalent Saybolt Universal Viscosity, Sus
inematic Viscosity, cSt
1.81 2.71 4.26 7.37 10.33 13.08 15.66 18.12 20.54 43.0 64.6 86.2 108.0 129.5 139.8 151.0 172.6 194.2 215.8
At 100°F
At 210°F
32.0 35.0 40.0 50.0 60.0 70.0 80.0 90.0 100.0 200.0 300.0 400.0 500.0 600.0 648.0 700.0 800.0 900.0 1000.0
32.2 35.2 40.3 50.3 60.4 70.5 80.5 90.6 100.7 202.0 302.0 402.0 504.0 604.0 652.0
xtracted from ASTM Method D-2161-87. Copyright ASTM. Used with permission.
344
APPENDIX 7
API Correlations Xr = a + b x (R.) + c x (VG) Xn = d + e x (R.) + f x (VG) Xn = g + h x (R.) + i x (VG)
Where constants vary with molecular weight range given below:
Constants a b c d e f g h j
Heavy Fractions 200 < MW < 600 +2.5737 +1.0133 -3.573 +2.464 -3.6701 +1.96312 -4.0377 +2.6568 +1.60988
. = Refractivity Intercept VGC = Viscosity Gravity Constant R, K
K ~-R i(20)
Where:
j(2())
'
= Refractive Index @ 20°C = Density @ 20°C
ource: Riazi, M. R., and Daubert, T. E., "Prediction of the Composition of Petroleum ractions," Ind. Eng. Chem. Process Dev., Vol. 19, No. 2, 1982, pp. 289-294.
345
46
Fluid Catalytic Cracking Handbook
VGC = SG ~ °-24 - °-022 x log(V210 ~ 35.5) 0.755
Where:
V = Say bolt Universal Viscosity @ 210°F in seconds
Refractive Index @ 20°C (68°F):
I = A x exp(B x MeABP + C x SG + D x MeABP x SG) x MeABPE x SGF
Constants A B C D E F
2.341 * 10~2 6.464 x IQ"4 5.144 -3.289 x 10-4 -0.407 -3.333
MW = a x exp(b x MeABP + c x SG + d x MeABP x SG) x MeABP6 x SGf
Where:
Constants a b c d e f
20.486 1.165 x 10~4 -7.787 1.1582 x 10-3 1.26807 4.98308
APPENDIX 8
Definitions of Fluidization Terms
Aeration. Any supplemental gas (air, steam, nitrogen, etc.) that increases fluidity of the catalyst. Angle of Internal Friction—a. Angle of internal friction, or angle of shear, is the angle of solid against solid. It is the angle at which a catalyst will flow on itself in the nonfluidized state. For an FCC catalyst, this is about 80°. Angle of Repose—p. The angle that the slope of a poured catalyst will make with the horizontal. For an FCC catalyst, this is typically 30°.
SoHdSurfaca
347
48
Fluid Catalytic Cracking Handbook
Apparent Bulk Density—ABD. The density of the catalyst at which it is shipped either in bulk volume or bags. It is density of the catalyst at minimum fluidization velocity. ed Density—pb. The average density of a fluidized bed of solid particles and gas. Bed density is mainly a function of gas velocity and, to a lesser extent, the temperature. Minimum Bubbling Velocity (Umb). The velocity at which discrete bubbles begin to form. Typical minimum bubbling velocity for an FCC catalyst is 0.03 ft/sec. Minimum Fluidization Velocity (Umf). The lowest velocity at which the full weight of catalyst is supported by the fluidization gas. It is the minimum gas velocity at which a packed bed of solid particles will begin to expand and behave as a fluid. For an FCC catalyst, the minimum fluidization velocity is about 0.02 ft/sec. article Density—p . The actual density of the solid particles taking into account any volume due to voids (pores) within the structure of the solid particles. Particle density is calculated as follows:
Po =
Skeletal density (Skeletal density x PV) + 1
ore Volume—PV. The volume of pores or voids in the catalyst particles. Ratio of Minimum Bubbling Velocity to Minimum Fluidization Velocity (Umb/Umf). This ratio can be calculated as follows: Umb
=
2300 x p°' 2< x n°"3 x exp° 7lteF
Umf= Where: pg = ji = F = dp = p = g =
d°8xg<"34x(pp-p!)0'34 gas density, kg/m3 gas viscosity, kg/m/sec fraction of fines less than 45 microns mean particle size particle density, kg/m3 gravitational constant = 9.81 m/sec2
The higher the ratio, the easier it is to fluidize the catalyst.
