RFCC Process Technology Manual

UOP Fluid Catalytic Cracking Process Process Technology Manual

ORPIC Sohar, Oman September 2012

– LIMITED DISTRIBUTION – This material is UOP LLC technical information of a confidential nature for use only by personnel within your organization requiring the information. The material shall not be reproduced in any manner or distributed for any purpose whatsoever except by written permission of UOP LLC and except as authorized under agreements with UOP LLC.

157048 Table of Contents Page 1

FCC PROCESS TECHNOLOGY

TABLE OF CONTENTS I.

INTRODUCTION

II.

PROCESS FLOW Reactor Regenerator Main Column Gas Concentration and Recovery

III.

PROCESS CONTROL Reactor Regenerator Main Column Gas Concentration

IV.

EQUIPMENT Process Equipment and Its Use Metallurgical Corrosion

V.

FLUIDIZED SOLIDS Theory Applications to Fluid Catalytic Cracking

VI.

CATALYST History Modern FCC Catalysts Time and Temperature Effects Poisons Catalyst Management Catalyst Properties and Testing

VII.

PROCESS VARIABLES Reactor and Regenerator Process Variables Feedstock

VIII.

PROCESS CALCULATIONS FCC Flow Corrections and Mass Balance Liquid Product Cutpoint Corrections Reactor and Regenerator Heat Balance FCC Unit Mechanical Summaries Additional FCC Unit Calculations

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IX.

FEED AND PRODUCT TREATING Feed Treating Product Treating – Reasons and Methods

X.

ANALYTICAL METHODS Minimum Sample Size Typical Sampling Schedule Outline of FCCU Laboratory Methods

XI.

PROCEDURES Refractory Dryout Startup Shutdown Emergencies Catalyst Handling FCC Unit Evaluation

XII.

SAFETY General Additional Safety Precautions for Entering Vessels High Temperature Problems Chemical Hazards

XIII.

ENVIRONMENTAL Emissions Sources and Solutions

157048 Introduction Page 1

Introduction UOP Company History For more than 80 years, UOP has been one of the world’s leading licensors of new and innovative technology. Today, UOP continues in this role with 30 offices on 4 continents and 9 manufacturing facilities worldwide. UOP currently licenses and designs more than 60 different processes and has a total of 5,500 units licensed worldwide. For the last 50 years, the fluid catalytic cracking (FCC) process has been an important and successful part of UOP's licensing activities.

The Early Years UOP was founded in 1914 as the National Hydrocarbon Company on the strength of patent rights developed from the pioneering work of Jesse A. Dubbs, a California inventor. The company was financed by a noted Chicagoan, J. Ogden Armour. In 1915, the company name was changed to Universal Oil Products Company. From the beginning, the goal of the company was to develop and commercialize technology for license to the petroleum refining industry. Under the direction of C. P. (Carbon Petroleum) Dubbs, son of Jesse Dubbs, research and development work continued at the company's small site near Independence, Kansas, where the famous Dubbs Thermal Cracking process was successfully demonstrated in 1919. The then-revolutionary process became the foundation of UOP's rapid growth and its early worldwide recognition by the industry. The early period of growth was ably directed by its president, Hiram J. Halle, and by Dr. Gustav Egloff, one of the world’s leading petroleum chemists. In 1931, UOP established its headquarters in Chicago and its research laboratories in nearby Riverside, Illinois. That same year the ownership of UOP passed to a consortium of its major licensees, led by Shell and Standard Oil of California.

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During this stage, the company benefited immensely by the addition to its research staff of Prof. Vladimir Ipatieff, a famous Russian scientist known internationally for his work in high-pressure catalysis. His contributions in catalytic chemistry gave UOP a position of leadership in the development of catalysis as applied to petroleum processing. The first project of Ipatieff and his research team was catalytic polymerization. Other eminent scientists were also attracted to UOP’s research center in Riverside during this period. With the outbreak of World War II, UOP scientists and engineers focused their knowledge and talents on developing new catalytic processes, notably alkylation that helped meet wartime energy requirements, especially for aviation fuel. UOP also cooperated with other companies to develop the FCC process. In 1944, the owners of UOP divested themselves of their holdings in the company, and UOP’s stock was placed in trust. The American Chemical Society was named as the beneficiary. Thus, the Petroleum Research Fund was created with the understanding that income from the trust was to be used for advanced scientific education and fundamental research in the petroleum field. In spite of some financial and legal setbacks suffered by UOP during this period, strong management succeeded in steering the company back to its original course: taking creative research from concept to commercial reality. UOP was recognized as a company employing the world’s most knowledgeable scientific and technical personnel, who understood petroleum refining and the need for improved processing methods and techniques. In 1949, UOP's research staff developed a radically different refining process that used a catalyst containing platinum. Called the Platforming™ process, it revolutionized the art of reforming to produce gasoline with substantially improved octane number. The process was also instrumental in making benzene available in a quality and quantity never before realized on a commercial scale. With the Platforming™ process and other innovative processes, UOP became a vital contributor to the emergence and growth of the petrochemical industry.

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In the early 1950s, UOP also began to manufacture its own proprietary catalysts and a variety of refining chemicals at a newly constructed plant in Shreveport, Louisiana. Later UOP built manufacturing plants at McCook, Illinois; Brimsdown, U.K.; and other locations. In 1952, UOP moved its headquarters and engineering activities to Des Plaines, a suburb of Chicago. Soon after, the construction of a new research center at the same location was begun.

The Recent Era In 1959, UOP assumed its fourth different corporate form when it was sold to the public for the first time in its history. As a publicly owned company, UOP entered a new era marked by growth and diversification. The 1960s saw UOP grow from essentially a process-licensing company to a diversified corporation through many acquisitions and mergers with other companies. By 1975, UOP Inc. included more than 20 different divisions involved in such areas as aerospace and automotive technology. During the 1960s and 1970s, UOP's tradition of innovative process development and commercialization continued with the licensing of the first Sorbex™ simulated moving-bed countercurrent adsorption process in 1961 and the introduction of UOP's CCR Platforming™ process early in the 1970s. In 1975, Signal Companies Inc. acquired 50.5% of UOP and in 1978 acquired the remaining 49.5%, making UOP a wholly owned subsidiary of the company. When the Signal Companies merged with Allied Corporation in 1985, UOP Inc. became a subsidiary of Allied-Signal Inc. As the result of reorganizations and restructuring by its parent companies during the 1980s, UOP’s business scope was refocused on the development and licensing of process technology and the marketing of products associated with its licensing activities. Of the 20 different divisions, only the Process Division and UOP Management Services remain in the present UOP.

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In 1988, Allied-Signal entered into an agreement with Union Carbide Corporation that resulted in the creation of a unique joint venture company called simply UOP. The new UOP combined the resources of Allied-Signal’s UOP Inc. with the Catalysts, Adsorbents and Process Systems (CAPS) Division of Union Carbide. The joint venture brought together in synergistic union the strong R&D traditions of both companies. The joint venture now contains the new materials R&D of the CAPS Union Carbide researchers and the scale-up and commercialization skills of UOP research. In addition, the joint venture brings together the commercial experience and worldwide marketing presence of both partners. The result is unprecedented growth for UOP and the development of valuable new technologies, products, and services for its customers. Table 1 summarizes some of the historical highlights of UOP as a process technology company.

Table 1 UOP's History 1914

National Hydrocarbon Company formed to hold Jesse Dubbs patents for a process to recover heavy oil from water

1915

Name changed to Universal Oil Products Company -patents for Dubbs cracking process issued

1921

Dubbs continuous cracking process commercialized

1930

Ipatieff joins UOP beginning a wave of new process developments: alkylation, catalytic polymerization, C4 isomerization

1941

FCC technology developed

1949

Platforming™ introduced, many aromatics processes followed

Late 1950s 1961 Early 1970s

Hydrocracking introduced First Sorbex™ unit licensed CCR Platforming™ introduced

1988

UOP merged with the EP&P and CAPS groups of Union Carbide

1995

UOP acquires the Unocal hydroprocessing business

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In the last 20 years, UOP has developed and commercialized a variety of new and innovative processes for the refining and petrochemical industry including the Penex™, Molex™, BenSat™, Oleflex™, Ethermax™, Merox™, Styro-Plus™, Alkylene™, Isal™, Isomar™ and Detal™processes. UOP transfers this technology to its clients through its licensing activity. In the technology transfer process, UOP licenses technology; assists in the planning, design, engineering and commissioning of new installations; provides management services and advises on the efficient performance of processing facilities throughout the world. Behind the successful performance record of UOP is a highly qualified and strong team continuously at work on ideas and projects. The scientific disciplines are strongly represented in UOP's team of personnel. UOP has about 4,000 employees worldwide. With a wide array of highly specialized talents, UOP offers its clients the complete capability necessary in meeting the demands of today, and the challenges of the future. UOP licenses or maintains a position of technical expertise for more than 60 different processes in the petroleum and petrochemical industry. Approximately 175 process units are licensed yearly, and to date UOP has licensed more than 5,500 individual process units and provided technical know-how in designs for more than 1,000 additional non-licensed units. UOP presently holds in excess of 9,000 unexpired patents. UOP's worldwide licensing activities are supported by a network of offices and representatives. UOP is centered in Des Plaines, Illinois, and has a district office in Houston. UOP Limited, a 100% UOP owned subsidiary for operations in Europe, Africa and the Middle East, has its main European office in Guildford (near London) and district offices in New Delhi, Jakarta, Jeddah, Beijing and Moscow. UOP Asia Pacific, located in Tokyo, is an affiliate company of UOP for the licensing of UOP processes in Japan and certain other areas in the Far East and Southeast Asia.

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UOP has catalyst manufacturing facilities in the United States and in Europe. UOP Asia Pacific operates a catalyst plant in Japan. The international scope of UOP activities is evidenced by the fact that process units have been designed for installation in more than 80 countries around the world. UOP activities related to these installations have ranged from preparation of engineering designs for single process units to extensive planning studies involving market analyses, feasibility and optimization studies, designs for entire grassroots refineries (both process units and offsites), and complete plant commissioning services. The services provided by UOP for these units includes plant design, inspection, commissioning, performance testing, and training of refinery operating personnel. Since 1955, UOP has provided, or is providing, engineering designs for more than 125 grassroots refineries and petrochemical complexes. UOP also provided design specifications for all offsite equipment for many of these installations.

Historical Origins of FCC Technology The advent of the petroleum refining industry can be traced to the rapidly increasing demand for kerosene to fuel kerosene lamps for lighting in the latter half of the 1800s. With the invention of electric lighting and the automobile in the early 1900s, the high value product of petroleum refining shifted from kerosene to gasoline. The increasing demand for gasoline soon outstripped the availability of straight-run gasoline from crude oil distillation. This shortage of gasoline provided the impetus for the development of technologies to increase the gasoline yield from a barrel of crude oil. Table 2 shows a summary of the progression of cracking technology which has led to the FCC process as we know it today.

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Table 2 Historical Origins of Fluid Catalytic Cracking 1913 - 1936

Thermal cracking  Burton thermal cracking process (1913)  Dubbs thermal cracking process (1915)  Current use – visbreaking, coking

1936 - 1941

Fixed-bed catalytic cracking  Houdry Process Company (1931) - Multiple reactors -- cyclic process (1937) - Silica-alumina catalyst (acid-activated clay)

1941 - 1955

Moving-bed catalytic cracking  Thermofor catalytic cracking (TCC) developed by Socony-Vacuum (Mobil)  Houdryform catalytic cracking - Continuous process - Macro-catalyst, moving bed

1942 - Present Fluid catalytic cracking (FCC)  Joint development (1938) - Continuous process - Micro-catalyst, fluidized bed

Thermal Cracking The first thermal conversion process was the Burton process first practiced commercially in 1913 by Standard Oil of Indiana. In the original Burton process, oil was exposed batch-wise to high temperature at elevated pressure to achieve thermal conversion to lighter products. Because of the batch nature of the Burton process, commercial units contained a large number of individual cracking stills in order to achieve acceptable daily throughputs.

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Following the commercialization of the Burton process, the Dubbs thermal cracking process was developed and patented in 1915 (UOP). The Dubbs process was a continuous process for the thermal conversion of oil to lighter products at elevated temperature and pressure. The Dubbs process was widely used in refineries through the 1920s and into the 1930s. Thermal cracking processes continue to be used in refining today. Examples of currently used thermal processes are visbreaking and various forms of coking.

Fixed-Bed Catalytic Cracking In the mid 1920s, a French mechanical engineer and racecar enthusiast named Eugene Houdry became interested in gasoline quality. After the trial and error screening of hundreds of catalyst formulations, Houdry found that acid-activated clay (silica and alumina) was an effective catalyst for cracking heavy oil to lighter products, particularly high octane gasoline. In 1931, Houdry, in partnership with Socony-Vacuum (now Mobil), founded the Houdry Process Company to develop Houdry's fixed-bed catalytic cracking process. The Houdry catalytic cracking was a cyclic process which typically used four timephased reactors, each of which cycled through a sequence of steps outlined below: 1. 2. 3. 4.

Hot heavy oil is cracked by contact with a fixed bed of catalyst. The reactor is purged to remove hydrocarbon. Coke deposited on the catalyst is burned off with air. The combustion gases are purged from the reactor and the reactor is ready to begin the next cracking cycle.

A number of technical innovations were required to make the Houdry cracking process successful. Among these were the development of automatic valves and the use of control algorithms to control the reaction-regeneration cycles. Many of the innovations associated with the commercialization of the Houdry cracking

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process were considered revolutionary in the field of process engineering at the time they were first introduced. The Houdry catalytic cracking process was first commercialized at the Sun-Marcus Hook refinery in 1937. The Houdry process was technically attractive to refiners and by 1940, 14 commercial Houdry units were in operation. Interest in the Houdry process declined after 1941 because of further advances in catalytic cracking technology.

Moving-Bed Catalytic Cracking The next advance in catalytic cracking was the development of a continuous moving-bed cracking process. The Thermofor Catalytic Cracking (TCC) and Houdryform Catalytic Cracking (HCC) processes were developed in parallel in the 1940s and early 1950s. Both processes used a similar concept and had approximately equal success. In the TCC process, the catalyst pellets continuously move through the reactor to the regeneration vessel and are then returned to the reactor. The key to the TCC process was the Thermofor kiln used to regenerate the spent catalyst (the kiln had been originally developed to burn coke off of Fuller’s earth used to filter lube oils). In the TCC process, regenerated catalyst flows by gravity from a surge vessel elevated above the reactor, into the reactor vessel where the catalyst contacts hot oil and the cracking reactions take place. The air environment of the catalyst surge vessel is buffered from the hydrocarbon environment of the reactor by steam injected into the catalyst transfer line. Both the hydrocarbon vapors and catalyst flow down through the reactor to a lower section where the cracked products exit the reactor through separation pipes. The spent catalyst continues to flow by gravity down through a steam stripping zone into the regeneration kiln where coke is burned off the spent catalyst with air. The steam stripping zone also serves to provide a barrier between air in the regenerator and hydrocarbon in the reactor. In early TCC units, the hot regenerated catalyst pellets were mechanically conveyed

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back up to the catalyst surge vessel by bucket elevators. Later units employed pneumatic air lift systems to transfer the regenerated catalyst back up to the surge vessel. Socony-Vacuum was the principle developer of the TCC process and the first semi commercial unit started up at the Paulsboro refinery in 1941. The TCC units were licensed and operated by Socony-Vacuum and others from 1941 to about 1955 when the TCC gave way to the more versatile FCC process developed in the during the late 1930s and early 1940s. A few TCC units still continue to operate today.

The FCC Process Early development of the FCC process took place late in the 1930s. A number of motivations were behind the development of the FCC process. Among these were the high fees required to license the Houdry cracking process, the diffusion and heat transfer limitations associated with both the Houdry fixed-bed process and the TCC process (both used large size catalyst pellets), and the increasing demand for high octane aviation gasoline brought on by World War II. Initial FCC process development efforts were led by Standard Oil of New Jersey (now Exxon) in association with two researchers from the Massachusetts Institute of Technology, Warren Lewis and Edwin Gilliland (consultants to Standard-NJ). Lewis and Gilliland had found that under the proper aeration conditions, finely divided solid particles (powders) could flow through pipes and in many respects act similarly to liquids. This was the advent of fluidization. The use of finely divided cracking catalyst offered a means of overcoming the diffusion and heat transfer limitations encountered with the large size catalyst pellets used in the earlier catalytic cracking processes. In 1938, Standard-NJ and some of the other major oil companies, as well as M. W. Kellogg Co. and Universal Oil Products (UOP), formed Catalytic Research Associates (CRA) to jointly develop a fluidized catalytic cracking technology. The first commercial-scale (13,000 BPD) FCC unit, designated the Model I, started up at

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Standard-NJ's Baton Rogue refinery in May 1942. Two other Model I FCC’s were designed but were not built as the improved Model II FCC design came very quickly. When Standard-NJ announced the construction and imminent startup of the first FCC Model I, they also announced that Universal Oil Products (UOP) and M. W. Kellogg would be designing and licensing the new FCC technology. In the threeyear period between 1942 and 1945, 34 new FCC units came on stream in the refineries of 20 different oil companies. The installed capacity of these new FCC units was over 500,000 BPD. Thirteen of these units were licensed from UOP. Following the commercialization of the Model I and Model II FCC units within the CRA partnership, the FCC unit design and development diverged with the partner companies largely going their separate ways with regard to future FCC technology development and commercialization.

UOP and Fluid Catalytic Cracking During the 1940s, military requirements resulted in widespread commercialization when UOP designed about 40% of the 34 units that were built and operated. Following this period, UOP was in the forefront with commercialization of the "stacked" FCC unit design which featured a low-pressure reactor stacked directly above a higher pressure regenerator. The stacked design not only met the economic needs of smaller refiners, it was a major step toward shifting the cracking reaction from the dense phase of the catalyst bed to the dilute phase of the riser. In the mid-1950s, UOP introduced the "straight-riser" or side-by-side design. In this unit, the regenerator was located near ground level with the reactor to the side in an elevated position. Regenerated catalyst, fresh feed and recycle were directed to the reactor by means of a long, straight riser located directly below the reactor. Compared to earlier designs, product yields and selectivity were substantially improved. A major breakthrough in catalyst technology occurred in the mid-1960s with the development of the zeolitic catalysts. These catalysts demonstrated vastly superior activity, gasoline selectivity and stability characteristics compared to the amorphous

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silica-alumina catalysts then in use. The availability of the zeolitic catalysts served as the basis for most of the process innovations that have developed in recent years. The continuing sequence of advances in both catalyst activity and process design culminated in the most significant concept to date in the field of the FCC process – the achievement of transport-phase cracking entirely in the riser, or all-riser cracking. The key to all-riser cracking is the design of a system that initiates a plug-flow reaction and then stops the cracking reaction at the optimum yield of desired products. UOP commercialized a new design based on this concept in 1971. This design was also applied to existing unit revamps. Commercial results confirmed the expected advantages of the system compared to the older designs. The quick quench design provided the desired high selectivity to gasoline, reduced coke yield, and a reduction of secondary cracking of desired products to lighter, less valuable material. The next major improvement in the FCC technology was the development of catalysts and regenerator systems for the complete internal combustion of carbon monoxide (CO) to carbon dioxide (CO2). In 1973, an existing UOP unit was revamped to include a new combustor concept in regeneration technology to achieve direct conversion of CO within the unit. This was followed by the start-up in 1974 of a new FCC unit specifically designed to incorporate the combustor regenerator technology. This development in regenerator design and operating technique resulted in reduced coke yields, lower CO emissions which satisfy environmental standards and higher circulating catalyst activity that improved product distribution and quality. Table 3 summarizes some of the major achievements in UOP's FCC process technology development and commercialization.

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Table 3 Milestones in FCC Technology 1942

UOP begins licensing the FCC Process

1945

13 Units licensed by UOP

1947

UOP commercializes stacked unit design  Economical for small refiners  50 Stacked designs over 10-year period

1950s

UOP commercializes side-by-side design  Straight riser  Better suited for larger units  Riser extension and termination (more reaction in riser)

1973

First complete combustion regenerator

1983

First two-stage regenerator with external dense-phase cooling for highly contaminated resid feed commissioned

1983

First elevated distributors commissioned

1991 - 1995

Newest generations of highly contained riser termination devices commercialized (VDS™ and VSS™)

1994

First Optimix™ feed distributor commissioned

1994

First MSCC™ unit commissioned

2006

First AF™ Packing commissioned

Recent Developments Advances in riser termination devices occurred at a rapid rate in the 1980s to the mid 1990s. Early riser termination devices such as the open Tee resulted in very long residence times for the hydrocarbon products in the reactor vessel. This extended residence time resulted in nonselective thermal cracking and secondary catalytic cracking reactions. Recent improvements have resulted in better containment of the hydrocarbon vapor to the riser and therefore lower post riser residence time. This reduced delta coke and dry gas and improved gasoline selectivity. Early versions of these high containment riser terminations included the vented riser and

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SCSS (suspended catalyst solids separation) devices. In 1991, the first VDS™ (vortex disengager stripper) was commissioned. This technology further minimized the post-riser residence time resulting in further improvements in product yields. In 1995, the first VSS™ (vortex separation system) was commissioned. Improvements in feed distribution systems also occurred rapidly in the late 1980s and 1990s. Elevated, radially oriented feed distributors minimize nonselective thermal cracking reactions by providing more uniform feed/catalyst contacting with less back mixing than the earlier wye feed distributors. Acceleration zone technology which pre-accelerates the catalyst into a uniform, moderate density flow pattern for optimum oil penetration and uniform catalyst/oil contacting further improved the performance of the elevated feed distributors. The first UOP elevated feed distributors were commissioned in 1983. Developments in spray nozzle technology resulted in the Optimix™ feed distributor which has a smaller, more uniform oil droplet size and a spray pattern that distributes the oil uniformly over the entire riser area for superior catalyst/oil contacting and performance. The first Optimix™ feed distributor was commissioned in 1994. Since then, the number of refiners using Optimix™ feed distributors has grown to over 80. Resid processing in FCC units began in the mid-1970s. During this same period, reactor temperatures were being increased to maximize gasoline octane. The need for higher conversion, combined with the desire to process residue feeds significantly increased coke yields and ultimately limited the FCC regenerator capacity. The RCC®, or Reduced Crude Conversion, process was developed jointly by UOP and Ashland Oil in the late 1970s to address residue processing. It is an extension of UOP's FCC design experience that incorporates many innovations and modifications from the UOP-Ashland Oil development program. In addition to cold-flow modeling work, a large-scale pilot plant was constructed at Ashland's Catlettsburg, Kentucky refinery. Testing in this 200 BPSD plant examined processing variables and new mechanical designs on a wide range of residual feedstocks. In 1983, Ashland commissioned a 40,000 BPSD RCC unit at the Catlettsburg refinery.