Definitions of Fluidization Terms
349
keletal Density—SD. The actual density of the pure solid materials that make up the individual catalyst particles. For an FCC catalyst, the skeletal density can be calculated as follows:
SD =
3.4
2.1
Where: Al = Alumina content of the catalyst, wt% Si = Silica content of the catalyst, wt% lip Factor. The ratio of vapor velocity to catalyst velocity. tick Slip Flow. The continuous sudden stoppage and resumption of catalyst flow in a standpipe. This is usually caused by underaeration. uperficial Velocity. The velocity of the gas through the vessel or pipe without any solids present. It is a volumetric flow rate of fluidization gas divided by the cross-sectional area.
APPENDIX 9
Conversion of ASTM 50% Point to TBP 50% Point Temperature
The following equation can be used to convert an ASTM D-86 50% emperature to a TBP 50% temperature. TBP (50) = 0.87180 x ASTM D-86 (50)1.0258
Where:
BP(50)
= true boiling point distillation temperature at 50 vol% distilled, °F ASTM D86(50) = observed ASTM D-86 distillation temperature at 50 vol% distilled, °F
xample:
Given ASTM D-86(50) = 547°F, determine TBP 50% temperature:
BP(50) = 0.87180 x (547) 1.058 BP(50) = 561°F
ource: Daubert, T. E., "Petroleum Fraction Distillation Interconversions," Hydrocarbon rocessing, September 1994, pp. 75-78.
350
APPENDIX 10
Determination of TBP Cut Points from ASTM D-86
The difference between adjacent TBP cut points can be determined y the following equation: Y. = A X B !
I
Where:
. Xi
= difference in TBP distillation between two cut points, °F = observed difference in ASTM D-86 distillation between two cut points, °F A,B = constants varying for cut points ranges, shown in the following table: i
Cut Point Range
A
B
1
100%–90% 90%-70% 70%–50% 50%–30% 30%– 10% 10%–0%
0.11798 3.0419 2.5282 3.0305 4.9004 7.4012
1 .6606 0.75497 0.820072 0.80076 0.71644 0.60244
7
3 4 5 6
ource: Daubed, T. E., "Petroleum Fraction Distillation Interconversions," Hydrocarbon rocessing, September 1994, pp. 75-78.
351
52
BP BP BP BP BP BP
Fluid Catalytic Cracking Handbook (0) (10) (30) (70) (90) (100)
= TBP(50) = TBP(50) = TBP(50) = TBP(50) = TBP(50) = TBP(50)
- Y4 - Y5 - Y6 - Y4 - Y~ - Y4 + Y3 + Y3 + Y2 + Y" + Y, + Y,
APPENDIX 11
Nominal Pipe Sizes Identification Steel
Wall Thickness (t) Inches
Inside Diameter
Sched. No.
Stainless Steel Sched. No.
STD
xs
40 80
40S 80S
.088 .119
.364 .302
0.675
STD XS
40 80
40S SOS
.091 .126
.493 .423
0.840
STD XS —
xxs
40 80 160 —
40S 80S .__ —
.109 .147 .187 .294
.622 .546 .466 .252
ominal Pipe Size nches
Outside Diameter Inches
Iron Pipe Size
1/4
0.540
3/8 1/2
(d) Inches
3/4
1 .050
STD XS — XXS
40 80 160 —
40S 80S — —
.113 .154 .219 .308
.824 .742 .612 .434
1
1.315
STD XS — XXS
40 80 160 —
40S 80S — —
.133 .179 .250 .358
1.049 .957 .815 .599
1-1/4
1.660
STD XS — XXS
40 80 160 —
40S 80S _ —
.140 .191 .250 .382
1.380 1 .278 1 . 1 60 .896
1-1/2
1.900
STD XS — XXS
40 80 160 —
40S 80S — —
.145 .200 .281 .400
1.610 1.500 1.338 1.100
353
54
Fluid Catalytic Cracking Handbook Identification
Siteei
Wall Thickness
Sched.
Stainless Steel Sched.
(t)
(d)
No.
No.
Inches
Inches
40 80 160 —
40S 80S
.154 .218 .344 .436
2.067 1.939 1.687 1.503
40 80 160 —
40S 80S — —
.203 .276 .375 .552
2.469 2.323 2.125 1.771
xxs
40 80 160 —
40S 80S — —
.216 .300 .438 .600
3.068 2.900 2.624 2.300
Nominal Pipe Size Inches
Outside Diameter Inches
Iron Pipe Size
2
2.375
STD XS —
xxs 2-1/2
2.875
STD XS
— XXS 3
3.500
STD XS —
— —
inside Diameter
3-1/2
4.000
STD XS
40 80
40S 80S
.226 .318
3.548 3.364
4
4.5
STD XS
40 80 120 160
40S 80S — — —
.237 .337 .438 .531 .674
4.0.26 3.826 3.624 3.438 3.152
xxs
40 80 120 160 —
40S 80S — — —
.258 .375 .500 .625 .750
5.047 4.813 4.563 4.313 4.063
STD XS — — XXS
40 80 120 160 —
40S 80S — — —
.280 .432 .562 .719 .864
6.065 5.761 5.501 5. 187 4.897
—
—
xxs 5
5.563
STD XS — —
6
6.625
—
Conversion Factors atmosphere (atm) atmosphere (atm)
atmosphere (atm) bar barrel(bbl), 42 U.S.gal. barrel/day Btu Btu Btu/hr Btu/hr Btu/hr Btu/lb Btu/lb Btu/lb—°F Btu/lb—°F Btu/hr-ft2—°F egree Fahrenheit (°F) egree Kelvin (°K) egree Rankine (°R) foot (ft or ') foot (ft or ') gallon (gal), U.S. gallon (gal), U.S. as constant (°R) as constant (°R) horsepower (hp) inch (in. or ") inch (in. or ") pound (Ib), weight Ib/ft2-sec lb/ft3 lb/ft ! lb/ft'
= 14.696 Ib (force)/in2, (absolute) = 1.013 x 105 Newton/square meter (N/m 2 ) = 1.013 bar = 105 pascal = 0.159 cubic meter (m3) = 6.625 x 10-' mYhr = 1,055 joule (J) = 252.0 calories (cal) = 3.93 x 10^ horsepower (hp) = 0.252 kcal/hr = 0.29307 Watts = 0.556 calorie/gram (cal/g) = 2.326 joules/gram (J/g) = 4.186 joules/gram—°C = 1.0 calorie/gram—°C = 4.882 kg-cal/hr-m2—°C = 1.8 x °C + 32 = °C + 273 = 460 + °F = 12 inches (in or ") = 0.3048 meter (m) = 3.785 x liters = 3.785 x 10~3 cubic meter (m3) = 10.73 (psia) x (ft3)/(lb-mole) x (°R) = 8314 N/m2 x mVkg-mole x °K = 746 watts (W) = 2.54 centimeters (cm) = 0.0254 meter (m) = 453.6 grams (g) = 4.8761 kg/m2-sec = 0.016 gram/cubic centimeter (g/cm3) = 0.016 gram/milliliter (g/ml) = 16.018 kilogram/cubic meter (kg/m3)
355
56
Fluid Catalytic Cracking Handbook
Ib /gal (U.S.) Ib (force)/in2 (psi) Ib (force)/in2 (psi) Ib (force)/in2 (psi) mile ton (short) ton (short) ton (metric) ton (long) toe (long)
= 0.1198 g/cm3 = 0.0689 bars = 0.0680 atmospheres (atm) = 0.0703 kg/cm2 = L61 kilometers = 2,000 pounds (Ibs) = 907.2 kilograms = 1000.00 kilograms = 1016.0 kilograms = 2,240 Ibs
Glossary
Advance Process Control (APC) is a mechanism which manipulates regulatory controls toward more optimum unit operation. Afterburn is the combustion of carbon monoxide (CO) to carbon dioxide (CO2) in the dilute phase or in the cyclones of the regenerator. Alkylation is one of the refining processes in which light olefin molecules are reacted with isobutane (in the presence of either sulfuric or hydrofluoric acid) to produce a "desirable" gasoline component called alkylate. American Society of Testing and Materials (ASTM) is the organization that develops analytical tests and procedures to facilitate commerce. Aniline Point is the minimum temperature for complete miscibility of equal volumes of aniline and the hydrocarbon sample. In cat cracking, aniline solution is used to determine aromaticity of FCC feedstocks. Aromaticity increases with reducing aniline point.