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Several major innovations from the pilot plant testing and first commercial design at Ashland have become the foundation of UOP's technical offering for catalytic cracking of residue feedstocks, including the following. • Acceleration zone and feed distribution system • Higher containment riser termination devices for quick disengagement • Two-stage catalyst regeneration • Catalyst cooler Since 1983, eight grass-root RCC units licensed by UOP have been commissioned. In addition, resid feedstocks are being processed in more than 30 existing UOP FCC units. In present times, the distinction between a gasoil FCC unit and a resid FCC unit has blurred to the point where most modern FCC units are capable of processing some level of resid. The term RFCC is used by UOP today to designate a new unit utilizing a 2 stage regenerator designed for the specific intent of processing resid feeds. Table 4 shows a brief summary of resid processing and UOP's activity in the area of resid processing.

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Table 4 Resid Processing Milestones 1940s

Resid component added to feed

1950s

Resid processing diminishes

1975

Resid processing regains attractiveness  Market conditions favor increased efficiency in gasoline production  Technology and catalyst advances increase resid processing potential  UOP units begin processing resid/gasoil blends

1976

UOP and Ashland Oil Cooperation  Research and development for reduced crude conversion  Semi-commercial demonstration

1983

First RCC unit commissioned

1984 - 2006

8 New RCC units operating ->30 Units processing resid

Commercial Experience Since commercialization of the FCC process, UOP has licensed more than 210 units, or over 50% of all non-captive installations. More than 140 of these units continue to operate throughout the world. The superior technology and operational reliability built into UOP FCC units are some of the reasons why 58 refineries worldwide have licensed new UOP FCC units since 1980, which is more than all other licensors combined during this period. UOP's commercial activity in the FCC/RCC/MSCC™ processes since 1980 is as follows:     

63 40 330 180 30+

New units licensed New units commissioned Revamps Major revamps Units processing resid with UOP technology

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Revamp activity is of equal importance in demonstrating technical expertise. In the period 1980-1998, UOP performed more than 330 unit revamps that encompassed virtually every major section of the FCC unit. This activity is vital to UOP's continuing advances in both process and design engineering. The depth of both grass-root and revamp experience gives UOP great capability to respond to the changing needs of the industry.

FCC Process Description The FCC process converts heavy crude oil fractions into lighter, more valuable hydrocarbon products at high temperature and moderate pressure in the presence of a finely divided silica/alumina based catalyst. In the course of cracking large hydrocarbon molecules into smaller molecules, a non-volatile carbonaceous material, commonly referred to as coke, is deposited on the catalyst. The coke laid down on the catalyst acts to deactivate the catalytic cracking activity of the catalyst by blocking access to the active catalytic sites. In order to regenerate the catalytic activity of the catalyst, the coke deposited on the catalyst is burned off with air in the regenerator vessel. One of important advantages of the FCC process is the ability of the catalyst to flow easily between the reactor and the regenerator when fluidized with an appropriate vapor phase. In FCC units, the vapor phase on the reactor side is vaporized hydrocarbon and steam, while on the regenerator side the fluidization media is air and combustion gasses. In this way, fluidization permits hot regenerated catalyst to contact fresh feed; the hot catalyst vaporizes the liquid feed and catalytically cracks the vaporized feed to form lighter hydrocarbon products. After the gaseous hydrocarbons are separated from the spent catalyst, the hydrocarbon vapor is cooled and then fractionated into the desired product streams. The separated spent catalyst flows via steam fluidization from the reactor to the regenerator vessel where the coke is burned off the catalyst to restore its activity. In the course of burning the coke a large amount of heat is liberated. Most of this heat of combustion is absorbed by the regenerated catalyst and is carried back to reactor by the fluidized regenerated catalyst to supply the heat required to drive the reaction side of the

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process. The ability to continuously circulate fluidized catalyst between the reactor and the regenerator allows the FCC unit to operate efficiently as a continuous process. The FCC units are large-scale processes and unit throughputs are typically in the range of about 10,000 to 130,000 barrels per day. This corresponds to catalyst circulation rates of around 7 to 130 tons per minute. The largest commercial FCC unit in operation was designed at 130,000 BPSD, pushed to ~184,000 BPSD, and in 2005 was revamped to a nominal 200,000 BPSD with a catalyst circulation rate in excess of 170 metric tons per minute. These large circulation rates of hot, abrasive catalyst constitute a very significant challenge to the mechanical integrity of the reactor, the regenerator and their associated internal equipment. Thus, mechanical design considerations are critical to the ultimate success of an FCC unit as a prominent refinery process unit. The main features of an FCC unit are:     



Catalytic process Mechanical process Cracks heavy molecules to lighter ones Pressure: 15-45 psig (1-3 kg/cm2g) Temperature: Reactor: 915-1050F (490-565C) Regenerator: 1200-1450F (650-790C) Reaction and regeneration sections intimately linked by heat balance and catalyst circulation

FCC Process Feedstocks FCC units process heavy oil from a variety of variety of refinery flow schemes. Generally, the feed comes from either the refinery crude unit or vacuum unit and constitutes the fraction of the crude boiling in the range of 650 to 1000+°F (350 to 550+°C). There may be additional feed preparation units upstream of the FCC unit such as a hydrotreater or deasphalter. Figure 1 shows a schematic diagram of the possible refinery flows providing feed to an FCC unit.

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In addition, the FCC units commonly process heavy fractions from other conversion units as part of the combined FCC feed blend. Examples of these types of streams are coker gasoil and hydrocracker fractionator bottoms. The types of heavy hydrocarbon streams that are commonly charged to an FCC unit are:        

Atmospheric gasoil Vacuum gasoil Atmospheric resid Coker gasoil Demetallized oil Hydroprocessed gasoil Hydroprocessed resid Lube oil extracts

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FCC Products The products obtained from the FCC unit are light hydrocarbon gases (C2-) which are normally used within the refinery as fuel gas, light olefins and paraffins (C3’s and C4’s) also referred to as LPG, gasoline, LCO and clarified oil commonly referred to as main column bottoms. In addition, flue gas is generated from the burning of coke in the regenerator. Heat is recovered from the flue gas and is used to make steam and in some cases power is also recovered from the flue gas in the form of electricity via a power recovery expander coupled to a motor/generator. Products produced from an FCC unit are:       

Light gas Light olefins LPG Light paraffins Gasoline Light cycle oil Main Column Bottoms Coke (burned in unit as fuel)



Most of the FCC product streams undergo further processing before leaving the refinery as marketable products. Figure 2 shows typical routes for the FCC product steams going to further processing and ultimately to blending into the refinery product pools.

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Figure 2 Typical Use of FCC Products Flue Gas Reactor & Regen.

C3/C4 Paraffins Fuel Gas

Alkylation

LPG Merox Gasoline Gasoline Merox LCO CLO

Alkylate Gasoline Pool

C3/C4 Splitter Main Column & Gas Con

LPG Pool

MTBE Distillate Hydrotreater

Diesel Pool Heavy Fuel Oil Pool

The light liquid products from the FCC process are LPG and gasoline. The LPG from an FCC unit is highly olefinic and has become an increasingly valuable stream for further processing in the present movement toward reformulated gasoline and as petrochemical unit feedstocks. The FCC olefins are an important feedstock for the production of MTBE and alkylate as gasoline blending components and for the production of polypropylene. The FCC gasoline generally has good octane properties (90-95 RON and 80-83 MON) and may make up 30 vol-% or more of the refinery gasoline pool. Some typical characteristics of light FCC products from highconversion operations (VGO Feed, 1.0 wt% sulfur) are:

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

LPG: 500 - 1500 wppm total sulfur 30 - 40 vol-% C3 olefins 34 - 45 vol-% total C4 olefins



Gasoline: C5 - 380F 90% point (193C 90% point) 92 - 94 RONC 0.1 - 0.2 wt-% sulfur 30 - 40 vol-% olefins 25 - 35 vol-% aromatics 0.5 - 1.0 vol-% benzene

The heavy liquid products from an FCC unit are normally LCO and clarified oil. The LCO product is normally used as a blending component in the diesel pool and/or in the heavy fuel oil pool. It is becoming increasingly common for LCO destined for diesel blending to be hydrotreated first for sulfur reduction. Clarified oil is usually blended off to the heavy fuel oil pool. In some cases, the FCC unit clarified oil is used in coker feed, for asphalt production or sold as feed for carbon black production. Some characteristics of heavy FCC products from high conversion operations (VGO Feed, 1.0 wt% sulfur) are: 

Light cycle oil: 600F 90% point (316C 90% point) 20 - 26 cetane index 1 – 1.5 wt-% sulfur 75 - 80 vol-% aromatics 3 - 3.5 cSt @ 122F (50C)



Clarified slurry oil: 2 - 3 wt-% sulfur 9 - 13 cSt @ 210F (100C)



Source: Middle Eastern light gasoil

157048 Introduction Page 23

157048 Introduction Page 24

Abbreviations and Definitions ABD

average bulk density

ACFM

actual cubic feet per minute

activity

conversion of oil by test catalyst compared to standard reference feed often referred to as MAT activity

adjusted

conversion or yields reported as corrected to standard product cutpoints

afterburning

burning of CO above the dense bed in the dilute phase or flue gas, characterized by temperature increase

AGO

atmospheric gasoil

Al

aluminum

Al2O3

alumina

APS

average particle size

AR

atmospheric column resid

ash

non-combustible particles remaining after burning of a main column bottoms sample

as produced

conversion or yields reported as percent of fresh feed at the actual product rates not adjusted to standard product cut points

ASTM

American Society for Testing and Materials

ßo

Coefficient of thermal expansion at 60°F, (1/°F)

behind in burning

insufficient coke combustion in regenerator, characterized by increased coke production in reactor and dark grey regenerated catalyst (high carbon on regenerated catalyst)

BPD

barrels per day

157048 Introduction Page 25

BS & W

bottoms sediment and water, normally reported in vol-%

C1, C2

methane, ethane, ...

C3=

olefin (propylene)

caustic

sodium hydroxide

CCR

catalyst circulation rate

CFR

combined feed ratio (volume of fresh feed plus recycle, divided by volume of fresh feed)

CN-

cyanide ion

CO2/CO

mole ratio of carbon dioxide to carbon monoxide, indicates degree of partial combustion

cold regenerator

operation in conventional controlled regenerator afterburning mode of regeneration

conversion

measure of the rate of gasoil disappearance (or conversion) from feed to products defined as

COS

carbonyl sulfide

CRC

carbon on regenerated catalyst

CSO

clarified slurry oil

Cu

copper

DA, DS, DG

reactor or regenerator purges using air, steam or gas, respectively

P, DP

pressure drop or pressure difference between two points

dry gas

gas from sponge absorber (usually refers to C2-)

EP

end point of distillation

F-1

research octane number (RON)

F-2

motor octane number (MON)

157048 Introduction Page 26

Fe

iron

Fines

catalyst particles less than 20 microns diameter

Fm

feed metals factor

Gasoline efficiency

ratio of liquid vol-% gasoline to vol-% conversion, indicates selectivity to produce gasoline

GC, GLC

gas chromatography, gas/liquid chromatography

Gb

Fluid gravity at base temperature (60°F)

Gf

fluid gravity at flowing temperature

gpm

gallons per minute

H2/C1

ratio of moles hydrogen to moles of methane

H2S

hydrogen sulfide

HC

hydrocarbon

HCN

heavy cat naphtha product drawn from the side of the main column

HCO

heavy cycle oil

HPS

high pressure separator

H

enthalpy (heat) difference

IBP

initial boiling point of distillation

K (UOP K)

measure of paraffinicity or aromaticity of hydrocarbon

lb/Bbl (#/Bbl)

pounds per barrel

LCO

light cycle oil

LV-%

liquid volume percent

M

prefix for thousand

MC

main column

157048 Introduction Page 27

MCB

main column bottoms product

MON

motor octane number

MW

molecular weight

N (or N2)

nitrogen

Na

sodium

NH3

ammonia

Ni

nickel

NOx

nitrogen oxides

O (or O2)

oxygen

ppm

parts per million

Pf

Pressure at flowing conditions (absolute)

recycle

normally refers to heavy oil from main column which has already passed through the reactor that is returned with the fresh feed to the reactor, this could also refer to light material such as LCO or gasoline; a stream which returns to its source.

RE (or Re2O3)

rare earth (or rare earth oxide)

Rg

regenerator

RON

research octane number

RSH

mercaptan sulfur

RVP

Reid vapor pressure

Rx

reactor

SA

surface area

SCF/Bbl (SCFB)

standard cubic feet per barrel of fresh feed

157048 Introduction Page 28

SCFD

standard cubic feet per day

selectivity

preferential towards specified goal or species

severity

combines different factors to give an overall qualitative measurement of extent or difficulty in cracking and regeneration

Si

silicon

Si2O3

silica

sintering

closure of catalyst pores

SOX

sulfur oxides

spillback

gas recycle, may also refer to liquid recycle

SS

stainless steel, also second stage

Tf

temperature at flowing conditions (absolute)

V

vanadium

VGO

vacuum gasoil

vol-%

volume percent

wt-%

weight percent

157048 Introduction Page 29

UOP P&I Diagram Abbreviations AR

Analysis Recorder

ARC

Analysis Recording Controller

DR

Specific Gravity Recorder

FA

Flow Alarm

FE

Orifice Flange Assembly

FFRC

Flow (ratio) Recording Controller

Fl

Flow Indicator

FIC

Flow Indicator Controller

FIF

Flow Indicator Flow Type

FQI

Flow Meter Displacement Type

FR

Flow Recorder

FRA

Flow Recording Alarm

FRC

Flow Recording Controller

FRCF

Flow Recording Controller Float Type

FRCQI

Flow Recording Controller Integrator

FRQI

Flow Recorder Integrator

FRQIA

Flow Recorder Integrator Alarm

HC

Hand Control

II

Current Indicator

LA

Level Alarm

LC

Level Controller

LG-B

Gage Glass Boiler Type—Visible Length Shown

157048 Introduction Page 30

LG-R

Gage Glass Reflex Type—Visible Length Shown

LG-RLT

Gage Glass Reflex Type Visible Length Shown—Low Temperature

LG-T

Gage Glass Through View Type Visible Length Shown

LG-TK

Gage Glass Through View Type Visible Length Shown—KEL-F

LG-TLT

Gage Glass Through View Type Visible Length Shown—Low Temperature

Ll

Level Indicator

LIA

Level Indicating Alarm

LIC

Level Indicating Controller

LR

Level Recorder

LRA

Level Recording Alarm

LRC

Level Recording Controller

PA

Pressure Alarm

PC

Pressure Controller

PDC

Pressure Differential Controller

PDI

Pressure Differential Indicator

PDIC

Pressure Differential Indicating Controller

PDR

Pressure Differential Recorder

PDRA

Pressure Differential Recording Alarm

PDRC

Pressure Differential Recording Controller

PDRCA

Pressure Differential Recording Controller Alarm

157048 Introduction Page 31

PI

Pressure Indicator

PIA

Pressure Indicating Alarm

PIC

Pressure Indicating Controller

PR

Pressure Recorder

PRA

Pressure Recording Alarm

PRC

Pressure Recording Controller

SI

Speed Indicator

SR

Speed Recorder

TA

Temperature Alarm

TC

Temperature Controller

TDR

Temperature Differential Recorder

TDRA

Temperature Differential Recording Alarm

TDRC

Temperature Differential Recording Controller

TI

Temperature Indicator

TIC

Temperature Indicating Controller

TIX

Temperature Indicator Skin

TR

Temperature Recorder

TRA

Temperature Recording Alarm

TRC

Temperature Recording Controller

TRX

Temperature Recorder Skin

TW

Thermowell

Zl

Valve Position Indicator

157048 Introduction Page 32

When Instruments Are Designated with an Alarm H

Indicates High

HH

Indicates High-High, typically in association with an Emergency Shutdown (ESD) system trip point

L

Indicates Low

LL

Indicates Low-Low, typically in association with an Emergency Shutdown (ESD) system trip point

157048 Process Flow Page 1

PROCESS FLOW INTRODUCTION The modern Fluid Catalytic Cracking unit is a large and complex process for cracking heavy gas oil to lighter hydrocarbons. FCC has largely replaced the old thermal crackers because it is a more efficient process, i.e. more production of valuable products at a lower overall cost by using catalyst and heat instead of simply heat. In its simplest form, the process consists of a reactor, a catalyst regenerator, and product separation. This is shown in Figure 1. Catalyst circulation is continuous, at very large mass flow rates. For this reason, the reactor and regenerator are usually discussed as one section. The product separation is usually divided into its low and high pressure components, i.e. the main column section, and the gas concentration and recovery section.

Figure 1: Fluid Catalytic Cracking Process Flue Gas

Regenerator

Air

Catalyst Transfer Lines

Products Reactor

Raw Oil

Product Separation

157048 Process Flow Page 2

Reactor-Regenerator This is the heart of the process, where the heavy feed is cracked. The reaction products range from oil which is heavier than the charge to a light fuel gas. The catalyst is continuously regenerated by burning off the coke deposited during the cracking reaction. This provides a large measure of the heat required for the process.

Main Column The main column cools the reactor vapors and begins the separation process. A heavy naphtha fraction and light and heavy fuel oils (LCO and CLO) come off the tower as products; gasoline and lighter materials leave the top of the tower together and are cooled and separated further into product streams in the gas concentration section.

Gas Concentration and Recovery This section separates the main column overhead into gasoline, liquefied petroleum gas, and fuel gas streams. The composition of each stream is controlled for maximum product value. Figure 2 shows a slightly more detailed schematic of an FCC unit.

Raw Oil

Catalyst Section

MCB HCO

Steam

Light Naphtha

Fuel Gas

Flue Gas

FCC-PF002

LPG

Gas Con Section

Flue Gas Cooler BFW

Heavy Naphtha

LCO

Main Column Section

Power Recovery Section

Figure 2 FCC Block Flow Diagram

157048 Process Flow Page 3

157048 Process Flow Page 4

PROCESS FLOW DESCRIPTION Reactor-Regenerator The FCC process was developed in the early 1940's. A number of companies participated in the early stages of the work, so most of the early units were virtually identical. The first design, the Model I, was installed at only three refineries and quickly replaced by a more successful Model ll. Thirty-one of these were built, thirteen designed by UOP. Figures 3 and 4 show the configurations of the Model I and the Model II FCC’s, respectively. The Model II units had double slide valves and long standpipes, which were a prime source of operating problems due to loss of catalyst fluidization in the standpipes. The raw oil charge passed through a dense bed of fluidized catalyst in the reactor vessel; however, evidence indicated that a large part of the desired cracking was occurring in the transfer line where the hydrocarbon first contacted the catalyst. The early units used low activity catalysts by today's standards, starting with natural clay and later progressing to amorphous synthetic silica/alumina catalysts. Large amounts of heavy oil were recycled back to the reactor in order to obtain the desired conversion levels of 40-60 vol-%. UOP introduced a major departure from the Model ll design in 1947. The regenerator riser was eliminated and the regeneration air was injected directly into the regenerator dense bed. Single slide valves and the more compact design cut construction costs. An important feature of the design was a long reactor riser, which was a major advantage as FCC technology advanced toward entirely riser cracking. The UOP Stacked FCC design proved to be quite popular and UOP designed about 50 Stacked FCC units. Figure 5 shows the typical arrangement of the UOP Stacked FCC unit.