Anti-dumping is a federal requirement to prevent refiners from blending the "good" gasoline components in "reformulated gasoline" and the "bad" components in conventional gasoline.
Antimony is a metal, in either hydrocarbon or aqueous solution, commonly injected into the fresh feed to passivate nickel.
API Gravity is an "artificial" scale of liquid gravity defined by: (141.5/SG) 131.5. The scale was developed for water = 10. The main advantage of using °API gravity is that it magnifies small changes in liquid density.
Apparent Bulk Density (ABD) is the density of catalyst as measured, "loosely compacted" in a specified container.
Average Particle Size (APS) is the weighted average diameter of a catalyst.
Back-Mixing is the phenomena by which the catalyst travels more slowly up the riser than the hydrocarbon vapors.
Basic Nitrogen is the nitrogen compounds in the FCC feed that react with the catalyst acid sites, thereby reducing the catalyst's activity and selectivity. Beta-Scission is splitting of the C-C bond two bonds away from the positively charged carbon atom.
357
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Fluid Catalytic Cracking Handbook
inder is the material used in the FCC catalyst to bind the matrix and zeolite components into a single homogeneous particle. California Air Resources Board (CARB) is a state agency which regulates and sets standards for air quality and emissions of various pollutants. Catalyst Activity is the conversion of feed (gas oils) to gasoline, lighter products, and coke in the MAT laboratory. Carbeniurn Ion is a positively charged (R-CH2+) ion that is formed from a positive charge to an olefin and/or by removing a hydrogen and two electrons from a paraffin molecule. Carbocation is a generic term for a positively charged carbon ion. Carbocation is further subdivided into carbenium and carbonium ions. Carbonium Ion is a positively charged (CH5+) ion which is formed by adding a hydrogen ion (H+) to a paraffin. Cat/Oil Ratio is the weight ratio of regenerated catalyst to the fresh feed to the riser feed injection zone. Catalyst Cooler is a heat exchanger that removes heat from the regenerator through steam generation. Cetane Number is a numerical indication of a fuel's (kerosene, diesel, heating oil) ignition quality. Cetane number is measured in a singlecylinder engine, whereas cetane index is a calculated value. Coke is a hydrogen deficient residue left on the catalyst as a by-product of catalytic reactions. Coke Factor is coke-forming characteristics of the equilibrium catalyst relative to coke-forming characteristics of a standard catalyst at the same conversion, Coke (Carbon) on Regenerated Catalyst (CRC) is the level of residual of carbon remaining on the catalyst when the catalyst exits the regenerator. Coke Yield is the amount of coke the unit produces to stay in heat balance, usually expressed as percent of feed. Conradson Carbon, or Concarbon, is a standard test to determine the level of carbon residue present in a heavy oil feed. Conventional Gasoline is a non-RFG gasoline that meets exhaust benzene, sulfur, olefins and T90 specifications. Conversion is often defined as the percentage of fresh feed cracked to gasoline, lighter products, and coke. Raw conversion is calculated by subtracting the volume or weight percent of the FCC products (based on fresh feed) heavier than gasoline from 100, or: Conversion = 100 – (LCO + HCO + DO) vol% or wt%
Glossary
359
Cyclone is a centrifugal separator which collects and removes particulates from gases. D-86 is a common ASTM test method that measures the boiling point of "light" liquid hydrocarbons at various volume percent fractions. The sample is distilled at atmospheric pressure, provided its final boiling point (end point) is less than 750°F. D–1160 is an ASTM method that measures the boiling point of "heavy" liquid hydrocarbons at various volume percent fractions. The sample is distilled under vacuum (results are converted to atmospheric pressure). The application of D1160 is limited to a maximum final boiling point of about 1000°F. Decanted Oil, Slurry, Clarified Oil, or Bottoms is the heaviest and often the lowest priced liquid product from a cat cracker. Delta Coke is the difference between the coke content of the spent catalyst and the coke content of the regenerated catalyst. Numerical value of delta coke is calculated from: Delta coke = coke yield (wt%/catalyst-to-oil ratio)
Dense Phase is the region where the bulk of the fluidized catalyst is maintained. Dilute Phase is the region above the dense phase which has a substantially lower catalyst concentration. Dipleg is the part of a cyclone separator that provides a barometric seal between the cyclone inlet and the cyclone solid outlet. Distributive Control System (DCS) is a digital control system that has a distributive architecture where different control functions are implemented in specialized controllers. Dynamic Activity is an indication of conversion per unit coke using data from the MAT laboratory. Equilibrium Catalyst (E-cat) is the regenerated catalyst circulating from the reactor to the regenerator. Exhaust Benzene is the amount of benzene toxins released. Exhaust benzene is a function of aromatics and benzene. Expansion Joint is a mechanical assembly designed to eliminate large thermal stresses in the piping. Faujasite is a naturally occurring mineral, having a specific crystalline, alumina-silicate structure, used in the manufacturing of the FCC catalyst. Zeolite faujasite is a synthetic form of the mineral. Filler is the inactive component of the FCC catalyst.
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Flapper Valve, Trickle Valve, or Check Valve is often attached to the end of a dipleg to minimize gas leakage up the dipleg and also catalyst losses during the unit start-up. Free Radical is an uncharged molecule formed in the initial step of thermal cracking. Free radicals are very reactive and short-lived. Gas Factor is the hydrogen and lighter gas-producing (C1-C4) characteristics of the equilibrium catalyst relative to the hydrogen and lighter gasproducing characteristics of some standard catalyst at the same conversion. Heat of Cracking is the amount of energy required to convert FCC feed to the desired products.
Hydrogen Transfer is the secondary reaction that converts olefins (predominantly iso-olefins) into paraffins while extracting hydrogen from larger, more hydrogen-deficient molecules. K Factor is an index designed to balance density and boiling point such that it relates solely to hydrogen content of the hydrocarbon. Microactivity Test (MAT) is a small, packed-bed catalytic cracking test that measures activity and selectivity of a feedstock-catalyst combination. Matrix is a substrate in which the zeolite is imbedded in the cracking catalyst. Matrix is often used as a term for the active, non-zeolitic component of the FCC catalyst. Methyl Tertiary Butyl Ether (MTBE) is an ether added to gasoline to improve its octane and reduce air pollution.
Mix Zone Temperature is the theoretical equilibrium temperature between the regenerated catalyst and the uncracked vaporized feed at the bottom of the riser.
Molecular Sieve is a term applied to zeolite. Zeolite exhibits shape selectivity and hydrocarbon absorptions.
Motor Octane Number (MON) is a quantitative measure of a fuel to "knocking," simulating the fuel's performance under severe operating conditions (at 900 rpm and at SOOT).
-d-M is an ASTM method that estimates chemical composition of a liquid stream.
Octane Barrel Yield, as used in the FCC, is defined as (RON + MON)/2 times the gasoline yield. Oxygenate is an oxygen-containing hydrocarbon. The term is used for oxygen-containing molecules blended into gasoline to improve its combustion characteristics.