157048 Process Flow Page 5

Figure 3: Model I FCC Cottrell Precipitator

Flue Gas

Catalyst Fines

Products to Main Column

Cyclones

Hoppers Air or Steam

Steam

Standpipes

Regenerator Reactor

Steam Regenerator Riser Water

Catalyst Recycle Cooler

Raw Oil Charge

Reactor Riser

Fired Heater FCC-PF003

157048 Process Flow Page 6

Figure 4 Down-flow Model II Catalytic Cracking Unit

Flue Gas

Cottrell Precipitator

Waste Heat Boiler

Multicyclones Regenerator Products to Main Column

Reactor

Steam to Stripping Section Raw Oil Charge MCB Recycle

Air

FCC-PF004

157048 Process Flow Page 7

Figure 5 UOP Stacked Fluid Catalytic Cracking Unit To Main Column Cyclone

Flue Gas

Reactor

Orific Chamber Spent Catalyst Stripper

Flue Gas Slide Valve

Stripping Steam

Regenerator Spent Catalyst Slide Valve

Regenerated Catalyst Slide Valve Air Slurry Recycle

Raw Oil Charge

HCO Recycle FCC-PF005

157048 Process Flow Page 8

The next advance in reactor-regenerator design was the Side-by-Side unit, shown in Figure 6. This design was better for larger units, where stacking the reactor on top of the regenerator became more expensive. The Side-by-Side layout has also been used for many of the new small units. The straight riser showed less erosion than the curved riser of the Stacked unit. Some of the Side-by-Side units were designed with a reactor dense bed. This bed was eliminated with the advent of zeolite cracking catalysts, and the riser was extended within the reactor to minimize thermal and catalytic cracking by reducing vapor residence time in the vessel. Initially, cyclones were installed on the riser to separate the oil and catalyst, but this was not particularly successful due to poor cyclone performance. The riser cyclones were replaced by a "Tee" shaped termination at the top of the riser. The riser cyclone could then be moved over to allow room for the addition of another cyclone at the reactor outlet, providing two stages of cyclone separation. Later advances in riser termination devices concentrated on maximizing hydrocarbon containment or minimizing the post-riser residence time in the reactor shell where non-selective, thermal and secondary cracking reactions occur. Side-by-Side units have won good acceptance by the industry and over 75 UOP designed Side-by-Side FCC units have been built. Zeolite cracking catalysts were developed in 1963 and gradually accepted by the industry over the next ten years. These catalysts proved to be much more active than amorphous catalysts and were ideally suited for the short contact time riser cracker. Conversion levels rose as high as 80% without requiring excessive reactor temperature. Another significant improvement in FCC reactor technology was the use of elevated feed distributors. The older wye feed distributors injected the raw oil charge into a highly back-mixed catalyst flow that resulted in non-uniform oil/catalyst mixing and excessive light gas and coke formation. In newer systems, multiple, radially oriented feed distributors elevated in the riser inject raw oil more uniformly to maximize selectivity to desired products.

157048 Process Flow Page 9

New regenerator designs were also developed over the years. The old perforated plate air distributor was changed to a pipe grid for better air distribution. Two stage cyclones replaced single stage cyclones and reduced catalyst losses. The burning of coke in the old regenerators was not complete, i.e., not all the carbon went to CO2, and the flue gas normally contained 6-10 vol-% CO. The unit ran with no excess oxygen. This prevented afterburning in the cyclones and the resultant heat damage to the cyclones. An extra furnace to generate steam, the CO boiler, was added to utilize heat that would otherwise be lost. All of the excess CO in the flue gas could be burned in the CO boiler, but capital costs were high. The obvious solution to this problem was to burn all of the CO to CO2 in the regenerator, where the catalyst can absorb the heat. Although this could be done in a standard “bubbling bed” regenerator, a new, “high efficiency” type regenerator design proved more efficient. In the high efficiency or combustor style regenerator, shown in Figure 7, the air and catalyst is mixed in a fast fluidized environment in the lower part of the regenerator or combustor. The fluidized catalyst is then carried up the combustor riser to the upper regenerator. The fluidization in the combustor provides excellent air/catalyst mixing and heat transfer to maximize coke burning kinetics. The high efficiency regenerators on stream average less than 100 ppm CO, and less than 40 ppm NOx in the flue gas. This design enables refineries to get greater thermal efficiency from the unit while simultaneously meeting more stringent air quality standards.

157048 Process Flow Page 10

Figure 6 UOP Side by Side Fluid Catalytic Cracking Unit Rxn Products to Main Column

Reactor Down-Turned Arm

Flue Gas

Flue Gas Slide Valve

Stripping Steam

Bubbling Bed Regenerator

Spent Catalys Slide Valve

Main Distributo

Air

Regenerated Catalys Slide Valve

Wy Section Raw Oil Feed

157048 Process Flow Page 11

Figure 7 Modern UOP Side by Side Fluid Catalytic Cracking Unit With High Efficiency Regenerator, Elevated Feed Distributors and Vortex Separation System Riser Termination

157048 Process Flow Page 12

The process flow of the reactor and regenerator section of a typical, modern FCC unit with a high efficiency regenerator can be described as follows: Lift steam and/or light hydrocarbon is injected at the base of the riser or Wye to accelerate the catalyst from towards the elevated feed distributors, which are located about 1/3 the way up the riser. The preheated raw oil charge is pumped through the feed distributors and atomized with the addition of steam then injected into the regenerated catalyst stream. The heat from the catalyst and reduced hydrocarbon partial pressure in the riser both act to help vaporizes the oil. The catalyst, oil and steam travel up the riser to a region of lower pressure in the reactor where the cracked hydrocarbon products are separated from the catalyst in the riser termination device and cyclones before going to the main column for initial product separation. During the cracking reaction, a carbonaceous by-product called coke is deposited on the circulating catalyst. This catalyst (referred to as spent catalyst) drops from the reactor disengager and cyclones into the stripping section where a countercurrent flow of steam is used to remove both interstitial and some adsorbed hydrocarbon vapors. The stripped catalyst flows from the reactor stripper through the spent catalyst standpipe to the regenerator, where the coke is continuously burned off. The catalyst flow through the spent catalyst standpipe is controlled to balance the circulating catalyst flow by maintaining a constant catalyst level in the reactor. In the regenerator, the spent catalyst mixes with air and hot regenerated catalyst from the recirculation catalyst standpipe at the base of the combustor. Here the coke deposited during in the reactor is burned off to reactivate the catalyst and provide heat for the net endothermic cracking reactions. The heat of combustion raises the catalyst temperature in the regenerator to a range of 1200°F-1375°F (648°C-746°C). The catalyst and air flow up the combustor riser and separate at a "Tee" shaped head. The flue gas is further "cleaned" of catalyst in the cyclones in the upper regenerator. The recirculation catalyst standpipe returns some of the hot regenerated catalyst to the combustor either on temperature or density control to provide heat for initiation of the carbon burn. The remainder or the regenerated

157048 Process Flow Page 13

catalyst flows down the regenerated catalyst standpipe on reactor temperature control to the riser Wye to complete the cycle. The flue gas exits the regenerator through the flue gas slide valves on pressure control to the regenerator. An orifice chamber located downstream of the slide valves acts to reduce the pressure drop and velocity across valves to minimize mechanical deflection of the body and erosion to the internals. Many units have a power recovery unit in place of the slide valve and orifice chamber to recover electrical energy by letting down the high volume, moderate pressure flue gas across a turbo-expander connected to a motor/generator. Finally the sensible heat energy in the flue gas is recovered through steam generation in either a CO boiler or flue gas cooler depending on the mode of operation in the regenerator. Many units also have an electrostatic precipitator or wet gas scrubber to remove catalyst fines from the flue gas before it is discharged to the atmosphere. The reasons and methods for varying the high efficiency regenerator operation will be discussed in more detail later in the PROCESS VARIABLES section. RFCC Regenerator As a result of the crude oil embargoes and oil price rises of the 1970’s, interest in processing heavier feeds in FCC units grew. However, FCC technology at that time could not handle highly contaminated heavy feeds while maintaining a reasonable degree of conversion. In the mid 1970’s, UOP and Ashland Oil Company embarked on a joint development project to develop catalytic cracking technology capable of processing very heavy, highly contaminated feeds, i.e. feeds with high metals and Conradson carbon contents. The result of this development program was the commercialization of the RCCSM (Reduced Crude Conversion) process at Ashland’s Catlettsburg refinery in 1983. The main feature of the RCC unit is a two stage regenerator equipped with a catalyst cooler to remove heat from the regenerator. The upper or first stage regenerator burns approximately 2/3 of the coke from the catalyst in partial combustion mode to limit the heat of combustion and therefore the temperature of

157048 Process Flow Page 14

the catalyst. A portion of the partially regenerated catalyst entering the lower or second stage regenerator flows through the catalyst cooler(s) where heat is removed from the catalyst to generate steam. The cooled catalyst and the remainder of the hot catalyst from the first stage regenerator mix in the second stage regenerator where the coke burning is completed under conditions of complete CO combustion in the presence of excess O2. Carbon is burned off the catalyst to low levels in the second stage regenerator at moderate temperature to maximize catalyst activity. The combustion gases from the second stage regenerator pass into the first stage regenerator where the pre-heated excess O2 improves coke burn kinetics, and is completely consumed. The combined flue gas exits through two stages of cyclones in the first stage regenerator and out through a single common flue gas line. The overall mode of combustion for the two stage regenerator is partial burn with the additional benefit that all of the catalyst returning to the reactor is fully regenerated due to the full burn environment of the second stage regenerator. Figure 8 shows the arrangement of the regenerator of an RFCC unit. The reactor is the same as the modern Side by Side unit shown in Figure 7. Figure 9 shows the process flow for a catalyst cooler. Although catalyst coolers are not a new idea for FCC service, past attempts to employ catalyst coolers on FCC’s have been largely unsuccessful from both mechanical and process points of view. UOP’s catalyst cooler represents an improved design developed and refined to provide both mechanical reliability and a wide range of heat removal flexibility. Heat removal varies with the rate of fluidization air injected to the cooler and the catalyst slide valve opening. The operation of the catalyst cooler is as follows; catalyst enters the cooler shell where the tube bundle is immersed in hot fluidized catalyst. Fluidization air is injected at the bottom of the cooler shell to control the fluidization and heat transfer. Annular bayonet type water tubes are used in the tube bundle. Water entering the bundle flows up through the inner tube, flows out the top of the inner tube and down through the annular space between the inner and outer tubes where heat transfer occurs and water is vaporized to steam. In flow-through style coolers cooled catalyst exits the cooler shell through a standpipe and slide valve and is returned to

157048 Process Flow Page 15

the regenerator to allow hot catalyst to enter the top of the cooler to maximize the cooler duty. Back-mix type coolers rely only on fluidization and back-mixing to transfer hot catalyst from the regenerator rather than using catalyst flow through a standpipe. Back-mix coolers have a simpler mechanical configuration but can only remove approximately 70% of the heat transfer capable through a flow-through cooler. A large excess of water is circulated through the tubes where heat transfer generates steam to ensure that the tube walls are always wet and cooled. The steam and water mixture returns from the cooler bundle to a steam drum where the steam and water are separated. Water from the drum is circulated back to the cooler and the saturated steam from the steam drum is routed to the refinery steam system.

157048 Process Flow Page 16

Figure 8 UOP RFCC Regenerator Process Flow Flue Gas

Spent Catalyst from Reactor

1st Stage Regenerator

Vent Tubes

Catalyst Cooler

First Stage Air

2nd Stage Regenerator

Water/ Steam Water

Recirculation Catalyst Standpipe Regenerated Catalyst to Reactor Second Stage Air

157048 Process Flow Page 17

Figure 9 UOP FCC Catalyst Cooler Process Flow Water and Steam Saturated Steam to Superheater Fluffing Air

Makeup BFW Blowdown

Cooled Catalyst Slide Vlave

Circulating Water

FCC-PF009

157048 Process Flow Page 18

Main Column The main column is the first step in the separation and recovery of the cracked hydrocarbon vapors from the FCC reactor. The reaction products enter the column at high temperatures, 900-1022°F (480-550°C). The main column is similar to a crude tower, with two important differences: 1) The vapors must be cooled before fractionation can begin, and 2) a large quantity of lighter gas passes overhead with the gasoline. Figure 10 shows the general process flow for an FCC main column. Large quantities of heavy oil are circulated over a series of disc and doughnut trays to cool the vapors and wash down entrained catalyst. The heat removed by the main column bottoms and the heavy cycle oil is used for feed preheat, steam generation, reboiler heat in the rest of the unit, or some combination of the three. The catalyst washed out of the reactor is concentrated in the main column bottoms stream. Most of the bottoms flow is directed through exchangers for heat removal and returned to the disc and doughnut trays. The return line must be free draining to avoid plugging problems with catalyst fines settling in low points. Some of the cooled bottoms material from the steam generators may be returned directly to the bottom of the tower as quench to reduce the temperature of the liquid and minimize coking and fouling in the bottoms system. Figure 11 shows a typical process flow for the main column bottoms pumparound and product circuit. Many older units used a slurry settler to separate and return catalyst fines to the reactor with a slurry stream off the bottom of the settler. The main column bottoms product comes off the top of the settler and is normally called clarified slurry oil. In reactors with two stages of cyclones and in units with modern riser termination devices, the use of slurry settlers has normally been discontinued. Heavy bottoms product comes directly from the main column bottoms circulating stream, as does any slurry recycle to the reactor.

157048 Process Flow Page 19

Figure 10 UOP FCC Main Column To Wet Gas Compressor CW

To Sour Water

To Primary Absorber

1 5 6

Gas Concentration Unit

FI 7

19 21

FI 22

Equalizing Line to/from Feed Drum

Steam

26 27 29

30

Heavy Naphtha Product Gas Concentration Unit

Steam

32

Light Cycle Oil Product

33

Flushing Oil

34 35

Gas Concentration Unit

36

Flushing Oil 37

Torch Oil 38

Reactor Vapors

Steam

BFW

Main Column Bottoms Product CW

Raw Oil from Surge Drum

Raw Oil to Reactor

157048 Process Flow Page 20

Figure 11 UOP FCC Main Column Bottoms Process Flow

6 Minimum Spillback

MCB Product Circuit Minimum Flow Valve

3

1

E FRC

Main Column

Rx Overhead

E

MCB Product Pumps

E

E

Tempered Water

Quench

Steam

Steam Main Column Bottoms Circulation Pumps

MCB Product

MCB Steam Generators Water

Raw Oil Water

Circulating Bottoms/ Raw Oil Exchanger

Raw Oil

Net Bottoms/ Raw Oil Exchanger

157048 Process Flow Page 21

As previously mentioned, most FCC units with modern riser terminations and reactor cyclones do not require the use of a slurry settler and new units currently being designed do not include settlers. If a refiner has strict specifications on the ash content of the main column bottoms product, then more advanced alternate fines removal equipment is usually employed to reduce the catalyst fines to very low levels. The two most common types of catalyst removal equipment used today are the micromesh filter and the electrostatic separator. Cyclonic separation devices have also been used, but are typically limited to smaller capacity installations. A typical micromesh filter system will have 2 or 3 vessels with up to 100 filter elements in each. When multiple filtration vessels are used, each filtration vessel is sized for 100% of the design flow rate. One vessel is typically in filtration mode while another is in backflush mode to remove the filter cake from the elements. When enough catalyst fines have deposited on the filter elements to increase the pressure drop across the filter to a pre set limit, the vessel is taken off line for back flushing. Once the filter vessel is off line and drained the vessel is filled with backflush liquid, either HCO or LCO, and allowed to soak to help dissolve any heavy aromatic compounds on the elements. The top of the vessel is then pressured up with either fuel gas or nitrogen to provide the driving force for a high velocity back flush. The back flush material is collected in a receiver vessel and pumped back to the reactor riser. A typical process flow for the micromesh filtration system is shown in Figure 12. A typical electrostatic slurry oil filtration unit consists of 4-16 skid mounted cylindrical shells (modules) depending on the volume of filtrate to be process. Each module contains a high voltage cylindrical electrode surrounded by conductive glass beads, with a ground rod located in the center of the module assembly. During the separation cycle, the glass beads are ionized in an electrostatic field. As catalyst particles flow between the beads, they are electrostatically collected on the surface of the beads. Each module is sequentially back-flushed while the remaining modules in the system continue the separation. In the backflush cycle, the electrode is de-energized and the beads are fluidized, resulting in a circulating motion up through the center of a 9-inch annular electrode and down the outside. The circulation up the center annulus and down the walls of the module creates a scrubbing action, to mechanically scrub the beads clean. Mechanically scrubbing

157048 Process Flow Page 22

the beads as opposed to solvent soaking as with the micro-mesh filters, raw oil feed, or any compatible oil can be used as the backflush medium to an electrostatic filter. The back flush material is directed back to the reactor riser.

Figure 12 Main Column Bottoms Product Filtration System Backwash Gas (N2 or Fuel Gas)

Clean MCB Product to Storage

Filter #1

Filter #2

Filter #3

N2 or Fuel Gas Vent

Back Flush Liquid (LCO/HCO) MCB Product

Receiver Vessel

Catalyst Backwash to Reactor

157048 Process Flow Page 23

There are typically three side-cuts withdrawn from the main column, heavy cycle oil (HCO), light cycle oil (LCO), and heavy naphtha (HCN). The refiner may withdraw all three, only two or one, depending on product needs and tower design. On relatively rare occasions, the main column is designed with a fourth side-cut to discretely fractionate a heavy LCO cut and a light LCO cut. The side-cut streams that go out as product are usually stripped to meet flash-point specifications. Pumparound loops from these side-draws are used to heat balance the main column by exchanging heat with the gas concentration unit reboilers, the raw oil charge or boiler feed water. The heat removed in the bottom and side pumparounds determines the amount of reflux in each section of the tower and must be properly balanced for proper column operation. Gasoline and light gases pass up through the main column and leave as vapors. After being cooled and condensed, unstabilized gasoline is pumped back to the top of the column as reflux to control the top temperature in the column. Figures 13, 14, 15 and 16 show typical process flows for the HCO pumparound, the LCO pumparound, the Heavy Naphtha pumparound and the Main Column Overhead system, respectively.

157048 Process Flow Page 24

Figure 13 Main Column HCO Pumparound

E

Main Column

LIC

E

Gas Con Unit Debutanizer Reboiler

HCO Internal Reflux (Pumped)

To MCB/Feed Exchanger Outlet (for startup) Filling Line (from feed pump)

E

Heavy Cycle Oil Circulation Pumps

To Feed Surge Drum (normally no flow) To Pump Flushing Oil Supply Header

157048 Process Flow Page 25

Figure 14 Main Column LCO Pumparound and Product

FI FIC LIC Steam

FIC LCO Product

BFW CW Preheater

FIC

LCO to Flushing Oil

Debutanizer Feed Exchanger

FIC Stripper Reboiler Rich Oil from Sponge Absorber Lean Oil to Sponge Absorber FCC-PC403

157048 Process Flow Page 26

Figure 15 Main Column Heavy Naphtha Pumparound and Product

Main Column C3/C4 Splitter Reboiler

E

E

Circulating Naphtha/ Debutanizer Feed Exchanger

Steam E FRC

Heavy Naphtha Stripper

LCO Stripper

Reflux (Gravity Flow)

LIC

Heavy Naphtha Circulation Pumps Heavy Naphtha Product Pumps Heavy Naphtha Product Cooler HeavyNaphtha to and from NHT Unit Signal from HCN Hydrotreater

CW FRC E

Hydrotreated Heavy Naphtha Product

157048 Process Flow Page 27

Figure 16

Main Column Overhead System

FCC/DS-R00-37

157048 Process Flow Page 28

The raw oil feed system is included in the main column section for better process efficiency, i.e. to take advantage of the heat from the main column. Feed enters the unit from storage or directly from upstream processes, such as a vacuum tower or a hydrotreater. The latter scheme is more efficient because the feed will not have to be cooled before storage and then reheated flowing into the FCC. The number and type of exchangers used will depend on cost and process factors that will vary with each refinery. Most newer units do not use fired charge heaters. Fired charge heaters have become unpopular due to the increases in fuel costs, operational safety and impact on overall refinery stack emissions. The feed goes directly to the riser after the raw oil/main column bottoms exchanger. Figures 17 and 18 show typical process flow schemes for the FCC feed preheat system without a fired charge heater and with a fired charge heater, respectively.

Figure 17 Feed Preheat Equalizing Line To/From Main Column

Raw Oil Surge Drum

Raw Oil from Crude Unit/ Hydrotreating

To Reactor LCO Product

MCB Product

Circ. MCB

MCB Recycle HCO Recycle FCC-PF401

157048 Process Flow Page 29

Figure 18 Feed Preheat with Fired Heater

Equalizing Line To/From Main Column

Raw Oil Surge Drum

Fired Heater Raw Oil from Storage/ Upstream Unit

MCB Product

To Reactor

Circ. MCB

Fuel Gas

FCC-PF401

157048 Process Flow Page 30

Gas Concentration and Recovery This section further separates the main column overhead products into stabilized gasoline, LPG and fuel gas. The normal configuration is shown in Figure 18. Unstabilized gasoline from the main column overhead receiver is pumped to the primary absorber, where it is used to adsorb C3’s and C4’s in the gas stream at much higher pressure than the main column. From here the liquid stream goes to the high pressure receiver (separator), then the stripper column, where H2S and C2- are removed. The gasoline off the bottom of the stripper is pressured to the debutanizer for separation of LPG and gasoline and vapor pressure adjustment of the gasoline. The overhead of the debutanizer is olefins rich LPG which is often further processed, for C3 and C4 separation and propylene recovery. The gas from the main column overhead receiver goes first to the wet gas compressor. From here it is pressured to the HPS, the primary absorber, and finally the sponge absorber. Valuable light products such as LPG are removed in the first of two vessels by absorption into the gasoline. The second vessel is the sponge absorber which uses a LCO pumparound from the main column as a final absorption stage before the gas goes out as fuel. Wash water, typically clean condensate, is injected into the inlet to the wet gas compressor interstage condenser. From the interstage receiver it is pumped to the high pressure receiver then to the main column overhead condensers. This water washes out salt forming and corrosive and species such as H2S, NH3, cyanides and phenols. The wash water flow is shown in Figure 19.