Glossary
361
article Density is the actual density of solid particles, taking into account volume due to any voids (pores) within the structure of the solid particles. article Size Distribution (PSD) is the particle size fractions of the FCC catalyst expressed as percent through a given sized hole, Plenum is a means of collecting gases from multiple sets of cyclones before they are exhausted from the unit. Pore Diameter is an estimate of the average pore size of the catalyst. Pore Volume is the open space in the FCC catalyst, generally measured by mercury, nitrogen, or water. Mercury is used to measure large pores, nitrogen measures small pores, and water is used for both. Ramsbottom, similar to Conradson Carbon, is a quantitative indication of carbon residue of a sample. Rare Earth is a generic name used for the 14 metallic elements of the lanthanide series used in the manufacturing of FCC catalyst to improve stability, activity, and gasoline selectivity of the zeolite. Reformulated Gasoline (RFG) is the gasoline sold in some ozone nonattainment metropolitan areas designed to reduce ozone and other air pollutants. Refractive Index, similar to aniline point, is a quantitative indication of a sample's aromaticity. Refractory is a cement-like material used to stand abrasion and erosion. Reid Vapor Pressure (RVP) is gasoline vapor pressure at 100°F. Resid refers to a process, such as resid cat cracking, that upgrades residual oil. Residue is the residual material from the processing of raw crude (for example, vacuum residue and not vacuum resid). Research Octane Number (RON) is a quantitative measure of a fuel to "knocking," simulating the fuel's performance under low engine severity (at 600 rpm and 120°F). Riser is a vertical "pipe" where virtually all FCC reactions take place. Selectivity is the ratio of yield to conversion for the "desired" products. Silica Oxide to Alumina Oxide Ratio (SAR) is used to describe the framework composition of zeolite. keletal Density is the actual density of the pure solid materials that make up individual particles. lide Valve or Plug Valve is a valve used to regulate the flow of catalyst between reactor and regenerator. lip Factor is the ratio of catalyst residence time to hydrocarbon vapors residence time in the riser.
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oda Y Zeolite is a "crystallized" form of Y-faujasite before any Ion exchanges occur. pent Catalyst is the coke-laden catalyst in the stripper. tandpipe is a means of conveying the catalyst between reactor and regenerator, tick-Slip Flow is erratic circulation caused when the catalyst packs and bridges across the standpipe. True Boiling Point (TBP) is the distillation separation which has characteristics of 15 different theoretical plates at 5 to 1 reflux ratio. Ultra Stable Y is a hydrothermally treated Y-faujasite which has a unit cell size at or below 24.50 °A and exhibits superior hydrothermal stability over Soda Y faujasite. Unit Cell Size (UCS) is an indirect measure of active sites and SAR in the zeoliteZeolite is a synthetic crystalline alumina-silicate material used in the manufacturing of FCC catalyst.
Index complex model, 314 Carbenium ion, 131 Carbon black feedstock, 199 Carbon residue. See Feed. Carbonium ion, 131 Catalyst, 84, 296 activity, 20 CRC and, 111 additives. See each additive. aging, 109 alumina balance, 107, 114 binder, 95 changeover, 113 chemical properties, 107 circulation, 166, 236, 293 limited, 236 erratic, 243 coke level, 13 components, 84 CRC, 109 developments, 212, 332 distributor, 119 equilibrium (E-cat), 102 evaluation, 115 filler, 95 fluidization, 169, 221, 241 heat capacity, 164 history, 128 hopper, 170 design guidelines, 221 losses, 21, 244 management, 21, 109 manufacturing, 96 Englehard, 99
Additives, See each. Advanced process control, 179 Afterburn, 18, 259 Air distributor, 15, 119, 204 design guidelines, 226 debottlenecking, 295 Alkaline earth metals. See Sodium. Alkylation, 322 Alpha-scission, 132 Amine tyreating, 34 Ammonia, 29 Ammonium bisulfide, 29 Aniline point. See Feed. Antimony, 122 API gravity. See Feed. Aromatics, 43, 75 polynuclear aromatics, 43 in fuels, 319 Asphaltene, 53
enzene. See Gasoline. eta-scission, 132 inder, 95 MCI, 199 ottoms. See Slurry. romine number. See Feed.