D HPR IR PA SA S WGC

HPR

Wash Water to MC OVHD Receiver

IR

Debutanizer High Pressure Receiver Interstage Receiver Primary Absorber Sponge Absorber Stripper Wet Gas Compressor

Legend:

Wash Water

WGC

Gas from MC Overhead Receiver

Gasoline from MC Overhead Receiver

40

P A

9

1

Rich Lean Oil Oil

S A 36

S

1

Typical Gas Concentration Unit Process Flow

Figure 19:

40

D

1

Stabilized FCC Gasoline

HCO

LPG

Fuel Gas

157048 Process Flow Page 31

Main Column

Hydrocarbon Water

Condensate

To Sour Water Stripper WGC 1st Stage

Main Column Receiver

Interstage Drum

Wash Water Flow

Figure 20:

WGC 2nd Stage

High Pressure Receiver

157048 Process Flow Page 32

157048 Process Control Page 1

PROCESS CONTROL Reactor-Regenerator Control Systems Most of the control systems in a Fluid Catalytic Cracking unit are similar to those used elsewhere in the refinery. Good control of the catalyst circulation through the reactor and regenerator is critical for stable operation. The catalyst circulation control scheme is shown in Figure 1. This figure shows a side by side unit with a bubbling bed regenerator but the catalyst circulation control between the reactor and regenerator is the same on all FCC and RFCC units. The circulation of hot regenerated catalyst from the regenerator to the reactor is controlled to maintain a constant reactor temperature with the regenerated catalyst slide valve. The circulation of spent catalyst from the reactor to the regenerator is controlled to maintain a constant catalyst level in the reactor with the spent catalyst slide valve. The controls on both the spent catalyst and regenerated catalyst slide valves also include a low differential pressure override. If the differential pressure across either slide valve drops to a very low or negative value the override will close the slide valve. This minimizes the possibility of reverse flow in the standpipes, either air entering the reactor or hydrocarbon entering the regenerator, which are hazardous situations. Figure 1 shows signals from the reactor temperature controller and level controllers going to low level selectors (LSS). The low signal selectors also receive signals from the differential pressure controllers on the corresponding slide valve. If the differential pressure across the slide valve is greater than the override setpoint, typically 2 psi (0.14 kg/cm2) the LSS will select the process variable (level or temperature) to control the slide valve. If the slide valve pressure drop falls below the override setpoint the LSS will send that output to the slide valve which will start closing. The low differential pressure override controllers should always be in automatic.

157048 Process Control Page 2

Figure 1 Catalyst Circulation Controls Products to Main Column

TIC PDIC

Reactor

Riser Termination Device

HSS

> LIC

Flue Gas

PIC

Regenerator

LSS

LI

< PDIC

XI (Density)

Spent Catalyst Slide Valve LSS

< PDIC

Raw Oil

Air

Regenerated Catalyst Slide Valve Lift Gas/ Steam

FCC-PC001

157048 Process Control Page 3

For steady control of the catalyst circulation between the reactor and regenerator the differential pressure across the slide valves must be constant. To ensure steady slidy valve DP’s the differential pressure between the reactor and regenerator is controlled with the double disc flue gas slide valves on the outlet of the regenerator. In addition to the slide valves an orifice chamber is also used to take approximately 2/3 of the total flue gas system pressure drop to minimize erosion in the flue gas slide valves. A typical flue gas system without power recovery is shown in Figure 2. The reactor pressure is not controlled directly and floats on the main column overhead pressure. The reactor-regenerator differential pressure controller allows the regenerator pressure to change along with the reactor pressure. Pressure control for units with power recovery is discussed later.

Figure 2 Regenerator Pressure Control/ Flue Gas System HSS >

PIC

PDIC

Signal from Reactor Pressure Tap

Flue Gas

Flue Gas Slide Valves

Orifice Chamber Steam

CO Boiler Air Air

Electrostatic Precipitator

Water

FCC-PC002

157048 Process Control Page 4

Reactor Control (Figure 3) Reactor temperature is controlled by the flow of hot regenerated catalyst as described above. The temperature controller may be located in the reactor vapor line or in the upper vapor space of the reactor vessel depending on the type of riser termination device. The reactor pressure is not directly controlled. Reactor pressure floats on the main column overhead pressure. Thus, the reactor pressure is the sum of the main column overhead receiver pressure plus the pressure drop through the main column and MC overhead condensers plus the pressure drop through the reactor cyclones and reactor vapor line. The reactor catalyst level is controlled by the flow of spent catalyst to the regenerator as described above. Modern reactor designs include a wide range level indicator and a more accurate, narrow range level controller. Also, the density in the spent catalyst stripper is measured to allow compensation of the level indication for actual catalyst density. The level is typically controlled off of the wide range LIC because the signal has less noise but a switch is included in newer unit designs to allow use of the more accurate narrow range LIC if desired. Steam is used to atomize the feed in the elevated Optimix feed distributors (Figure 4). The amount of steam used determines the both the pressure drop and extent of atomization as well as the velocity out of the distributor tip and penetration into the catalyst. During normal operation the steam and oil enter the nozzle separately and are mixed internally near the tip. The steam flow is normally set at design rates, typically1-2 wt% of the design charge rate. During operation at turndown additional steam may be used to maintain the optimal pressure drop across the nozzles (typically 50-75 psig) to ensure adequate atomization is maintained. An alternative that allows maintaining near design pressure drop and atomization at turndown with minimal additional steam is to mix a small amount of steam directly with the oil before it enters the nozzle. This results in a large increase in pressure drop with minimal increase in velocity out of the tip. The oil and steam flows to each nozzle have restriction orifices with pressure drop indicators and globe valves to ensure that the flows to each nozzle are equal.

157048 Process Control Page 5

Figure 3 Reactor Control and Instrumentation PDI TIC LIC

LIC

X

SWITCH

LI (Density Compensated)

XI (Density) XI (Density)

Prestripping Steam

PDIC

< LSS

Stripping Steam

MPS

Fluffing Steam

Spent Catalyst LSS PDIC

Regenerated Catalyst

<

Atomizing Steam For Feed/Steam controls see Figure 4 Raw Oil

FIC Lift Gas FIC

Lift Steam

MPS FIC Start-up/Emergency Steam

FCC-PC003

157048 Process Control Page 6

The feed bypass system is also shown in Figure 4. When a situation requiring a quick shutdown is encountered, a control board mounted switch is activated which trips a solenoid valve controlling the pneumatic signals to the feed bypass valves, causing these valves to move to their failure positions, i.e. the valve in the line to the riser closes and the valve in the bypass line to the main column opens. Normally, the next course of action is to open the emergency steam to the riser to either maintain catalyst circulation if the regenerated catalyst slide valve remains open or to clear the riser of catalyst if the regenerated catalyst slide valve is closed. On units with elevated feed distributors, another important operating variable affecting the yield pattern is the lift zone velocity. The lift zone velocity is a function of the lift steam and/or lift gas flow rates which are controlled on straight flow control. As the lift zone velocity is increased the catalyst density at the point of the feed injection decreases allowing better penetration of the atomized oil droplets into the catalyst. The optimum lift zone velocity is typically in the range of 10 – 15 ft/sec (3 – 4.5 m/sec). Either lift gas, lift steam or a combination of both may be used to achieve the optimum lift zone velocity. Stripping steam flows are also controlled on straight flow control. The optimal total stripping steam rate is typically 1.7 – 2.5 lb/M-lb catalyst circulation but this can vary significantly with depending on the unit design. The stripping steam rate should be changed when any process conditions are changed that result in a significant change in the catalyst circulation rate. The stripping steam rate should also be tested occasionally to ensure that the optimum steam rate is used.

157048 Process Control Page 7

Figure 4 Feed/Atomizing Steam Control

FI

FI FI

To other nozzles

PI

Steam FO

Local PI must be readable from valve PI

Local FI must be readable from valve

FI FI

To other nozzles

Local FI must be readable from valve To other nozzles

NORMALLY NO FLOW HEADER PURGE

HS

Feed Bypass Switch

Raw Oil to Main Column

S Instrument Air

FO Vent

FC

Raw Oil from Preheat

FCC-PC004

157048 Process Control Page 8

Bubbling Bed Regenerator Control In full combustion units without a catalyst cooler the regenerator temperature is not directly controlled and is a function of a number of process variables. In simple terms, the regenerator temperature is a function of delta coke, i.e. the wt% coke on spent catalyst entering the regenerator minus the wt% coke on regenerated catalyst leaving the regenerator. The concept of delta coke is discussed in more detail later in the section of this book covering Process Variables. The regenerator may be operated to burn the coke on catalyst completely to CO2 (complete combustion mode) or may be operated so that some of the coke is burned to CO (partial combustion mode). In units that operate in partial combustion mode the CO2/CO ratio of the flue gas may be controlled to adjust the heat of combustion and therefore adjust the regenerator temperature. The CO2/CO ratio is controlled primarily with the amount of combustion air entering the regenerator. Partial combustion operation is discussed in more detail in the Process Variables Section. In units with a catalyst cooler, the regenerator temperature may be controlled by controlling the amount of steam generated in the cooler. The controls of the catalyst cooler are discussed in more detail later in this chapter. The regenerator catalyst inventory serves as the surge capacity for catalyst in the system and there is no control instrumentation on the regenerator catalyst level. A level indicator is provided to monitor the regenerator catalyst level. The regenerator catalyst level changes with catalyst additions, withdrawals and losses. In most units the level is controlled by intermittent withdrawals of equilibrium catalyst. It is important that the level be maintained above the terminations of the cyclone diplegs and below the level that would cause the catalyst in the cyclone diplegs to back up into the cyclone dustbowl.

157048 Process Control Page 9

The air rate to the regenerator is controlled either to maintain a minimum of excess oxygen in the flue gas (typically 2%) for full combustion operation or to control the CO2/CO ratio and therefore the heat of combustion and regenerator temperatures in partial combustion operation. High Efficiency Regenerator The instrumentation and controls for a FCC unit with a high efficiency, combustor style regenerator is shown in Figure 5. The pressure, regenerated catalyst temperature and combustion air rate for the high efficiency regenerator are the same as the bubbling bed regenerator described above. The difference between the high efficiency regenerator and a conventional "bubbling bed " regenerator is that the regenerator is divided into two sections. The lower section is called the combustor and is where the spent catalyst and air mix and coke combustion occurs. The combustor operates in the fast fluidized regime of fluidization. All the catalyst entering the combustor is transported up the combustor riser into the upper regenerator where the regenerated catalyst disengages from the flue gas. The upper regenerator holds the cyclones, provides volume for the regenerated catalyst to disengage from the flue gas and provides the surge capacity for catalyst in the system. An important feature of the high efficiency regenerator is the recirculating catalyst standpipe and slide valve. The recirculation of hot regenerated catalyst from the upper regenerator to the combustor is important in controlling the coke combustion rate. By controlling the amount of catalyst recirculated, the following parameters are controlled in the combustor: the pre-combustion mixing temperature, the catalyst density, catalyst flux and catalyst residence time. This, in turn, allows the coke combustion rate and catalyst regeneration to be optimized. The recirculating catalyst slide valve is controlled through a low signal selector and a slide valve PDIC, similarly to the other slide valves. On early designs, this slide valve position was set on hand control. In current designs, the recirculation slide valve position is controlled by a temperature or density controller located in the upper section of the combustor. A switch is used to select the input signal to the recirculation slide valve low signal selector.

157048 Process Control Page 10

Figure 5 High Efficiency Regenerator Controls and Instrumentation >

TI's (1 Each Cyclone)

PIC

PDIC Signal from Reactor Pressure Tap

LI

XI (Density)

Recirculating Catalyst Slide Valve

FIC Fluffing Air Regenerated Catalyst to Reactor

PDIC TIC LSS

XIC (Density) Spent Catalyst from Reactor

<

SWITCH

FCC-PC005

157048 Process Control Page 11

In a high efficiency regenerator the dense bed catalyst level in the upper regenerator provides the surge volume for the unit and is not controlled directly except by catalyst additions and withdrawals. A small flow of fluffing air to the upper regenerator on straight flow control is required to ensure proper fluidization and flow into the regenerated and recirculation catalyst standpipes.

RFCC Two Stage Regenerator Control The regenerator control systems for the RFCC unit with a 2 stage regenerator are shown in Figure 6. The reactor control systems are identical to those described above for the conventional reactor-regenerator. The principal difference is that the coke combustion is completed in 2 stages. The upper regenerator, or 1st stage, is a bubbling bed regenerator operating in partial combustion mode without excess oxygen. The catalyst exiting this stage still contains a significant amount of coke which is burned off in the 2nd stage operating in full combustion mode with excess oxygen. This allows the benefits of both partial combustion (lower regenerated catalyst temperature) and full combustion (very low carbon on regenerated catalyst for maximum activity). In the RFCC the catalyst level in the second stage is controlled by the flow of catalyst from the first stage to the second stage through the recirculation catalyst standpipe and slide valve. This slide valve has a differential pressure override which closes the slide valve if the pressure drop across the valve drops to low as described above for the other slide valves. The level in the first stage regenerator is the surge volume for the unit and is typically controlled by periodic withdrawals of equilibrium catalyst. Since RFCC units are designed for heavily contaminated feed stocks one or more catalyst coolers are included in the design and are used to control the temperature of the regenerated catalyst circulated back to the reactor. The total air rate which controls

157048 Process Control Page 12

the CO2/CO ratio in the flue gas and therefore the heat of combustion is also used to adjust the temperature of the regenerated catalyst.

Figure 6 RFCC Regenerator Controls and Instrumentation

LI Spent Catalyst from Reactor XI

Side View

Vent Tubes PDIC

PDIC LIC

First Stage Air LSS

LSS

<

< TIC

Recirculation Catalyst Standpipe

Cooled Catalyst Standpipe

Regenerated Catalyst to Reactor Second Stage Air

Second Stage Air

FCC-PF006

157048 Process Control Page 13

Air Blower Control The control system for the main air blower depends on whether the blower is an axial or centrifugal machine and whether it is turbine or motor driven. The most common configuration built today is a turbine driven axial blower because these are generally more efficient, but the choice is unit specific. Figure 7 shows a conventional control scheme for a turbine driven axial blower. Flow is measured by the venturi in the discharge line and is controlled by varying the speed of the turbine. Alternatively, the air rate may be controlled by varying the stator vane position for a fixed speed, axial blower. A variable vent line, called a snort, is located on the air blower discharge and is used to prevent air blower surge. An anti-surge system control system constantly monitors the flow through the blower and the discharge pressure and compares the operating condition to the surge line on the blower curve. If conditions close to the surge line are detected the anti surge-controller opens the snort valve to increase the flow to atmosphere and reduce the discharge pressure of the blower. Modern anti-surge control systems available from specialized vendors continuously monitor a number of process variables and calculate deviation from surge to allow operation closer to the surge line while providing better protection for the equipment. The process variables monitored by the modern anti-surge controllers are shown in Figure 7. Occasionally, on older partial combustion units, an additional, smaller vent is used to fine tune the air flow to the regenerator. This may be tied into a differential temperature controller (DTC) which controls the temperature difference between the regenerator dense and dilute phases to limit or control afterburn in partial combustion units. On RFCC units control valves are used on the air lines to the first and second stages of the regenerator (Figure 8). Typically the air to the second stage is set at ~25-30% of the total air flow. In some units an additional control loop is included to minimize the discharge pressure and thereby minimize the energy consumed by controlling the axial air blower stator vane position. This control loop minimizes the blower discharge pressure until one of the 2 flow control valves on the discharge is nearly wide open.

157048 Process Control Page 14

Figure 7 Main Air Blower Control Vent to Atmosphere

Silencer Snort Valve

Low Flow Shutdown

FO

FIC (Temp, Pressure Corrected)

Instrument Air

XIC (Anti-Surge)

S

Steam

Vent

ST SIC Suction Filter Housing

T TT PT

Air Cylinder FT

TT PT

TT PT

FT

Air to Regenerator

Special Check Valve

FCC-PC007

Figure 8 RFCC Main Air Blower Control (Anti Surge and Special Check Valve Details Not Shown) Vent to Atmosphere

Silencer

FIC (Press, Temp Corrected)

TT PT

ZIC

ZT To First Stage Regenerator

XIC (Anti-Surge) Steam

Suction Filter Housing

FO

Special Check Valves

FIC (Press, Temp Corrected)

T TT PT

ZT To Second Stage Regenerator

Stator Vanes

FCC-PC008

157048 Process Control Page 15

The check valve in the air blower discharge line isolates the blower from the regenerator if the blower trips. This minimizes the possibility of hot catalyst backing up into the blower and minimizes the volume of air in the piping if surging occurs. The closing action is assisted by a spring loaded air cylinder, which operates when the air flow falls below a certain limit. When the flow drops below the low limit a solenoid valve deenergizers and vents the air from the cylinder allowing the spring to move a cam which bumps the check valve to assist it in closing. An air line from the catalyst hoppers or plant air is used to provide plant air to clear catalyst from the discharge line if the blower is down. It can also be used to supply warm blower air to the catalyst hoppers.

Power Recovery Controls A typical process flow and control scheme for the flue gas system on an FCC with a power recovery unit is shown in Figure 9. On units with a power recovery turbine, butterfly valves are used in the flue gas line instead of slide valves for pressure control. The butterfly valves operate on a single, split range controller which first opens the valve on the line to the expander inlet then opens the butterfly valve on the bypass around the expander if the capacity of the expander is exceeded. The bypass valve may also be controlled to limit the speed of the expander or the pressure in the expander. In the past, flue gas power recovery systems were designed with regenerator pressure held steady to minimize fluctuations to the power recovery expander-motor/generatorair blower train. With this control strategy, a regenerator pressure controller output signal was used to control the power recovery butterfly valve positions. In this control scheme, the regenerator pressure is fixed and the reactor-regenerator differential pressure is allowed to float within reasonable limits. Recently, the control system for flue gas power recovery pressure control has been modified so that the reactor-regenerator differential pressure is controlled with a differential pressure controller as in conventional flue gas systems. The objective of this change in pressure control strategy is to insure that the reactor-regenerator pressure balance and catalyst circulation are maintained at steady conditions. Now the output signals of the reactor-regenerator PDIC and the regenerator PIC are fed to a high signal

157048 Process Control Page 16

selector (HSS). The high signal selector directs the appropriate control signal (normally the PDIC) to set the positions of the expander inlet and bypass butterfly valves via a split range signal. The expander turbine is discussed further in the section covering Equipment. Modern power recovery controls, while controlling in the same manner, are more complicated than discussed here. Power recovery controls and anti-surge controllers are provided by specialized vendors. The additional functions performed by these controllers are beyond the scope of this manual. On units with power recovery, a steam turbine may be used for startup of the air blower. The turbine can also provide auxiliary power if necessary. The motor/generator imports or exports power to maintain a constant speed on the power recovery train. If more energy is being supplied to the expander and the turbine than is required by the blower there will be surplus electricity generated and exported. If the blower needs more power than the expander is providing then electricity will be consumed to hold the train at normal speed. The flue gas leaving the regenerator flows through a series of small cyclonic devices in the third stage separator for catalyst fines removal to minimize erosion in the expander. The underflow catalyst stream from the third stage separator, carried by a small gas flow, bypasses the expander and typically leaves the system with the main flue gas stream downstream of the expander turbine. The underflow from the bottom of the separator is controlled by a restriction orifice called the critical flow nozzle. A flue gas quench system is also included on units with power recovery. If the flue gas temperature exceeds the design temperature of the expander a flow of cooling steam is started to the regenerator plenum through a split range controller. If the steam fails to cool the flue gas sufficiently a flow of cooling water, typically BFW, is started to the plenum. Most power recovery vendors also include an emergency steam quench at the inlet of the expander for additional protection.

157048 Process Control Page 17

Figure 9 Power Recovery Controls Quench Connection See Detail Below

TIC > PIC

HSS

Split Range

Flue Gas

Orifice Chamber

PDIC

Butterfly Valves

Third Stage Separator

Flue Gas Cooler

Critical Flow Orifice

Motor/ Air Blower Generator M

Electrostatic Precipitator

Steam T

Steam Turbine

Expander Air

Regenerator Plenum Chamber Quench Detail HighTemperature Signal Opens This Control Valve First

TIC

Split Range

LPS

Atomizing Steam FC

Purge Steam

FI BFW

Signal From Regenerator Temperature Transmitter

To Other Nozzles FC

RO

Concentric Sleeve Purge

Stuffing Box

Plenum Shell

Retracta ble Tip

FCC-PC009

157048 Process Control Page 18

Catalyst Cooler Controls Catalyst coolers provide FCC operating flexibility, permitting direct control over the regenerated catalyst temperature. The regenerated catalyst temperature is a major variable impacting cracking reactions since it sets both the catalyst to oil ratio and determines the temperature of the catalyst surface at its first contact with the oil feed. Both of these variables are important in determining the overall conversion and yield pattern in the FCC. Figure 10 shows a schematic of a flow-through catalyst cooler installed on a high efficiency regenerator with the associated cooler fluidization air, water circulation, steam drum and control instrumentation. As already discussed, hot catalyst from the upper regenerator flows through the cooler shell, around the water tubes of the inserted tube bundle and out of the cooler through a cooled catalyst standpipe and slide valve into the combustor. The fluidization air lance system delivers fluidization air to the cooler shell to maintain catalyst fluidization and mixing in the shell and to ensure that catalyst flows smoothly through the cooler and out through the standpipe. The combination of mixing and net catalyst flux through the cooler provide the driving force for heat transfer by maintaining contact of hot catalyst with the tube wall. The cooler duty is therefore controlled by controlling the amount of fluidizing air and the flow through the cooled catalyst standpipe. Minimum and maximum fluidizing air rates are typically specified to ensure that the air lances do not plug with catalyst and that high velocities in the cooler do not cause erosion on the tube surface. In back mix type coolers without a standpipe and slide valve the cooler duty is controlled only with the fluidizing air. To protect the tubes from thermal damage and oxidation, a large excess of water is circulated through the tubes to ensure that the tube walls remain wetted. The ratio of water circulated to steam generated is typically in the range of 20:1 to 25:1 lbs water circulation per pound of steam generated. In the most recent designs the water circulation rate is determined by the pump curve and no control valve is provided. In earlier designs water circulation is controlled by a control valve. A low flow signal from the flow indicator or controller activates a spare water circulation pump autostart. A low low flow signal from the water circulation flow indicator or a low low steam drum level

157048 Process Control Page 19

effectively causes the cooler to shutdown by closing the cooled catalyst slide valve and closing the fluidization air control valve. The mixture of steam and water exiting the cooler tube bundle flows to the steam drum where the steam and water are separated. The saturated steam flow from the drum is metered and flows out to the refinery steam system. Normally the saturated steam leaving the steam drum is superheated in some type of downstream steam superheater. The water returned to the steam drum is recirculated back to the cooler tube bundle. A small continuous blowdown flow of water is removed from the drum to control accumulation of impurities in the circulating water. The outputs of the steam product flow meter and the steam drum level indicator are summed and control the flow of boiler feed water (BFW) makeup entering the steam drum.

157048 Process Control Page 20

Figure 10 Catalyst Cooler Controls Steam and Water

Low Low Flow or Low Low Level Fluffing Air Trip

Steam

Low Low Level (2/3 Voting)

LG

LT

I

FI

LI PI

VENT

FIC

S

I

LT

TI

DE

PG

STEAM SEPARATOR

Air

LIC

LT

TI

I

FI

Low Low Flow or Low Low Level Slide Valve Trip

Intermittent Blowdown

Continuous Blowdown

Low/Low Low Flow (2/3 Voting)

FIC

I

FI

I

Low Flow Pump Auto Start FT

Makeup BFW VENT S DE

To Motor Control Circuit

To Motor Control Circuit

Steam

FO

T

M

M

FCC-PC010 I

Interlock System

157048 Process Control Page 21

Emergency Interlock Systems For many years automated interlock systems that removed feed from the FCC were not used because it was assumed that a well trained operator would make a better decision regarding stopping feed to the unit than any logic system looking at only a limited number of inputs. Also, since many of the inputs into the interlock system relied on pressure taps around the reactor and regenerator which were prone to plugging, the threat of spurious trips was too great. Recently, however, many refiners understand the value of a properly designed interlock decision to automatically remove feed from the FCC and place it in a safe condition. Also, in many refineries the turnover of operations personnel has increased so that many FCC operators have limited experience. There have been a number of incidents where operators tried to keep the unit running during upset conditions warranting shutting down the unit only to greatly increase the mechanical damage done to the unit. Instead of a shutdown lasting only a few hours an extended shutdown resulted. UOP now includes an emergency interlock system on all new units and major revamps. The purpose of this system is to move the unit to a safe condition during an abnormal event while permitting a safe, fast restart of the unit once the problem is resolved. This system automatically performs the steps necessary to accomplish these goals that were once left to the operator. The system is also designed to minimize the risk of spurious trips or shutdowns resulting from false indications when no abnormal condition existed. The emergency interlock system monitors the air flow rate, regenerated and spent catalyst slide valve pressure drops and valve positions, feed flow rate, reactor temperature, and reactor stripper level to determine and verify the existence of an abnormal event warranting a shutdown of the FCC. Once the abnormal event is detected and verified by 2 out of 3 voting systems or a by a combination of abnormal process readings the following actions are automatically initiated:

157048 Process Control Page 22

    

The feed is bypassed to the main column Raw oil flow rate is reduced (now going to the main column) The spent and regenerated catalyst slide valves are closed Steam is increased to the riser All related controllers are placed in manual as required

Torch Oil Nozzle Control Figure 11 below shows a typical torch oil arrangement for an FCC regenerator. The torch oil nozzles provide a means of injecting heavy oil, usually raw oil feed or HCO, into the regenerator when extra heat is needed, e.g. during startup. Details of the nozzle will be described in the section covering equipment. Control of the torch oil flow and the torch oil atomization steam is often by hand control on older units. New unit designs usually use a flow controller with a split range output to the torch oil flow and the torch oil atomization steam to ensure that the atomizing steam is commissioned before the oil. The torch oil assembly is provided with a continuous steam purge to the annular space around the torch oil nozzle to keep it clear of catalyst. This purge steam flow is regulated to less than 50 lb/hr (25 kg/hr) steam with a restriction orifice. In addition, a small flow of steam is sent to the nozzle tip through a restriction orifice when torch oil is not in use to help cool the nozzle and keep it from plugging with catalyst.

157048 Process Control Page 23

Figure 11 Torch Oil Control Atomizing Steam

PI FC

Steam

Annular Sleeve Purge Steam

STRAINERS

Split Range (Steam Valve Opens First)

Manifold Purge Steam

FI MUST BE READABLE FROM CONTROL VALVE

PI

RO

FI

FC

FIC

Regenerator Shell and Lining

PI MUST BE READABLE FROM CONTROL VALVE

I

Stuffing Box

S

Instrument Air

FC VENT

DE

Emergency Interlock Unit Shutdown Logic

From Raw Oil Pumps

From HCO Oil Pumps FCC-PC011

157048 Process Control Page 24

Direct Fired Air Heater (DFAH) Figure 12 shows the control system of a modern direct fired air heater, present on all FCC units between the main air blower and the combustion air inlet to the regenerator. The air heater is used for refractory curing and dry-out following repair or renewal of regenerator linings as well as during normal startup to heat the regenerator catalyst inventory. The DFAH outlet temperature is controlled by the fuel gas rate. A high temperature shutdown trips the fuel gas to the heater to prevent damage to the air grid in the regenerator. Modern air heater controls trip the fuel gas on a flame out signal from a flame sensor or on low air flow rate. Separate shutoff valves are provided so that the tight shut off of only the fuel gas control valve is not relied upon for emergency trip conditions.

Figure 12 Direct Fired Air Heater Control To Regenerator

Air Ignitor Start Flame Sensor

High Temperature Shutdown I

PDI Must Be Readable From Damper

I

TIC

Damper

Ignitor

Sight Port

Burner

Pilot/ Ignitor

Sight Port

Pilot Gas PI

I

Sight Port (sighting flame)

PI

Sight Port (sighting opp. wall

PI Pressure Control Valve PI

PI

PIC

S S

Interlock Shuts Down the Air Heater on:

S

VENT

DE

FC

INSTRUMENT AIR

Fuel Gas

  

Low Air Flow High Outlet Temperature Loss of Flame Detection

FCC-PC012

157048 Process Control Page 25

Catalyst Addition Controls Continuous additions of fresh catalyst are required to maintain the activity of the catalyst in the reactor and regenerator and to replace catalyst fines lost from the unit through the cyclones. Regular additions of small batches of catalyst results in more stable operating conditions and yields than larger batches added less frequently. The typical UOP catalyst addition system uses a volume pot and automated valve sequence to add a constant volume of catalyst at timed intervals. This system is shown in Figure 13. Specialized valves, designed to close on the catalyst transfer lines when they are full of catalyst, are required to ensure reliable service. Other systems are available including weight addition systems which include a load cell to add weighed batches of catalyst on a regular interval.

157048 Process Control Page 26

Figure 13 Catalyst Addition Control Fresh Catalsyt Hopper INSTRUMENT AIR

I

I

VENT

S

S

FC

Instrument Air

Plant Air Vent

Plant Air FC

RO

Volume Pot

RO

I

S

Instrument Air

Plant Air Vent

FC

FI FI Plant Air

Sight Flow Indicator

To Regenerator

FCC-PC013

157048 Process Control Page 27

Main Column Control The main column is the first step in the product separation sequence. The superheated reactor vapors need to be cooled so that fractionation can be conducted. In large measure, operation of the main column becomes an exercise in controlled heat removal coupled with sufficient liquid-vapor contacting to effect the desired degree of fractionation into desired product streams, typically main column bottoms (MCB), light cycle oil (LCO), heavy naphtha (HCN), unstabilized gasoline and wet gas. MCB, LCO and HCN are drawn as products directly from the main column, although on many FCC units HCN is not removed from the unit as a product stream. On some units, Heavy Cycle Oil (HCO) is drawn as a product from the main column between the MCB and LCO products. Main column sidecut products are often steam stripped in sidecut strippers for flash point adjustment. The unstabilized gasoline and wet gas are further separated in the gas concentration section. The arrangement and integration of heat exchange from an FCC main column varies from refinery to refinery based on the specific requirements and economics of a given installation. The following discussion describes typical heat exchange circuits used on an FCC unit. Figure 14 shows a simplified FCC main column schematic.

157048 Process Control Page 28

Figure 14 Main Column Overview Wash Water

WGC Spillback OVHD Vapor to WGC

CW

Net OVHD Liquid to Gas Con

HCN from Gas Con

Sour Water

LCO from Gas Con Steam HCN Product

HCO from Gas Con Steam

LCO to Flushing Oil LCO Product

HCN to Gas Con LCO to Gas Con HCO to Gas Con Reactor Product Vapor Steam Raw Oil to Reactor

BFW

MCB Product

FCC-PC400

157048 Process Control Page 29

Main Column Bottoms Pumparound Circuit The main column bottoms circulation system (Figure 15) is designed to desuperheat the reactor vapors, condense the bottoms product and scrub entrained catalyst fines from the reactor product vapor. Main column bottoms (MCB), also commonly called slurry oil, is removed from the bottom of the main column and typically pumped to a circulating bottoms/raw oil exchanger and one or more steam generators. MCB may also be used to provide heat to reboilers. The reactor vapors are desuperheated by contact with a large stream of cooled slurry oil on the disk and donut trays. The bottoms flow over the disc and donut trays also washes catalyst fines out of the reactor vapors. The total bottoms circulation rate over the disc and donut trays is normally set at 150% to 200% of the feed rate or 6 gpm per ft2 of column area (15 m3/hr per m2 of column area), whichever is greater. A minimum spillback valve is provided to maintain this minimum flow during turndown operation. This ensures enough circulating MCB is returned to the main column to adequately wet the disc and donut trays, thereby cleansing the reactor vapors of catalyst fines and preventing coke formation on the disk and donut trays due to insufficient liquid flow over the trays. Typically the bottoms temperature is maintained at 670-700°F (354-370°C) to minimize coking and fouling in the slurry circuit. The bottoms temperature is controlled by the LCO product draw rate (or HCO product draw rate if HCO is withdrawn as a product). This flow is adjusted manually by the operators. If the bottoms temperature is too high the LCO product rate is reduced to drop more LCO to the bottom of the column which lowers the bubble point of the MCB product. Alternatively, if a higher LCO endpoint is desired than can be achieved while controlling the bottoms temperature in this manner, a stream of cooled bottoms from the steam generator (quench) may be returned directly to the bottom of the column to sub cool the liquid there. In this manner the temperature in the bottom of the column is no longer composition dependent and the LCO/MCB cutpoint may be varied independently of the bottoms temperature.

157048 Process Control Page 30

Figure 15 Main Column Bottoms and HCO Flow and Control

FIC

FIC

LIC

HCN Stripper Reboiler

Debutanizer Reboiler

Torch Oil Pump Flushing Oil

MCB Quench

Disc and Donut Minimum Flow

FIC

Reactor Product Vapor

Steam FI

FIC

LIC

Raw 

FIC

BFW FIC FIC

MCB Product CW

Raw

FCC-PC401

157048 Process Control Page 31

During normal operations the heat input from the circulating main column bottoms to reboilers and preheat exchangers will be set by product and process considerations. This is true of the heat removed in the other pump around loops as well. Heat input to the steam generators is generally the only variable available to the operator for adjustment of the column heat balance and reflux rates. Increasing the bottoms flow through the steam generators will cause more heat to be removed. This will reduce the amount of hot vapor traffic up the column and eventually will reduce the overhead reflux rate and the heat removed in the overhead system. Each exchanger in the main column bottoms pumparound circuit is designed for oil containing catalyst particles. Main column bottoms flows through the exchanger on the tube side and the velocity should be kept between 3.75 ft/s and 7.0 ft/s (1.14 m/s and 2.13 m/s) for straight tubes and between 3.75 ft/s and 5.75 ft/s (1.14 m/s and 1.75 m/s) for U-tubes. Below the minimum velocity, catalyst can accumulate on the tube walls and slowly plug the tube while greatly reducing heat transfer. If the velocity is above 7 ft/s (2.13 m/s), there is a risk of erosion on the tube walls. If the circulating MCB is very low in ash content ( 5.8 Wt-%), higher liquid volume yields, and capacity increases (1.7) at equivalent conversion. These changes resulted from several major process changes, described below:

157048 Process Variables Page 57

∆Hcomb: ∆Hrecycle:

9055 14,150 BTU/lb coke Zero BTU/lb feed

With the combustor operation, complete CO combustion (< 50 ppm CO in flue gas) can be achieved thermally, without promoter at temperatures greater than 1270°F and 2 mol % excess oxygen in the flue gas. However, Mobil discovered that certain Group VIII metals, particularly platinum, could be used at very low levels (1-3 wt ppm) to effectively catalyze the combustion of CO to CO2 either as an integral part of the fresh FCC catalyst or as a separate additive. In the combustor operation these additives are also frequently used but at lower concentrations than the conventional bubbling bed regenerators. The optimization of the combustor operation is simply the optimization of coke burning kinetics, with a slight twist over conventional bubbling bed systems because of the fast-fluidization regime. The combustor operation is normally quite stable, and with only a little attention from the operator, optimal conditions can be maintained without difficulty. The focus of where to begin lies in only two areas: carbon on regenerated catalyst and afterburning control. If these are both within acceptable values, no further optimization is required. To adjust these values, we need to examine the areas of control for the combustor. Temperature

The primary control point for optimization is the combustor temperature. Although the other factors are important, the temperature can be considered the one truly independent variable for the operator.

157048 Process Variables Page 58

The catalyst recirculation slide valve controls the combustor temperature by adjusting the amount of hot catalyst recycled to the vessel. It is normally operated on TRC control. With adequate excess oxygen, combustor density and normal velocity, an upper combustor temperature of 1275°F should be sufficient. This temperature may need to be increased for the following reasons: •

Dilute temperatures are too high due to afterburn

•

Combustor density is low due to high air rates

•

Flue gas oxygen is low due to blower limits

•

CRC is high due to above situations

The ability to increase temperature in the combustor is, however, limited. Once the recirculation slide valve is full open, no further increase in temperature can be gained with hot recycle catalyst. The effect of other factors, such as the use of CO promoter or increased ∆coke operation, can lead to additional temperature increase. It should be noted that a flow-through catalyst cooler can limit the effectiveness of the recirculation catalyst to raise combustor temperature. The cooled catalyst flow will require additional hot catalyst recycle to maintain desired temperatures. Somewhat compensating is the high ∆coke operation that calls for the catalyst cooler in the first place. However, if it becomes necessary, the cooled catalyst slide valve can be adjusted to help optimize the combustor temperature. Density

Density and velocity together determine the residence time of the catalyst within the combustor. It is desirable to maintain at least 6 lb/ft3 density in the combustor and if possible, 9 - 10 lb/ft3 should be a normal operating target.

157048 Process Variables Page 59

The combustor density is a dependent variable, since the temperature controls the recirculation slide valve, unit capacity determines the velocity based on air rate, and spent catalyst flow is dependent on the catalyst/oil ratio. However, where possible the density (residence time) should be increased for the following situations: •

CRC is high due to insufficient combustion time

•

Dilute temperatures are too high due to afterburn

Increased density will promote better thermal mixing and will increase residence time to help resolve these conditions. Also keep in mind that as density increases, the combustor catalyst inventory increases and the upper regenerator level will decrease. Velocity

Combustor velocity is a dependent variable determined by the amount of air needed to complete the combustion and thus is not a variable available to the operator to optimize other combustor conditions. However, excessive levels of flue gas oxygen (>2.5%) provide little benefit and only serve to increase combustor velocity unnecessarily. A good range for combustor velocity is 4 to 5.5 ft/s. The lower end of the range is generally better when there is a choice. When velocity exceeds 5.5 ft/s, the unit can become combustion limited and increased afterburning may be observed. At high combustor velocities due to capacity demands, CO promoter can be a valuable addition. Although velocities in excess of 6 ft/s have been observed commercially, they have been run in a promoted operation. Combustor velocity may be too high if the following conditions are observed: •

Afterburn increases despite other efforts to minimize

•

CRC increases even with high combustor temperature

157048 Process Variables Page 60

Excess Oxygen

The air rate and regenerator pressure will determine the oxygen partial pressure in the combustor. Although higher levels of excess oxygen via increased air rate are beneficial to the combustion kinetics, they are counterproductive to the unit economics. Velocity will increase as air rate is increased, negating some of the advantages of higher oxygen partial pressure. Thus, flue gas oxygen should be maintained at or below 2 mol%. There are certain situations when an increase in oxygen beyond normal flue gas levels may be warranted: •

Velocity is low and combustor temperature cannot be increased further

•

CRC is high despite other efforts to minimize

One element that can be considered to increase oxygen partial pressure in the combustor without increasing air rate is the use of pure oxygen injection. In certain situations this may be economically feasible. CO Combustion Promoter

If all other conditions in the combustor are optimized, the use of promoter will not be required. However, it is perhaps inevitable that the FCC unit will be pushed to the limits such that the combustor conditions will generally exceed optimum values. In this case, the use of promoter can provide an additional measure of safety for controlling afterburn and avoiding excessive dilute phase temperatures. Particularly when maximum unit capacity is required beyond design levels, the use of promoter can be beneficial. The following conditions might suggest that promoter should be considered: •

Velocity is high, density is low and recirculation catalyst is maximized

•

Dilute temperatures are high due to the previous conditions

•

Afterburn is high despite high combustor temperatures

157048 Process Variables Page 61

Maximum Capacity

Since many units operate on the basis of pushing the FCC capacity to the limit, it is appropriate to discuss what will happen in the combustor. Prior comments on the operating variables have addressed some of the concerns of higher capacity operation, but an overall perspective is needed. In general, pushing capacity to the maximum will result in a combustor operating at high velocity, low density and low catalyst residence time. Directionally, this will tend to increase afterburn and carbon on regenerated catalyst. From earlier data shown regarding coke burning kinetics, carbon is reduced to levels around 0.15 wt% very quickly, but to get to 0.05 wt% requires substantial additional time. In a maximum capacity operation, there may not be sufficient time in the combustor to reach carbon levels of 0.05 wt% or below. It is important to point out that this is not all bad, and in fact may be economically advantageous. With the use of promoter to maintain control of after burning, it is possible to allow the CRC to increase as a trade off for more feed capacity. Although effective catalyst activity is reduced, the reduction may be small enough to justify the higher capacity operation. Optimization Summary

All of the above variables need to be thoroughly understood, especially how they interact with each other. Application of this knowledge properly will result in arriving at the best operation for each particular unit. Optimization should be considered as a continuous effort, since what is optimal in one set of circumstances may not be in another. As a final note, it should be emphasized that UOP is always interested in obtaining feedback on the operation and experiences of these units from the refiner. Our Technical Service department is always available to provide assistance or consultation as needed. Only through working together can we hope to continue to improve our designs and their performance capability.

157048 Process Variables Page 62

MAIN COLUMN BOTTOMS AND SLURRY SETTLER

INTRODUCTION

The main column bottoms system on an FCC presents some unusual mechanical and operating problems. The vapors entering the column are superheated and contain catalyst fines which may cause erosion or plugging. Heavy oil cascading down the disc and doughnut (side-to-side on some units) trays cool and condense the heavy vapors so that they can be fractionated. The catalyst fines are washed down to the bottom of the column by this cascading stream.

COKE

Many units have some coke buildup in the reactor, vapor line, and bottom of the main column. The amount in the reactor and vapor line can be minimized by good insulation. The coke buildup in the main column is influenced by three factors: 1.

Hydrocarbon characteristics

2.

Temperature

3.

Residence Time

Some hydrocarbons have a greater tendency to thermally crack and produce coke than others. The Conradson or Ramsbottom carbon residue tests may be used to get some idea of this tendency, but it is difficult to compare one stock with another. The amount of catalytic cracking that has taken place in the reactor will influence the coke production of the hydrocarbons in the main column. The operation of the FCC unit and of the upstream units that produce FCC feed are controlled by other, more important variables than the main column bottoms coke make. Temperature and residence time control are used to minimize the bottoms coke problem.

157048 Process Variables Page 63

The maximum allowable temperature in the bottom of the tower is usually given as 700°F (370°C). Experience with a particular feed may raise or lower this somewhat, but it is better to run lower to prevent coke problems. Many refiners use 680-700°F (360-370°C) as the normal operating point. The major control on the bottoms temperature is by the composition of the material, but the quench line may also be used to sub-cool the bottoms material. This line returns cool oil directly to the bottoms level, instead of to the disc and doughnut trays. It is generally used at high throughput, when there is adequate flow to the discs and doughnuts, and the heat input to the column is high. Residence time refers only to the time spent in the column, not in the entire slurry system. The oil begins to cool as soon as it leaves the tower. There may be some minor coking in the slurry settler and associated piping, but usually this is not enough to affect plant operations. If the oil is in the tower too long, at too high a temperature, serious coking can plug the bottom. Removal is difficult, because chipping away with hammers usually leads to lining damage. Other methods, such as chemical cleaning, are usually ineffective. At low charge rates, the bottoms circulation through the exchangers must be reduced to heat balance the tower. The oil cascading down the discs and doughnuts decreases to a point where the trays run dry. Because the hot reactor vapors are not cooled, they can cause warping and distortion of both the disc and doughnut trays and those above as this large volume of vapors tries to pass up the column. Catalyst fines will be carried up into the HCO and LCO circuits, which can be quite serious, because these areas are not designed for fines removal. Product specifications will be adversely affected. The solution to this problem and that of long residence times at low circulation rates is to use the minimum flow line which bypasses the exchangers and returns oil directly to the column, over the disc and doughnut trays. The flow over these trays should be at least 50 gpm/ft tower ID (37.3 m3/hr/meter tower ID). This line is usually used on startup, when the charge rate is low and the tower is cool. As

157048 Process Variables Page 64

charge rate increases, the flow rate through the minimum line is decreased to force oil through the heat exchange circuits. There should always be a small amount of oil flowing through the line, so that the oil does not set up. The same is true for the quench line to the bottom of the fractionator. Some units have used a large impingement baffle to break up the reactor vapors as they enter the column. Many of these baffles proved to be a good starting point for coke, and the baffles were removed. Inevitably, there is some coke buildup; pieces may break off with thermal shock or other stresses. If a chunk of coke could enter the bottom line, it could cause plugging or pump damage. A coke trap made of Type 405 or 410 stainless steel is used at the bottom of the column to keep these large chunks out of the line. Smaller pieces of coke pass through it and are caught by the pump suction screens. These can be cleaned on stream, while the larger chunks of coke might prove difficult to remove from the line.