AAA, 187, 311 simple model, 313
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atalyst (continued) matrix, 93 properties, 99 apparent bulk density (ABD), 106 coke factor, 104 gas factor, 104 microactivity (MAT), 104, 111 particle size distribution (PSD), 100, 107 pore volume, 107 surface area, 100, 104 rare-earth, 89, 101 and activity, 102 and octane, 93, 190, 192 and yield, 93 and hydrogen transfer, 134 separation, 17 sodium, 92, 100, 108. See also Zeolite, octane and, 94 atalyst cooler, 204 atalyst-to-oil ratio, 162, 186 austic treating, 31, 36 BFS, 199 etane, 197 O boiler, 17 O promoter, 20, 117 oke, 135 calculation, 149 delta coke, 201 hydrocarbons, 13 sulfur, 62 yield, 135, 149, 200, 324 factor, 104 oking/fouling, 248, 252 oking, delayed, 199 ombustion modes, 18, 117 onradson, 52 ontrol system, 177, 304 onversion apparent, 156
definition, 142 nitrogen and, 55 troubleshooting, 264 Copper, 68 Correlations. See Feed. CRC, 109 Cyanide, 29, 56, 260 Cyclones, 9, 17, 204, 225 design guidelines, 229 flapper valve, 286
D D-1160, 48 D-2887 (SIMDIS), 48 D-86, 47 Debottlenecking, 276 Debutanizer, 28, 275 debottlenecking, 304 Decant oil. See Slurry. Delta coke, 201 Distillation. See Feed. Dry gas, 183 troubleshooting, 268
Economics, 182, 202 Equilibrium catalyst (E-cat), 22, 102 activity, 112 analysis, 102 ESP, 329 Expander, 18, 263 Expansion joints, 204 design guidelines, 230
Feed aniline point, 49 API gravity, 45 conversion to S.G., 45 octane and, 188 temperature correction, 46
bromine number, 50 carbon residue, 52 characterization, 40 coking tendency, 52, 324 contaminants, 54, 324 correlations, 68 API, 78 aromatic content, 75 hydrogen, 74 K-factor, 69 molecular weight, 74 n-d-M, 75 refractive index, 74 TOTAL, 74 UOP, 72 distillation, 47 hydrotreating, 81, 310 injection system, 201, 214 metals, 62 nozzles, 214, 292 physical properties, 45 preheat, 6, 278 refractive index, 49 residue, 323 split feed injection, 185 sulfur, 58 viscosity, 51 low reversal, 251 prevention, 254 shutdown matrix, 257 lue gas, 15, 296 CO, 331 emissions, 327 heat recovery, 17 NOX, 310, 332 particulates, 328 SOx, 118, 331 ouling. See Coking.
as factor, 104 as plant, 25
Index debottlenecking, 297 fouling/corrosion, 29, 56 Gasoline aromatics, 319 benzene, 190, 311, 319 end point, 188, 310 MTBE, 322 NOX and, 310 octane, 188 troubleshooting, 272 ZSM-5 and, 121 olefins, 308, 321 reformulated (RFG), 308 sulfur, 190, 193, 194, 315 sweetening, 37 vapor pressure, 308 VOC's, 308 yield, 121, 186 troubleshooting, 270 Gasoline splitter, 29
H H2S, 29, 34, 58 HCO, 24, 198 Heat balance, 158 Heptane insoluble, 53 Hydrogen, 135 in coke, 13, 166 from nickel, 64, 135 Hydrogen blistering, 29, 260 Hydrogen cyanide, 29 control, 260 Hydrogen transfer, 134, 184 Hydrotreating, 81 I
Iron, 68
K factor, 69
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Fluid Catalytic Cracking Handbook
CD, 24 quality, 196 quench, 291 yield, 195 icensors elco, 330 ABB Lummus, 284 Exxon Research and Engineering (ER&E), 2 Kellogg Brown & Root—KBR, 284 Phillips Petroleum, 122 Stone & Webster, 3, 288, 291 UOP, 3, 72, 287 PG, 25 olefin content, 121, 184 recovery, 302 treating, 36 yield, 121, 183 ZSM-5, 121
M
Main fractionator, 22 debottlenecking, 297 pool quench, 298 Magna-Cat, 122 MAT, 105 Material balance, 140 Matrix, 93 active, 95 and slurry yield, 95 and octane, 190 Mercaptans, 31 Metals, 62 passivation, 122 activity indexes, 63 balance, 108 copper, 68 feed hydrotreating, 81 iron, 68
Micro-method, 53