CATALYST CARRYOVER

The amount of catalyst carryover from the reactor will depend on unit design, cyclone efficiency, catalyst type, and unit throughput. Unusual problems such as high reactor level, cyclone failure through cracks or plugged diplegs, or pressure surges can increase the catalyst carryover to unacceptable levels. A conventional unit of older design (single stage cyclones) should lose less than 0.4 LB/bbl charge over to the main column, with most of this returning with the slurry recycle. A new unit with a riser "Tee" and high efficiency cyclones should lose less than 0.05 LB/bbl. It is not practical to actually measure the catalyst content of the main column inlet vapors. The catalyst content of the circulating bottom stream is, however, a good indication of the catalyst losses. A general assumption may be made that very little catalyst is entrained up into the HCO or LCO products. If this is not the case, the bottoms circulation over the disc and doughnut trays should be increased as much as possible, within the limits imposed by other operating variables. The lower part of

157048 Process Variables Page 65

the tower should be inspected carefully at the next turnaround to determine if the problem is here. The main column bottoms pump discharge manifold is constructed to direct the catalyst fines and coke particles out of the circulating system. The lines to the various exchangers come off the top of the manifold. The line to the slurry settler will come off the end of the manifold from the bottom. If there is no settler, the reactor recycle and bottoms product lines will both come off the bottom of the manifold. This obviously will not achieve a complete removal of catalyst fines or coke, but it will help prevent a buildup in the main column bottoms circulating system.

SLURRY SETTLER

The slurry settler process flow is shown in Figure 16 in Process Control section. Main column bottoms enter the settler through a tangential nozzle; this gives it a swirling motion that promotes a more even distribution of the heavy oil as it moves up towards the outlet. The catalyst fines settle out and are carried back to the reactor. The carrying medium will depend upon the operation. 1.

Low activity catalyst with large amounts of recycle:

This would be typical of older units that require the high recycle rates to get desired conversion. The recycle consists of main column bottoms and HCO, which is cooler and has a higher flowing specific gravity than the bottoms material. The HCO is injected into the upper diluent point and flows down with some of the bottoms and most of the catalyst that enters the settler. A plant operating in this mode would have a CFR of 1.2 to 2.0.

157048 Process Variables Page 66

2.

Higher activity catalyst with little or no recycle:

The use of the new higher activity catalysts and better reactor design enable the refiner to crack most of the feed on the first try. The heavy oil product make is lower; it is also a more refractory material which tends to go to coke and dry gas when recycled to the reactor. The carrying material used in this case may be HCO, but for many units, raw oil is the better choice. The bottom diluent injection point is used to minimize the amount of raw oil that will go up into the settler. If the raw oil thermally cracks or goes out with the CSO, it may cause problems with the CSO product specifications. To help prevent this, the flow back to the reactor should always be higher than the diluent flow to the settler.

SETTLER OVERHEAD

The amount of CSO that comes off the top of the settler will also depend upon the operation. The decreased heavy oil make of the newer units can lead to a buildup of catalyst in the main column because there is less main column bottoms leaving the tower to take it out. This refers to oil leaving the system, not to the bottoms circulation streams. The concentration of fines in the tower may build up enough to cause serious erosion and plugging problems. The most effective solution is to use a return line from the top of the settler to the main column disc and doughnut trays. The flow to the settler can then be increased by the amount of oil which is returned to the main column relatively free of catalyst. Typical settler flow rates are shown in Table 11. For this table, it was assumed that catalyst is coming overhead from the reactor at a rate of 100 LB/day. The catalyst can only leave the system through the outlet to the settler; therefore, the concentration will be equal to the amount of catalyst (100 LB/day), divided by the flow to the settler.

157048 Process Variables Page 67

TABLE 11 SLURRY SETTLER FLOW CASE A

CASE B

CASE C

10,000

10,000

10,000

Main Column Bottoms To Slurry Settler, BPD

1500

500

1500

Main Column Bottoms Catalyst Concentration, lb/bbl

0.067

0.20

0.067

500

500

500

CSO Product, BPD

1,000

500

500

Recycle to Reactor, BPD

1,000

500

500

0

0

1,000

Feed to FCC, BPD

Diluent, BPD

CSO Return to Main Column, BPD

Case A would be an operation with lower activity catalyst and a higher CSO product make. Case B would be a conversion to higher activity operation, decreasing the amount of recycle and the CSO product. Neither Case A or B uses a return line from the top of the settler to the main column. Case C is similar to B in that a higher activity catalyst is used, with a small amount of main column bottoms produced. The use of the return line from the top of the settler back to the main column allows the flow to the settler to be increased to 1500 BPD. A limit on flow to the settler is the velocity of fines settling. The liquid velocity should never exceed 30 BPD/ft2 (50 m3/d/m2) of settler cross sectional area. Above this, there may be problems with catalyst carry over into the overhead stream.

157048 Process Variables Page 68

FCC UNITS WITHOUT SLURRY SETTLER

Some of the new units have incorporated a reactor cyclone configuration which decreases the amount of catalyst carryover to the main column. The reactor riser ends with a pair of outlet arms, and a two stage cyclone system on the reactor outlet eliminates most of the catalyst in the overhead vapors. With this system, some refiners have elected not to use a slurry settler. The recycle to the reactor and the bottoms product come directly from the main column bottoms pump discharge manifold.

PLUGGING

There are occasional problems with plugging in the lines or exchangers of catalyst bearing streams. If there is a plant upset which causes large amounts of catalyst to go overhead, such as a sudden dip in column pressure, immediate action should be taken to remove the extra catalyst from the system. Increased slurry recycle and clarified oil product would be the two most important steps. These should be continued until the laboratory confirms that the BS & W content of the bottoms stream is back to normal. In the event of a major upset that completely fills the bottoms of the column with catalyst, care must be taken to avoid further complications. Catalyst holds heat fairly well and conducts it poorly, so the cool-down period may be on the order of several days. The thermocouples in the column will be well insulated by the catalyst close to the wall, so the readings will probably be lower than the actual temperature of most of the catalyst. Introduction of cold oil or wash water will produce severe pressure surges that may damage the column internals. The oil soaked catalyst is both a fire and breathing hazard. The shutdown for cleanout can be estimated at one week, a good argument against hasty measures that could lead to excessive catalyst carryover to the column.

157048 Process Variables Page 69

MAIN COLUMN BOTTOMS EXCHANGERS

To prevent catalyst plugging or erosion in the exchangers, UOP calls for the following velocities in the tubes: 1.

Straight tubes: maximum velocity 8.0 ft/sec, minimum 4.00 ft/s.

2.

U-tubes: maximum velocity 8.00 ft/sec, minimum 4.00 ft/s.

In general, the optimum velocity is 8.00 ft/s. Straight tube construction is recommended. It is important to think of these numbers when changing raw oil charge or exchanger flow rates. Plugged exchangers are difficult to clean. A catalyst bearing stream is never routed through the shell side of an exchanger because the catalyst fines will settle to the bottom of the exchanger. There will be a progressive loss of heat transfer area as more and more tubes are covered by the fines.

MAIN COLUMN BOTTOMS HEAT REMOVAL

The main column bottoms circulation rate is adjusted to control the column's heat removal requirement. The bottoms stream generally exchanges heat with the raw oil feed and is utilized in the production of superheated steam in the steam generators. A reduction or increase in the bottoms heat removal must be compensated by an increase or reduction in heat removal in another section. Provided no other changes are made, the overhead reflux rate will compensate for any changes in bottoms heat removal. As illustrated in Table 12, a decrease in the bottoms stream heat removal results in an increase in reflux rate. The bottoms stream heat removal should be adjusted to minimize the reflux rate and maintain good product distillations.

157048 Process Variables Page 70

TABLE 12

Feed Rate, BPD

17550

17550

Overhead reflux rate, BPD Net ovhd light gas yield, MMSCFH Net ovhd liquid yield, BPD Ovhd heat removal, MM-BTU/hr

10488 13.81 7973 42.81

11581 13.81 7993 45.24

Net heavy naphtha yield, BPD Circ. heavy naphtha heat removal MM-Btu/HR

2782

2782

12.32

12.32

1790

1790

10.17

10.17

967

967

31.0

28.57

Net LCO yield, BPD Circ. LCO heat removal, MM-BTU/hr Net bottoms yield, BPD Circ. bottoms heat removal, MM-BTU/hr

GASOLINE/DISTILLATE PRODUCTION

The main column draw temperatures are dependent on the stream's composition and vary with changes in draw rate. As the gasoline product draw is reduced, liquid will drop down the column to the LCO draw tray and require an increase in LCO product draw. The reduced gasoline draw rate will result in a lighter gasoline product having a lower ASTM distillation end point and a lower draw tray temperature. The LCO product also becomes lighter because of light material dropping to the LCO draw tray. The LCO initial boiling point temperature will decrease with a slight decrease in the draw tray temperature. Table 13 illustrates the changes which result to the main column and product streams as the LCO yield is increased by reducing the gasoline endpoint and yield.

157048 Process Variables Page 71

TABLE 13

Gasoline: API Product rate, BPD 90% BP temp., °F RONC MC overhead temp., °F

57.9 15800 372 90.7 311

57.2 16200 381 90.6 316

56.3 16600 392 90.4 323

Light cycle oil API Product rate, BPD Flash point, F 10% BP temp., °F 90% BP temp., °F End point, F MC draw temp., °F

20.1 6000 198 462 624 664 482

19.8 5600 204 471 624 664 485

19.3 5200 212 482 624 664 499

Slurry °API Product rate, BPD

6.2 2800

6.2 2800

6.2 2800

GASOLINE CUT PROPERTIES

Gasoline or any other liquid stream can be broken up into numerous cuts, each having distinct properties. Examining the cuts which comprise a typical gasoline sample will show how overall product quality can be improved.

157048 Process Variables Page 72

A FCCU gasoline sample from a high severity operation was broken up into nine cuts, each having a narrow TBP range as illustrated in Table 14. The properties of the individual cuts are shown in Figures 6 and 7. TABLE 14 Cut No.

1 2 3 4 5 6 7 8 9

Cumulative Vol -% 20 30 40 50 60 70 80 90 100

TBP Fraction, °F 75 - 93 93 - 145 145 - 181 181 - 210 210 - 250 250 - 286 250 - 286 286 - 334 334 - 387 387+

157048 Process Variables Page 73

FIGURE 15: FCC GASOLINE CUT PROPERTIES Avg. BoilingPoint Temp, °F

400

300

200

100

0 1

2

3

4

5

6

7

8

9

6

7

8

9

6

7

8

9

Gasoline Cut 100

API Gravity

80

60

40

20 1

2

3

4

5

Gasoline Cut 100

RONC

96

92

88

84

80 1

2

3

4

5

Gasoline Cut

157048 Process Variables Page 74

FIGURE 16: FCCU GASOLINE CUT PROPERTIES 1.2 1

Sulfur, wt %

0.8 0.6 0.4 0.2 0 1

2

3

4

5 Gasoline Cut

6

7

8

9

5

6

7

8

9

100

Bromine

Number

90 80 70 60 50 40 30 20 10 0 1

2

80

3

4

Gasoline Cut

Paraffin/Naptha

70

Aromatic

Olefin

Liquid Vol %

60 50 40 30 20 10 0 1

2

3

4

5

Gasoline Cut

6

7

8

9

157048 Process Variables Page 75

The sulfur content of gasoline increases sharply in the last 387+ TBP fraction as indicated in the Wt-% SULFUR graph. The overall sulfur content of gasoline hence could be reduced by lowering the gasoline end point temperature. The RONC graph of the various cuts indicates a reduction in octane at CUT 4 (181-210 TBP) and CUT 9 (387+ TBP). A slight increase in the overall gasoline octane can be obtained by reducing the end point temperature of the gasoline. If it were possible to remove CUT 4, gasoline octane could be further increased.

157048 Process Variables Page 76

ROUTINE PROCESS VARIABLE CONTROL REACTOR REGENERATOR SECTION

This tabulation of process conditions is intended to assist the operator in selecting the optimum operating conditions for different operations. It may be noted here that process units rarely operate at their design conditions. Variable

Operating Conditions

Raw oil charge rate

As desired.

Raw oil temperature

To balance coke yield, conversion, and RON requirements.

Slurry recycle rate

Normally at minimum.

HCO or raw oil to slurry settler

Equal to or slightly less than total flow of slurry recycle to reactor, or until clarified oil gravity changes.

Heavy recycle rate

Heavy recycle rate can be varied to adjust conversion product yields or increase coke yield.

Combined feed temperature

Normally as high as possible provided neither reactor nor regenerator temperature is excessive.

Reactor temperature

As necessary to obtain desired conversion and RON.

Reactor pressure

Equals fractionation column receiver pressure plus fractionation pressure drop.

Reactor level (Riser cracking operation)

To cover top stripping grid.

Reactor level (Bed cracking operation)

Minimum level needed to achieve desired conversion.

Emergency steam to riser

Used to initiate catalyst circulation on startup and to avoid plugging the riser on an emergency shutdown. Normally not used.

157048 Process Variables Page 77

Variable

Operating Conditions

LCO or gasoline to riser

Used to control regenerator temperature when the unit is "behind in burning".

Stripping steam

Just enough to strip catalyst. This value can be arrived at by observing the effect of decreasing the stripping steam on regenerator temperature. 1.5-2 lb/1000 lb catalyst circulation is typical.

Steam to feed nozzle

Normally adjusted to maximize feed distributor pressure drop within feed pump hydraulic constraint. Usually 1-2 wt% of feed.

Regenerator air rate

As necessary to control regenerator temperature spread or to give good control using automatic snort. On total CO burning units to obtain 1-4% oxygen in flue gas.

Regenerator dense phase temperature

Normally cracking conditions are varied to optimize regenerator dense phase temperature. Too cold and catalyst will not be well regenerated. Too hot and reaction with oil will be thermal with resultant loss of gasoline and increase in gas.

Regenerator dilute phase or flue gas temperature (on conventional partial CO burning units)

The difference between regenerator dense, dilute and flue gas temperatures is an indication of the amount of excess oxygen present, and is the criterion by which the air rate is varied.

Regenerator pressure (conventional)

Equals the reactor pressure plus the reactor regenerator differential pressure.

Regenerator pressure (with flue gas power recovery)

Regenerator pressure controlled, reactor-regenerator differential allowed to swing within reasonable limits.

Regenerator level

The catalyst level can vary from about 35" to 80" of H2O with the unit inventory and regenerator velocity.

Reactor-regenerator differential pressure

Varied to obtain stable slide valve differentials and minimum utility consumption.

157048 Process Variables Page 78

Variable

Operating Conditions

Slide valve differential pressure

Dependent on vessel differential and catalyst condition. Over-rides are normally set at 1-2 psi.

Steam to spray and torch nozzles

Flow will be almost zero except when sprays are in service.

Torch oil rate

Used on startup and to control afterburning on partial CO combustion units.

Torch steam pressure

Only used when necessary to atomize torch oil and normally 5-10 psi higher than regenerator pressure.

157048 Process Variables Page 79

TABLE 15 MAIN COLUMN Variable

Operating Conditions

Main column bottoms temperature

Adjust clarified oil yield and quench to hold long enough to prevent coking.

Circulating slurry rate

As necessary to control fractionating column heat removal requirements. To minimize top reflux to obtain minimum gap (25-40°F) on gasoline between 90% and E.P.

Slurry return temperature

Of little interest. Dependent on cleanliness of exchangers, number in service and circulation rate.

Clarified slurry yield, or main column bottoms yield if no slurry settler

As necessary to control fractionating column bottoms temperature and level.

Main column bottoms BS&W

Adjust flow to slurry settler, clarified oil return to main column, and slurry recycle. If no slurry settler, adjust bottoms recycle and bottoms product.

Heavy recycle oil circulation

Rate is set as desired to transfer heat to various reboilers.

Heavy recycle deck temperature

Depends on distillation range on heavy recycle and on tower pressure.

Light cycle oil yield

Depends on charge rate and conversion, and is varied to maintain desired properties of light cycle oil. Also used to control bottoms level.

Flush oil

As required to keep catalyst out of instruments. Normally 1500-2000 BPD, but this will vary between different units.

157048 Process Variables Page 80

Variable

Operating Conditions

Light cycle oil stripping steam

Enough to meet flash point specifications.

Light cycle oil deck temperature

Depends on distillation range of light cat gas oil and on tower pressure.

Unstabilized gasoline top reflux rate

Is varied to control tower top temperature and depends on amount of heat removed lower in column.

Circulating top reflux temperature

Depends on water temperature and flow rates.

Fractionator column top temperature

Varied to control endpoint of unstabilized gasoline.

Overhead receiver pressure

Can be varied as discussed.

Overhead receiver temperature

Should always be as cold as is economically possible.

Unstabilized gasoline yield

Depends on charge rate and conversion.

Wet gas molecular weight

Minor variations in density will be due to changes in receiver conditions, but major changes will be due to increased hydrogen production. At low densities, the compressor will have an increased tendency to surge.

Wet gas flow

Dependent on compressor speed but must be adequate to handle production plus spillback for control. Must be above minimum to keep out of surge.

157048 Process Variables Page 81

TABLE 16 GAS CONCENTRATION SECTION Variable

Operating Conditions

Wet gas compressor Variable speed

Run at minimum governor until spillbacks close.

Fixed speed centrifugal

Butterfly valve opens away from limiting stop as spillbacks close.

Fixed speed reciprocating

On spillback control.

Wet gas spillbacks High pressure separator Temperature Pressure Primary absorber Top and intercoolers temperatures

Vary as needed to hold main column overhead receiver and interstage pressure. Used to keep centrifugal machine out of surge. 80°F-100°F (27°-38°C) Rides on primary absorber backpressure. Less than 100°F (38°C)

Intercooler flow rates

As needed for good absorption efficiency.

Pressure

Rides on sponge absorber backpressure.

Sponge absorber Top temperature

As cool as economically practical.

Lean oil flow rate

As needed for good absorption efficiency. Do not flood tower.

Pressure

Controlled to give good absorption efficiency and hold correct backpressure on HPS.

157048 Process Variables Page 82

Variable

Operating Conditions

Stripper Overhead vapor flow rate

Controlled by heat input to column to give sufficient C2- and H2S stripping.

Heat input to column

Controlled to give proper overhead vapor rates.

Pressure

Rides on high pressure separator backpressure.

Debutanizer Top temperature

Controls reflux to give desired RVP of gasoline.

Top reflux temperature

Should be as cold as economically possible.

Reboiler heat

As required to give good fractionation.

Pressure

Varied as needed for good fractionation, but must be high enough to allow condensation of C3-C4 stream by air fin fans or cooling water.

157048 Process Calculations Page 1

PROCESS CALCULATIONS INTRODUCTION Throughout the years the Fluid Catalytic Cracking process has been a very versatile and flexible tool for the refiner, and has become the basic conversion step in the modern refinery. This process has survived and prospered because of its ability to handle the many changes in catalyst, operating conditions, and feedstocks that have occurred over the years. The FCC Unit produces large volumes of high octane gasoline, olefinic LPG, fuel oil (LCO and MCB), fuel gas, steam, and electricity. The yields are mainly determined by process variables (i.e. feedstock, operating conditions, mechanical features, and type of catalyst). Process variables have varying degrees of interdependence and may change frequently producing changes in the yield structure of the products. A performance test conducted at least once a week is recommended to evaluate the effect of process variables on yield. The tests can be used to chart a history of the unit and to find conclusions at different operating conditions. The performance test provides accurate yield structures at a particular set of operating conditions and provides a base point for further testing. The Performance Test normally includes a heat balance, material balance, and a pressure survey. In those cases where more information is desired, a Mechanical Evaluation Test is recommended. The refiner can use this test to assess the potential of the unit and determine possible bottlenecks. This section explains how to accomplish an acceptable Heat and Material balance and how to do some of the most important calculations in the FCC Unit. An FCC Performance Test Procedure is available on request from the Technical Service Department. This procedure explains in detail how to conduct a performance test in the FCC Unit.

157048 Process Calculations Page 2

MATERIAL BALANCE A material balance on an FCC Unit is done by drawing an envelope around the unit in a manner that flow rates are known for all streams. This envelope includes the Reactor, Regenerator, Main Column, and Gas Concentration sections. Normally, the Gas Concentration Unit includes the primary absorber, sponge absorber, stripper column, and debutanizer column.

1.

Data

Flow rates, flowing temperatures and laboratory analyses are required for each stream. Pressure is also needed for the gas streams. The following table shows the information needed to do a material balance: INPUT DATA FOR HEAT AND MATERIAL BALANCE Stream

Flow Temp. Pressure

API

Distillation

Feed

Yes

Yes

Yes

D-1160

Air*

Yes

Yes

MCB

Yes

Yes

Yes

D-1160

Yes

LCO

Yes

Yes

Yes

D-86

Yes

Gasoline

Yes

Yes

Yes

D-86

LPG

Yes

Yes

Sponge Gas

Yes

Yes

Meter Factor Yes

Yes

Yes

Yes

Yes

Yes

Yes

Yes

Yes

Yes

Yes

Flue Gas Notes

GC

*Ambient temperature and relative humidity of air are needed. Reactor, Regenerator, and combined feed temperature are needed for Heat Balance

The method for including extraneous streams is straight forward as long as flow rates and analyses are known. These additional streams coming into the Main Column and Gas Concentration Unit are subtracted from the product streams.

157048 Process Calculations Page 3

2.

Liquid Streams

Calculate corrected liquid flow rates using the following equation: Q = K R (Gf)1/2 / Gb Where:

Q = flow rate K = flow meter constant ("K" factor) R = chart reading Gb = base gravity @ 60°F Gf = gravity at flowing temperature

The flowing gravity Gf is calculated using the following equation: Gf = Gb x VCF Where:

VCF = Volume correction factor VCF = EXP [ (-ßo) ∆T(1+0.8 ßo ∆T) ]

0.9545

Where: Go = density at 60°F in kg/m3 T = observed temperature in °F ∆T= T - 60 ßo = coefficient of thermal expansion at 60°F, (1/°F)

926 173 113

The set of correlations for the coefficient of thermal expansion based on API Data Tables are: ßo = (Ko + K1 Go + K2 Go² )/ Go² Where:

0.00039779

157048 Process Calculations Page 4

Product Crude Oils

°API Range 0 – 100

Ko

K1

K2

341.0957

0

0

0.2438

0

Gasoline

52 – 85

192.4571

Gasoline/Jet

48 – 52

1489.0670

0

-0.0018684

Jet Fuels

37 – 48

330.3010

0

0

Fuel Oils

0 – 37

103.8720

0.2701

0

0.3488

0

Lubricating Oils

-10 – 45

0

For FCC Raw Oil Feedstock (VGO); Ko = 341.0957, K1 = 0, K2 = 0 The following rounding is applied to the input and output of all routines. Temperature: Density: VCF:

0.1°F 0.1 °API or 0.5 kg/m3 five significant figures for computation

Example: Raw Oil Charge K = 3,380 in BPSD R = 8.9 Gb = 0.9260 VCF = 0.9545 Gf = 0.9260 x 0.9545 = 0.8839 Q = 3,380 x 8.9 x (0.8839)1/2 / 0.9260= 30,542 BPSD Q = (BPSD) x 5.614583 ft3/bbl x (Gb x 62.3635 lb/ft3 H2O)/24 hr/day Q = 412,615 lb/hr Similar calculations are done for all C5+ liquid streams. LPG streams are treated as follows:

157048 Process Calculations Page 5

Estimation of VCF for LPG The following equation approximates VCFs from API Tables #33 and #34:





VCF   10  A*Gb B   T  60   1.0 Where: T Gb A B

= = = =

Flowing Temperature, °F Specific Gravity @ 60°F, in g/ml 2.64641798 1.40583481

157048 Process Calculations Page 6

3.

Gas Streams

Calculate corrected gas and air flow rates as follows: Q = K x R x (Pf/(Tf x SG))1/2 Where:

Q = flow rate K = flow meter constant ("K" factor) R = flow meter reading Pf = Pressure at flowing conditions (absolute) Tf = Temperature at flowing conditions (absolute) SG = specific gravity of gas = MWgas/MWair (1.0 for air)

Example: Sponge Gas Q = K x R x [(Pf/(Tf x SG))]1/2 Where:

K = 91,848 scfh R = 6.5 chart reading Pf = 173 psig + 14.7 = 187.7 psia Tf = 113°F + 460 = 573°R SG = 0.7054 from Sponge gas calculations MW = 18.9 from Sponge gas calculations

Q = 91,848 x 6.5 x [(187.7)/(573 x 0.7054)]1/2 = 406,837 scfh M = Mass Flow = scfh x MW/379.67 = 20,261 lb/hr Where:

379.67 is the conversion factor for scf / mol

Similar calculations are done for all gas, vapor, and air streams.

157048 Process Calculations Page 7

4.

Calculate Coke

The coke make is calculated from the Heat Balance. (Refer to Heat Balance calculation section.)

5.

Calculate the as Produced Yields

A product yield is defined as the product rate divided by the raw oil rate. The volume percent of each product stream is: Vol-% (A) = (A, bpsd)100/Fresh Feed, bpsd The weight percent is: Wt-% (A) = (A, lb/hr)100/Fresh Feed, lb/hr Where:

A = any product stream

157048 Process Calculations Page 8

6.

Weight and Liquid Volume Recoveries

Once the weight and the volume flows are known for each stream, the Weight recovery and the Liquid volume recovery can be calculated. Proper data analysis requires that the Weight recovery must be 100.0  2.0 wt-%. Errors outside this range are significant and cast doubts on the validity of the test data. The Sponge Gas used to calculate the Weight recovery should not include the inert gases (N2, O2, CO, CO2). In addition, the as-produced Liquid Volume recovery does not include the C3+ from the sponge gas. Weight % Recovery = (Products, lb/hr x 100)/(Fresh Feed + Extraneous Feeds) lb/hr Liquid Volume % Recovery = (Products, bpsd x 100)/(Fresh Feed + Extraneous Feeds)bpsd

7.

Conversion

Conversion is defined as the volume percentage of raw oil converted to gasoline and lighter components. This is calculated as: Conversion, Vol% =

Feed - LCO - HCO - MCB  100 Feed

This conversion is called ‘as-produced’ or ‘apparent’ conversion because is not corrected for cut-points. The ‘corrected’ or ‘true’ conversion is calculated using the same equation after the gasoline and LCO yields are corrected for cut-points. 8.

Gasoline and LCO Yield Adjustments

It is important to correct the gasoline and LCO yields on a constant cut-point basis. It is inaccurate to compare gasoline yields at different 90% or EP temperatures. The gasoline yield must also be adjusted by removing the C4's and adding the C5's and C6's from the LPG and Sponge Gas streams. The procedures to adjust the liquid yields and to calculate the C4's in the gasoline are attached in this section.

157048 Process Calculations Page 9

9.

Gasoline Selectivity

The gasoline selectivity is the corrected gasoline yield divided by the true conversion: Gasoline Selectivity =

10.

(Corrected Gasoline Yield)  100 True Conversion

LPG and Sponge Gas Calculations

These procedures use the mol percentage from the GC analysis, the molecular weight, and the specific gravity to calculate the flow rate of each component. Also, the stream specific gravity and molecular weight are calculated. The procedure is attached in this section.

11.

C3 and C4 Recovery

The C3 and C4 recovery indicate how the Gas concentration is performing. The C3 recovery is calculated as:

C3 Recovery, Vol% =

C3 in LPG  100 (C3 in LPG + C3 in Fuel Gas)

The C4 recovery is calculated as: C4 Recovery, Vol% =

C4 in LPG  100 (C4 in LPG  C4 in Fuel Gas  C4 in Gasoline)

157048 Process Calculations Page 10

FCC Unit Material Balance Refiner: _______________________________

Location: _______________________

TI - Tag

FE - Tag

Gravity

Vol.Corr.

Flowing

Meter

Temp. °F

Readings

Gb@60°F

Factor

Gravity Gf

"K"

1

Fresh Feed (FF)

TI—10

2

Main Col Bottoms (MCB)

TI-20

lb/Hr

Vol%

Wt%

8.90

0.9260

0.9545

0.8839

3,380

30,538

412,601

4.80

1.0412

0.8806

0.9169

418

1,843

28,005

6.04

6.79

9.00

0.9200

0.9838

0.9051

700

6,515

87,450

21.33

21.19

7.60

0.7599

0.9489

0.7211

2,118

17,987

199,434

58.90

48.34

7.20

0.5612

0.9653

0.5418

632

5,964

48,833

19.53

11.84

0

0.00

0.00

FRC-20 462

TI-30

Light Cycle Oil (LCO)

BPSD

FRC-10 173

3

Date: __________

FC-30 101

4

Gasoline (DebBt)

TI-40

FRC-40

5

LPG

TI-50

6

LPG (ELPG) Extraneous Feed

TI-

7

Coke

See attached Heat Balance calculation sheet.

148 FRC-50 87 FRC-

0

25,187

6.10

6,589

18,821

4.56

6,996

20,887

0

0

PE-Tag Press, psig 8

TI-60

Sponge Gas (SGas)

FC-60 113

8

Gas (HGas) Extraneous Feed

TI-

9a

Air to Regenerator (Dry Back)

TI-70

9b

9

1.00

5.00

1.00

W/o Inerts

0.00

44.5

46,000

81,721

374,251

56

325

520

2,382

56

635

1,260

5,771

83,502

382,405

PI-72

FIC-74 230

1,518

PI-70 6.83

FIC-72

TI-74

Air to Distributor 2

W/o Inerts 173

PRC-

FIC-70

230 9c

0.6519

FC-

TI-72

Air to Cat. Cooler

PRC-60 6.50

415

SCFM

PI-74 6.20

1.00

Total Air to Regen (Wet Basis)

As produced Calculations

10

Weight Rec Inert Free

= [MCB lb/hr+LCO lb/hr+DebBt lb/hr+LPG lb/hr+SGas lb/hr+Coke lb/hr]*100/ [FF lb/hr+ELPG lb/hr+EGas lb/hr] =

=

98.82

Wt%

11

Liquid Vol. Recovery

= [MCB bpsd+LCO bpsd+DebBt bpsd+LPG bpsd]*100/[FF bpsd+ELPG bpsd]

=

105.8

Vol%

12

Conversion

= [FF bpsd-MCB bpsd-LCO bpsd]*100/[FF bpsd]

=

72.6

Vol%

For Liquids

:

BPSD = Units K [SQRT(Gf)]/Gb

:

lb/hr = [BPSD (Gb) (5.6146ft3/bbl) (62.3689lb/ft3)]/(24h/d) = (14.591) ( BPSD) (SG)

For Gases or Air

:

SCFM = K Units SQRT{psia/(°R*Gb)}

LV%

= (stream, bpsd)(100)/(FF, bpsd)

lb/hr = (scfm MW 60min/hr)/379.5 Wt% = (stream, lb/hr)(100)/(FF, lb/hr)

uop1292rc

157048 Process Calculations Page 11

FCC Unit Material Balance Continuation Refiner: _______________________________

Location: ____________________

Yields Adjustment for Gasoline @ 380°F - 90% and LCO @600°F - 90%. IBP, °F 10%, °F 70%, °F 90%, °F EP, °F Gasoline LCO

319

460

MCB

524

710

569

Date: _____________

°API

BPSD

lb/hr

367

410

54.71

17,987

618

653

22.30

6,515

87,450

4.40

1,843

28,005

26,346

314,889

Totals

199,434

Gasoline Gasoline Factor (F1)

= [(380-T90)/(EP-T90)](1/9)+1

= [ ( 380 – 367 )/(410-367)]

(1/9) + 1=

LCO Factor (F2)

= [(380-IBP)/(T10-IBP)](1/9)

=

(1/9) =

cGasoline, BPSD

= (Gasoline, bpsd)(F1) + (LCO,bpsd)(F2)

[ ( 380 - 319 )/(460-319)]

=

18,587

1.034 F1 0.05 F2

cBPSD

60.9

cLV%

c °API

= API + 37.5(LV% - cLV%) / LV%

=

52.13

c °API

0.771

cSG

cGasoline, lb/hr

= cBPSD x cSG x 14.591

=

208,974

c lb/hr

50.6

cWt%

lb/hr

BPSD

C4's in Gasoline

=

3,390

380

C5's + C6's in LPG

=

423

46

C5's + C6's in SGas

=

248

26

C5's + C6's in Extr Feeds

=

0

0

Gasoline Yield Adjustment for C4's, C5's, & C6's.

13 Corrected Gasoline, BPSD

= cGasoline-C4's+C5's+C6's Extr(C5+C6)

=

18,587

corrBPSD

60.9

corrLV%

14 Corrected Gasoline, lb/hr

= cGasoline-C4's+C5's+C6's-Extr(C5+C6)

=

208,974

corrlb/hr

50.6

corrWt%

15 Corrected °API

= ((141.5*BPSD*14.591)/(lb/h*24))-131.5

=

52.13

corr°API

LCO LCO Factor (F3)

= [(600-T70)/(T90-T70)](0.2)+0.7 =

[ (600-569)/ (618-569)] (0.2) +0.7 =

0.827 F3

MCB Factor (F4)

= [( 600-IBP )/( T10-IBP )]( 0.1 ) =

0.041 F4

Gasoline Factor (F5)

= Gasoline - cGasoline

[ ( 600 524 ( 17,680) -

=

16 Corrected LCO, BPSD

= [(LCO bpsd)F3 + (MCB bpsd)F4 + F5]/0.9

17 Corrected LCO °API 18 Corrected LCO, lb/hr

)/ (710-524)]

(0.1) =

(18,587) =

-907 F5

=

5,059

corrBPSD

16.6

corrLV%

= API + 5( LV%-corrLV% )/LV%

=

23.42

corr°API

0.913

corrSG

= corrBPSD x corrSG x 14.591

=

67,416

corrlb/hr

16.3

corrWt%

=

2,392

corrBPSD

7.8

corrLV%

=

35,780

corrlb/hr

8.7

corrWt%

MCB 19 Corrected MCB, BPSD

= Total C5+ Liq. Yield - Corr Gasoline - Corr LCO = = 17,689+6,515 +1,843-18,587 -5,059

20 Corrected MCB, lb/hr

= Total C5+ Wt. Yield - Corr Gasoline - Corr LCO = = 196,715+

87,450+

28,005-208, 974-67, 416

21 Corrected °API

= [( 141.5 * BPSD * 5.6146 * 62.3689 )/( lb/h * 24 )] - 131.5

=

6.5

corr°API

22

= [FF bpsd - corrMCB bpsd - corrLCO bpsd]100/[FF bpsd]

=

75.6

corrVol%

True Conversion

*Note that Total C5+ yield = cGasoline + LCO + MCB

API = (141.5/SG) - 131.5

Note- The Factor equations are valid only if the as produce 90% temperatures are between 360-400°F for gasoline and 580-620°F for LCO. The Factor Equations were developed from the 90% plus 10% Method. Used this method if the 90% Temp. are not in the specified range.

157048 Process Calculations Page 12

Sponge Gas Calculations Example Refiner: ______________________________Location: _____________________________ Date: _______________ A

B

C=AxB

D

E

g/100mol

Ib/hr

sp gr

0.04 0.00 1.47 4.31 5.82

1.3 0.0 64.7 120.7 186.7

14 0 716 1,335 2065

28.29 2.57 28.21 15.83 15.55 0.48 1.48 0.59 0.25 0.24 0.25 0.12 0.06 0.00 0.00 0.00 0.00 0.26 94.18 100.00

57.0 87.6 452.6 476.0 436.2 21.2 62.3 34.3 14.5 13.5 14.0 6.7 3.4 0.0 0.0 0.0 0.0 22.4 1,702 1,888

Component

MW

mole%

O2 CO CO2 N2 Total Inertes

32.00 28.01 44.01 28.01 (t)

H2 H2S C1 C2 C2= C3 C3= iC4 nC4 1-C4= i-C4= t-C4= c-C4= 1,3-C4== i-C5 n-C5 C5= C6+ Total Products TOTAL

2.02 34.08 16.04 30.07 28.05 44.10 42.08 58.12 58.12 56.11 56.11 56.11 56.11 54.09 72.15 72.15 70.14 86.18 (T)

MW = TC/TB = 18.9 SG = MW/28.966 = 0.6519 Mole % Inerts=tB= 5.82 Wt% Inerts=100tD/TD= 9.89 Inerts = tD = 2,065 Ib/hr

F bpsd (scfh)

(168) (0) (6,170) (18,091) (24,429)

631 969 5,006 5,265 4,825 234 689 379 161 149 155 74 37 0 0 0 0 248 18,821 20,887

0.5077 0.5220 0.5631 0.5844 0.6013 0.6004 0.6100 0.6271 0.6272 0.6248 0.6312 0.6496 0.6640

Pres. Psig = 173 Temp. F = 113 Meter Units = 6.5 Meter K = 91,100

(118,745) (10,787) (118,409) (66,445) (65,270) 31.6 90.4 46.2 18.8 17.0 17.7 8.4 4.1 0.0 0.0 0.0 0.0 25.6 259.8 (404,086)

=14.7=psia + 460=R

187.7 573.0

Gas with Inerts: V = Vol. Flow = KxUnitsxSQRT{psia/(R*SG)} = 419,742 scfh M = Mass Flow = scfh x MW/379.5 = 20,887 Ib/hr Inert-free Gas:

Mass = (M – Inerts) = M(100-wt%Inerts)/100 = Vol. Flow=(V)-(tF)=V-(V x M%Inerts/Tmoles) =

bpsd = [Ib/hr x24]/[ SGx5.614583ft3/bblx62.3689lb/ft3]

18,821 Ib/hr 395,313 scfh

scfh = V x Mol%/TB

D = C x M/TC = C x Total Mass Flow/(Total gr/100 mol)

157048 Process Calculations Page 13

Debutanizer Overhead (LPG) Calculation Example: Refiner: ______________________________Location: _____________________________ Date: _______________ A Component

MW

B mole%

C=AxE

D

g/100mol

SG

E=C/D

F=MxCftC G=VxE/tE

cc/lOOmol

lb/hr

H2S

34.08

0.0

0.0

0.7871

0

0

C2

30.07

0.0

0.0

0.3563

0

0

C2=

28.05

0.0

0.0

0.3680

0

0

C3

44.10

12.1

533.6

0.5077

1,051

5,216

704

C3=

42.08

36.2

1523.3

0.5220

2,918

14,891

1,955

iC4

58.12

10.7

621.9

0.5631

1,104

6,080

740

nC4

58.12

3.8

220.9

0.5844

378

2,159

253

1-C4=

56.11

8.5

476.9

0.6013

793

4,662

531

i-C4=

56.11

12.4

695.7

0.6004

1,159

6,801

776

t-C4=

56.11

9.2

516.2

0.6100

846

5,046

567

c-C4=

56.17

6.2

347.9

0.6271

555

3,401

372

1,3-C4==

54.09

0.3

16.2

0.6272

26

159

17

i-C5

72.15

0.4

28.9

0.6248

46

282

31

n-C5

72.15

0.2

14.4

0.6312

23

141

15

C5=

70.14

0.0

0.0

0.6496

0

0

0

C6+

86.18

0.0

0.0

0.6640

100.0

4,996

TOTAL (t)

0

0

8,899

48,838

bpsd

0 5,962

LPG S.G. =

tC/tE = 0.5614

Temperature, 0F

=

87

LPG MW =

tC/tB = 50.0

Vol. Corr. Factor

=

0.9653

Flow Gravity, Gf

=

0.5419

Meter Units Meter K

= =

7.2 631.5

V=Vol.Flow=Units x K x [SQRT(Gf)]/SG

=

5,962 bpsd

M = Mass Flow = (BPSD x SG x 5.6146ft3/bb1 x 62.3689lb/ft3)/24hr/d = =

48,838 lbs/hr

157048 Process Calculations Page 14

Propane and Butane Recovery Calculation Example: (1) (2) Sponge Gas Component C3 C3= Total C3’s (t)

bpsd 32 90 122

234 689 923

bpsd 704 1,956 2,660

lb/hr 5,215 14,888 20,103

(5) (6) Stabilized Gasoline (Debut. Bottoms) bpsd lb/hr 0 0 0 0 0 0

iC4 nC4 1-C4= i-C4= t-C4= c-C4= 1,3-C4= = Total C4’s (T)

46 19 17 18 8 4 0 112

379 161 149 155 74 37 0 956

740 253 531 776 567 372 17 3,257

6,078 2,159 4,661 6,800 5,045 3,400 159 28,301

0 47 34 34 112 44 109 380

lb/hr

(3) (4) Debutanizer Overhead

0 399 299 299 997 399 997 3,390

C3 Recovery = t(4) * 100 / [t(2) + t(4)] =

95.6

wt-%

C4 Recovery = [T(4) + T(6)] * 100 / [T(2) + T(4) + T(6)] =

97.1

wt-%

157048 Process Calculations Page 15

Yields Adjustment Composite 90% Plus 10% Method

This method uses the ASTM distillation of the liquid products to create a composite curve and correct the yields to any specified 90% temperatures. Procedure

1. Using the ASTM distillation and a straight line interpolation equation, calculate the LV% distilled every 20F for all the product streams. %x=[(Tx-Ta)/(Tb-Ta)](%b-%a)+%a

 Straight line interpolation equation

%a < %x < %b Ta < Tx < Tb 2. Calculate the composite volume and percent every 20F using the following equations: Cumulative BPD @ Tx = BPD = [(%Gasoline) (BPSD Gasoline) + (% LCO) (bpsd LCO) + (% MCB) (bpsd MCB)}/100 Cumulative LV% @ Tx = 100 BPD/Total BPD 3. Calculate Corrected Gasoline 90% @ 380F or Specified 90% Temperature: Corrected Gasoline (Composite Yield @ 380F)/(0.9) 4. Calculate Corrected LCO 90% @ 600F or Specified 90% Temperature: Corrected LCO = (Composite Yield @ 600F)-(Corr Gasoline)/(0.9) 5. Calculate Corrected MCB by difference: Corrected MCB = Total Liquid Yield – Correct Gasoline – Correct LCO

157048 Process Calculations Page 16

Data: (%) IBP 10% 30% 50% 70% 90% EP BPSD

Gasoline (°F) 94 124 165 222 294 367 410

LCO (°F) 319 460 503 533 569 618 653

MCB (°F) 524 710 759* 807* 876* 1100* 1200*

Total

17,680

6,515

2,227

26,422

Note - *These temperatures are not needed. %x = ((Tx-Ta)/(Tb-Ta))*(%b-%a)+%a %a < %x < %b Ta < Tx < Tb i,e, = ((100-94)/(124-94))*(10-0) +0 = 2.5 BPD = [(%Gasoline)(BPSD Gasoline) + (% LCO)(bpsd LCO)+(% MCB)(bpsd MCB)} / 100 %