N Naphthenes, 42 Nickel, 63, 108 dehydrogenation, 135 and hydrogen, 64 passivation, 122 Nitrogen basic, 54 effects, 56 and conversion, 55 NOx.See Flue gas. NSPS, 118, 328
0 Olefins, 42 Operating constraints, 278 Orifice chamber, 17 Oxygen enrichment, 204 Oxygenates, 311
Paraffins, 41 and K-factor, 68 Power recovery, 17 troubleshooting, 263 Pressure balance, 166 Primary absorber, 27, 301 Process control, 177, 304 Advanced, 179
Ramsbottom, 52 Reactions, 125 heat of reaction, 162 catalytic cracking, 128 mechanism, 131 coking, 135
dehydrogenation, 135 hydrogen transfer, 134 isomerization, 133 thermal cracking, 126, 283 thermodynamics, 136 Reactor, 7 component yields, 153 cyclones, 7 design guidelines, 218 effluent sampling, 142 heat balance, 162 material balance, 140, 145 mechanical limitations, 282 vapor quench, 288 Refractive index. See Feed. Refractory, 230 Regeneration modes, 18 Regenerator, 13, 296 afterburn, 18, 259 catalyst cooler, 204 cyclones, 17 effect on vanadium, 66 heat balance, 160 high temperature, 20, 256 mechanical constraints, 283 pressure balance, 166 residue feed and, 323 transport disengaging height, 17 Residue feed, 323 Revamp considerations, 206 Riser, 7 pressure drop, 171 design guidelines, 215 lift zone, 8, 214, 217 termination (RTD), 9, 216, 283
hutdown matrix, 257 lide valve, 15, 204, 222 design guidelines, 224 low differential, 178, 294
Index
3§7
pressure balance, 171 Slurry, 23, 198 Sodium, 67, 92 catalyst and, 92 chloride and, 68 octane and, 94, 191 sources, 67 vanadium and, 66 Sour water, 31 SOX. See Flue gas. SOX additive, 118 Sponge oil absorber, 27 Standpipe, 15, 170, 175, 221, 240 debottlenecking, 294 design guidelines, 223 pressure balance, 170, 175 Steam to reactor, 140 to stripper, 219, 293 Stress corrosion cracking (SCC), 29 Stripper/de-ethanizer, 28, 301 Stripper (catalyst), 11, 293 debottlenecking, 301 design guidelines, 218 pressure balance, 171 steam distributor, 219 Sulfur. See also Gasoline. feed, 58 SOX additive, 118 effect of hydrotreating, 59 distribution in products, 58, 193
Tail gas. See Dry gas. Test run, 144 Thermal cracking. See Reactions. Third stage separator, 18 Treating amine, 34 caustic, 31 Troubleshooting, 234
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Fluid Catalytic Cracking Handbook
Vanadium, 65, 108, 325 CRC and, 66, 326 passivation, 122 sodium and, 66 iscosity. See Feed. VOC emissions, 308
W
Water wash, 29, 260 Wet gas compressor, 25 debottlenecking, 300
Z Zeolite, 96 chemistry, 85 development, 129 manufacture, 96 octane and, 192 properties, 88 amorphous catalyst vs., 130 rare earth, 89 silica-alumina ratio (SAR), 89 unit cdl
size (UCS)j 8g
ZSM-5, 87, 120, 29, 184, 333
About the Author
Since 1977, Reza Sadeghbeigi has been involved in various aspects of cat cracking at major and independent oil companies focusing on technical services, troubleshooting, process design, and project management—including major FCC revamps. Mr. Sadeghbeigi received his B.S. in chemical engineering from owa State University and his M.S. from Oklahoma State University. He is a registered professional engineer in Texas and Louisiana and s president of RMS Engineering, Inc. (RMS). RMS is based in Houston, TX, and is an independent engineering firm supplying engineering services to the refining industry in the area of fluid catalytic cracking (FCC) and related processes, since 1994. Should you have any questions or comments on this book or FCC's in general, please feel free to contact Reza at (281) 333-0464 (U.S.) or by e-mail [email protected]).
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