= 100 BPD / Total BPD

157048 Process Calculations Page 17

Data: Tx, °F 80 100 120 140 160 180 200 220 240 260 280 300 320 340 360 367 380 400 410 420 440 460 480 500 520 540 560 580 600 618 620 640 660 680 700

Gasoline

LCO

MCB

%x 0.0 2.0 8.7 17.8 27.6 35.3 42.3 49.3 55.0 60.6 66.1 71.6 77.1 82.6 88.1 90.0 93.0 97.7 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0 100.0

%x

%x

0.0 0.1 1.5 2.9 3.4 4.3 5.7 6.5 7.2 8.6 10.0 19.3 28.6 37.9 53.9 65.0 74.5 82.7 90.0 90.6 96.3 100.0 100.0 100.0

0.0 0.9 1.9 3.0 4.1 5.1 5.2 6.2 7.3 8.4 9.5

Composite BPD* 0 354 1,532 3,148 4,873 6,235 7,475 8,716 9,724 10,706 11,688 12,667 13,640 14,701 15,726 16,134 16,728 17,643 18,100 18,147 18,239 18,332 18,938 19,544 20,150 21,210 21,958 22,600 23,156 23,656 23,696 24,092 24,358 24,382 24,406

Percent 0 1.34 5.80 11.91 18.44 23.60 28.29 32.99 36.80 40.52 44.24 47.94 51.62 55.64 59.66 61.06 63.31 66.77 68.51 68.68 69.03 69.38 71.67 73.97 76.26 80.27 83.10 85.54 87.64 89.53 89.68 91.18 92.19 92.28 92.37

Corrected Gasoline = =

(Composite Yield @ 380F) / (0.9) (16,728 / 0.9) = 18,587 BPD

Corrected LCO

= =

((Composite Yield @ 600F) - (Corrected Gasoline)) / (0.9) (23,156 – 18,587) / 0.9 = 5,077 BPD

Corrected MCB

= =

(Total Liquid Yield – Corrected Gasoline – Corrected LCO) (26,422 – 18,587 – 5,077) = 2,759 BPD

157048 Process Calculations Page 18

REACTOR AND REGENERATOR HEAT BALANCE

Burning coke in the regenerator provides all the heat necessary for the operation of the unit. Yet, roughly 30-40% of the heat generated by the combustion of coke exits the regenerator in the form of hot flue gas. The remainder is absorbed by the regenerated catalyst which carries it to the reactor riser where it is used to vaporize and heat up the combined feed to the desired cracking temperature. The amounts of energy associated with the unit's operation are determined from a catalyst section heat balance. The energy balance equation at steady state may be written as: Energy in + Energy produced = Energy out + Energy consumed

(1)

Regenerator Energy Balance

Energy in

= Energy (air + spent catalyst + coke)

Energy produced

= Combustion heat of coke

Energy out

= Energy (flue gas + regenerated catalyst + removed + radiation losses)

Energy consumed

= 0

If the Regenerator temperature is the reference temperature then, -∆H Air - ∆H Spent Catalyst - ∆H Coke + ∆H Combustion of coke = ∆Hremoved + ∆Hradiation losses

or:

∆H Spent Catalyst = ∆H Combustion of Coke - ∆H Coke - ∆H Air - ∆Hremoved ∆Hlosses (2)

157048 Process Calculations Page 19

Reactor Energy Balance

Energy in

= Energy (feed + regenerated catalyst + diluents)

Energy produced

= 0

Energy out

= Energy (reactor vapors + spent catalyst + radiation losses)

Energy consumed = Heat of reaction If the Reactor temperature is the reference base temperature, then -∆Hfeed - ∆Hdiluents + ∆Hregenerated catalyst = ∆Hradiation losses + Heat of Reaction or ∆Hregenerated catalyst = ∆Hfeed + ∆Hdiluents + ∆Hradiation losses + Heat of Reaction

(3)

The enthalpy change for the spent and regenerated catalyst is given by ∆Hspent catalyst = mass flow Cp (Rg Temp - Rx Temp)

(4)

∆Hregenerated catalyst = mass flow Cp (Rx Temp - Rg Temp)

(5)

At steady conditions, ∆Hspent catalyst + ∆Hregenerated catalyst = 0

(6)

157048 Process Calculations Page 20

Regenerator Heat Balance Flue Gas Spent Cataly st Coke Ra diati on Los ses H Combustion of Coke He at Remov al

Regenerated Catal ys t Air

Reactor Heat Balance Reactor Vapors Regenerated Catalyst

H Reaction

Spent Catalyst Coke Feed

Diluents

Radiation Losses

157048 Process Calculations Page 21

Combining equations (2), (3) and (6) ∆H Combustion of coke = ∆H Air + ∆H Coke + ∆Hremoved + ∆Hregen.rad. losses + ∆Hfeed + ∆Hdiluents + ∆Hrx rad. losses + Heat of Reaction

(7)

The equation (7) demonstrates that all the energy in the Reactor-Regenerator system is provided by the combustion of coke. The radiation loss term in this equation is not a major item, but since vessel insulation is not perfect, some radiation losses do occur. The term ∆Hremoved refers to the heat duty of catalyst cooler(s). The heat of reaction is the energy required to convert the feed to products via the catalytic reaction mechanism. The heat produced by the combustion of coke, equation (7), can be calculated from the coke product rate and the mode of the regeneration operation. If all the CO were burned to CO2 in the regenerator, more heat would be available per pound of carbon than when the unit runs in this normal partial combustion mode. The heat liberated by carbon combustion to CO2 is 14,150 BTU/lb (7,860 kcal/kg or 32,910 kj/kg) of carbon whereas heat for combustion to CO is only 3,960 BTU/lb (2,200 kcal/kg or 9,210 kj/kg). Conversion Factors: kcal/kg  1.8 = BTU/lb BTU/lb  2.326 = kJ/kg kcal/kg  4.187 = kJ/kg The heat of reaction is endothermic. Energy is consumed by the reaction which breaks the heavy hydrocarbon molecules into smaller, light hydrocarbon products. The heat of reaction must be calculated from the energy balance using equation (7). The most important value that can be calculated from the energy balance is the catalyst/oil weight ratio. This ratio is important because it is a major factor in hydrocarbon conversion and coke lay-down.

157048 Process Calculations Page 22

The following sample outlines the energy balance calculation method: 1.

Data Required

This heat balance is for a 30,565 BPSD feed case with the FCC Unit in total combustion mode. The process conditions are: Temperatures:

Pressures:

Reactor Combined Feed Lift Gas Lift Steam Feed Steam Stripping Steam

970°F 375°F 100°F 380°F 380°F 380°F

Regenerator Regenerated Catalyst Flue Gas Average Hottest in Rg Air Blower Discharge HP Boiler Feed Water Catalyst Cooler Steam

1371°F 1368°F 1375°F 399°F 220°F 463°F

Catalyst Cooler Steam

452 psig

157048 Process Calculations Page 23

Flow Rates:

*

Fresh Feed (No Recycle) Lift Gas Lift Steam Feed Steam Stripping Steam Total Air to Regenerator* Catalyst Cooler Steam Catalyst Cooler Blowdown

3,250 lb/hr 12,900 lb/hr 1,800 lb/hr 5,000 lb/hr 382,405 lb/hr 56,033 lb/hr 6,952 lb/hr

Total air includes: air to combustor, air to upper regenerator, and air to catalyst cooler.

Flue Gas Composition, mol% (by GC Method)

CO CO2 O2 + Ar N2 SO2 NO2

2.

412,923 lb/hr

= = = = = =

0.0 15.50 3.47 81.03 0.0 0.0

Flue Gas Composition Adjustment

Unlike an Orsat analysis, a GC analysis includes Argon with the Oxygen. The first step is to adjust the flue gas O2 content for Argon (Ar). The Ar content is assumed to be 1.2% of the nitrogen, therefore, Ar = (0.012) (81.03) = 0.97 mol%.

157048 Process Calculations Page 24

The corrected analysis is now: CO CO2 O2

= = = = = =

N2 + Ar SO2 NO2

3.

0 15.50 3.47 - 0.97 = 2.5

81.03 + 0.97 = 82.0 0 0

Combustion Air Corrected to a Dry Basis

A psychometric chart is used to determine the moisture content of the regeneration air. At atmospheric conditions of 62°F and a relative humidity of 97%, the moisture content is: Moisture Content =

0.01152 lb H2O lb dry air

Wet Air = 380,200 lb/hr Dry Air = 380, 200 lb/hr wet air 

1 lb dry air (1 + 0.01152) lb wet air

= 375, 870 lb/hr

Water in Air = 380,200 lb/hr - 375,870 lb/hr = 4,330 lb/hr

4.

Calculate Flue Gas Rate

The flue gas rate can be calculated from the regeneration air rate. These two streams are related by the inert N2 + Ar content which remains constant through the catalyst regeneration. From a Nitrogen balance, Since, moles =

Weight Molecular Weight

157048 Process Calculations Page 25

Then Dry Air = (375,870 lb/hr)/(28.966 MW) = 12,976 lb mol/hr mol/hr (N2 + Ar) in dry air = mol/hr (N2 + Ar) in flue gas 79 mol inerts lb mol FG 82 mol inerts 12, 976 lb mol  =  100 mol air hr 100 mol FG hr

Flue Gas (FG) = 12,501 lb mol/hr

5.

Calculate the Carbon (C) Content of Coke

The carbon (C) content of the coke is calculated from the flue gas composition. One mol of C is burned for each mole of CO or CO2produced. C + O2 + H2 + S + N = CO + CO2 + H2O + SO2 + NO2 + O2 C=

0 mol CO + 15.503 mol CO2 1 mol C 12, 501 lb mol   mol CO/CO2 hr FG 100 mol FG

C = 1,938 lb mol/hr of carbon

6.

Calculate the Hydrogen Content of Coke

The hydrogen (H2) content of the coke must be calculated from an O2 balance: O2 in regeneration air = excess O2 in flue gas + + O2 reacted to CO (0.5 mol O2/mol CO) + O2 reacted to CO2 (1 mol O2/mol CO2) + O2 reacted to H2O (0.5 mol O2/mol H2O) + O2 reacted to SO2 (1 mol O2/mol SO2) + O2 reacted to NO2 (1 mol O2/mol NO2)

157048 Process Calculations Page 26

Where: O2 in regen. air =

12, 976 lb mol dry air 21 mol O 2 2,725 lb mol  = of O2 hr hr 100 mol air

Excess O2 in FG =

O2 reacted to CO =

12, 501 lb mol FG 2.5 mol O2 312 lb mol  = of O2 hr hr 100 mol FG 12, 501 lb mol FG 0 mol CO 0.5 mol O2   hr 100 mol FG mol CO

= 0 lb mol/hr O2 O2 reacted to CO2 =

12, 501 lb mol FG 15.5 mol CO 2 1 mol O2   hr 100 mol FG mol CO2

= 1,938 lb mol/hr of O2 O2 reacted to SO2 =

12, 501 lb mol FG 0 mol SO2 1 mol O2   hr 100 mol FG mol SO2

= 0 lb mol/hr of O2 O2 reacted to NO2 =

12, 501 lb mol FG 0 mol NO2 1 mol O2   hr 100 mol FG mol NO 2

= 0 lb mol/hr of O2 O2 reacted to H2O (by difference) is: O2 Reacted to H2O = 2,725 - 312 - 0 - 1,938 - 0 - 0 lb mol/hr O2 = 475 lb mol/hr of O2 The hydrogen burned by oxygen in the regenerator is: H2 burned by O2 =

475 lb mol 2 mol H2 950 lb mol  = hr O2 mol O2 hr H2

157048 Process Calculations Page 27

7.

Calculate Coke from Carbon and Hydrogen

The mass of coke combusted to CO + CO2 + H2O is: from carbon =

1, 938 lb mol 12.01 lb C 23, 275 lb  = hr C lb mol C hr C

from hydrogen =

950 lb mol 2.016 lb H 1, 915 lb  = hr H2 lb mol H2 hr H

Total = 23,275 + 1,915 = 25,190 lb/hr coke

8.

Calculate Coke Yield Percent

The quantity of coke produced from the fresh feed is:

Coke Yield =

Coke, lb/hr100 FF,lb/hr 

Coke Yield =

25, 190 lb/hr coke 412, 923 lb/hr raw oil

9.

 100 = 6.10 wt - % coke

Calculate Hydrogen in Coke

The H2 content of the coke is: H2 in Coke =

H2 in Coke 

H2 ,

lb/hr(100) Coke, lb/hr

1,915 lb/hr H 25,190 lb/hr coke

 100 = 7.6 wt - % hydrogen

157048 Process Calculations Page 28

10. Calculate the Air/Coke Ratio

Air to Coke =

(Air, lb/hr)(100) Coke, lb/hr

Air to Coke =

375,870 lb/hr dry air lb of dry air = 14.92 25, 190 lb/hr coke lb of coke

11. Calculate the Heat of Combustion of Coke

Combustion heats are calculated based on the average hottest temperature in the regenerator. The dense, dilute, cyclones, and flue gas average temperatures are calculated and the hottest is used as basis. The average hottest temperature is 1375°F.

Hc (2C + O2

2CO) = 46,216 + 1.47 (1375°F) = 48,237

= (0

BTU lb  mole

BTU  lb  mole  of O2 reacted to CO) (2) 48, 237  lb  mole hr

= 0 BTU/hr

Hc (C + O2

CO2) = 169,135 + 0.5 (1375°F) = 169,822

= (1,938

BTU lb  mole

BTU  lb  mole  of O2 reacted to CO2) (1) 169, 822  lb  mole  hr

= 329.12  106 BTU/hr

157048 Process Calculations Page 29

Hc (2H2 + O2

2H2O) = 104,546 + 1.585 (1375°F) = 106,725

= (475

BTU lb  mole

BTU  lb  mole  of O2 reacted to H2O) (2) 106, 725  lb  mole  hr

= 101.389  106 BTU/hr ∆HCombustion of Coke = (0 + 329.12 + 101.39)  106 = 430.51  106 BTU/hr Using as basis 1 lb of coke ∆H combustion of coke =

430.51  10 6 BTU/hr 25, 190 lb/hr coke

= 17,090 BTU/lb coke This heat of combustion must be corrected for the coke’s hydrogen content according to the equation Correction

= 1133 - 134.6 (wt-% H) = 1133 - 134.6 (7.6) = +110 BTU/lb coke

The net heat of combustion of coke is: ∆HCombustion = 17,090 + 110 BTU/lb coke = 17,200 BTU/lb coke

157048 Process Calculations Page 30

REGENERATOR HEAT BALANCE Basis: 1 lb of coke 12. Heat Consumed to Heat Up the Regeneration Air

Since ∆H = mass  Cp  ∆T Air is heated from the main air blower discharge temperature of 399°F to the average hottest temperature of 1375°F at an average specific heat of 0.26 BTU/lb °F. 375, 870 lb/hr air 0.26 BTU HAir =  (1375- 399F)  = 3, 787 BTU/lb coke 25,190 lb/hr coke lb F

13. Heat Consumed to Heat up the Regeneration Air Water Vapor Water vapor is heated from 399 to 1375°F at an average specific heat of 0.5 BTU/lb °F. 0.5 BTU 4, 330 lb/hr H2 O HH O    (1375 - 399F) = 83.9 BTU/lb coke 2 lb F 25, 190 lb/hr coke

14. Heat Consumed to Heat Up the Coke Coke is heated from the reactor temperature of 970°F to the average hottest temperature of 1375°F at an average specific heat of 0.4 BTU/lb °F. ∆HCoke = (1375-970°F)  0.4 BTU/lb °F = 162 BTU/lb coke

157048 Process Calculations Page 31

15. Heat Consumed to Generate Steam in the Catalyst Cooler Enthalpies for water and steam are obtained from steam tables. Water in at 220°F Steam out at 463°F

= (188 BTU/lb) (56,033 + 6,952 lb/hr) = 11.841  106 BTU/hr = (1,205 BTU/lb) (56,033 lb/hr) = 67.52  106 BTU/hr

Blowdown out at 463°F = (441 BTU/lb) (6,952 lb/hr) = 3.066  106 BTU/hr Catalyst cooler duty = (67.52 + 3.066 - 11.841)  106 BTU/hr = 58.75  106 BTU/hr Or

∆HRemoved =

58.75  106 BTU/hr = 2, 332 BTU/lb of coke 25,190 lb/hr of Coke

16. Regenerator Heat Balance Using a typical regenerator heat loss rate of 250 BTU/lb coke, the heat consumed to heat up the catalyst is: RgHeat = (∆HComb Coke) - ∆HCoke - ∆HAir - ∆HH2O - ∆HLoss - ∆HRemoved Hregen = 17,200 - 3,787 - 84 - 162 - 2,332 - 250 = 10,585 BTU/lb Coke

157048 Process Calculations Page 32

17. Calculate the Catalyst Circulation Rate The catalyst is heated from the reactor temperature of 970°F to the regenerated catalyst temperature of 1371°F at an average specific heat of 0.275 BTU/lb °F. Since Q = m Cp ∆T then m = Q/Cp ∆T (Coke lb/hr)(Rg Heat BTU/lb Coke) (0.275 BTU/lb F) (Cat T - RXT)

CCR =

CCR =

25,190 lb/hr coke  10,585 BTU/lb coke  40, 299 lb/min 0.275 BTU/lb F  (1371 - 970F)  60 min/hr

or CCR =

40, 299 lb/min = 20.15 ton/min 2,000 lb/ton

18. Calculate the Catalyst/Oil Ratio C/O =

CCR lb/hr FF lb/hr

C/O =

40, 299 lb/min catalyst  60 min/hr = 5.86 wt/wt 412,923 lb/hr fresh feed

19. Calculate the Regenerator Efficiency Rg Eff =

Rg Heat  100 HCombustion of Coke

=

(10, 585 BTU/lb of Coke) (100) 17, 200 BTU/lb of Coke

= 62% (This number will be higher without catalyst cooler)

157048 Process Calculations Page 33

20. Calculate the Delta Coke Wt-%: Coke

(100) (Coke, lb/hr) 25, 190  100 = = 1.04 wt% Cat. Circ. lb/hr 40, 299  60

REACTOR HEAT CALCULATIONS 1 lb of fresh feed is used as basis in the following calculations.

21. Heat Consumed to Heat and Vaporize the Combined Feed Enthalpies for the raw oil feed are obtained by using the equation as discussed after this section. The UOP K to use for entering the enthalpy table is calculated from UOP Method 375 as discussed in the following section. UOP K Factor is a function of °API and Engler distillation. A high UOP K value of 12.5 indicates a more paraffinic (saturated chain) hydrocarbon, while a lower value of 11.2 occurs for a more aromatic (unsaturated cyclic) stock. Higher UOP K paraffinic feeds crack easier yielding higher conversion at a given reactor temperature. Raw Oil: UOP K = 11.8

°API Gravity = 21.3

At the 375°F riser inlet temperature,

Hraw oil

= 252 BTU/lb

At the 970°F reactor temperature,

Hvapor

= 760 BTU/lb

∆Hraw oil = 412,923 lb/hr  (760-2891 BTU/lb) = 194.487  106 BTU/hr The heat required to heat up the combined feed is the ∆Hraw oil multiplied by the Combined Feed Ratio (CFR), where, CFR =

(raw oil + recycle) lb/hr = 1.0 raw oil lb/hr

So, ∆Hcomb feed = 194.487  106 BTU/hr

157048 Process Calculations Page 34

Hcomb feed =

194.487  106 BTU/hr = 471 BTU/lb raw oil 412, 923 lb/hr raw oil

22. Heat Consumed to Heat Up the Lift Gas: The lift gas is heated from 110°F to 970°F at an assumed average specific heat of 0.5 BTU/lb °F. Hlift gas =

3, 250 lb/hr lift gas  (970 - 110F)  0.5 BTU/lb F = 3.4 BTU/lb raw oil 412, 923 lb/hr raw oil

23. Heat Consumed to Heat Up Lift Steam, Feed Steam and Stripping Steam Steam is heated up from the header temperature of 380°F to the reactor temperature at 970°F at an average specific heat of 0.485 BTU/lb °F.  Hsteam =

(5, 000 + 12, 900 + 1, 800) lb/hr  (970 - 380 F)  0.485 412, 923 lb/hr raw oil

= 13.8 BTU/lb raw oil

24. Heat of Inert Gas Carried from Regenerator to Reactor by Regenerated Catalyst The inerts gas can be calculated from the sponge gas stream and the procedure is at the end of this section. If this number is unknown, use 0.007% of fresh feed. Use an average specific heat of 0.275 BTU/lb °F. ∆Hinerts = (inerts wt%) (Cp) (RxT - RgT) = 0.007  0.275  (970 - 1371) = -0.8 BTU/lb raw oil

157048 Process Calculations Page 35

25. Reactor Heat Balance The total heat consumed in the reactor equals the sum of the heats consumed for the combined feed, lift gas, all steam, the reactor losses, plus the heat of reaction. Using a typical reactor heat loss rate of 2 BTU/lb raw oil, the heat balance is: Rx Heat = ∆Hcomb feed + ∆Hlift gas + ∆Hsteam + ∆Hinerts + ∆Hloss + ∆HRxN Hreactor = (471 + 3.4 + 13.8 - 0.8 + 2) BTU/lb raw oil + ∆Hreaction Hreactor = 489.4 BTU/lb raw oil + ∆Hreaction

26. Overall Heat Balance The heat consumed in the reactor is supplied by the hot catalyst circulated to the riser. At steady state, the heat consumed in the reactor must balance the heat produced in the regenerator. The reactor heat is based on a per lb of fresh feed basis while that of the regenerator is on a per lb of coke basis. These two can be equated using the raw oil to coke weight fraction to determine the heat of reaction: lb coke   Hregenerator BTU / lb coke  = Hreactor [BTU/lb raw oil] lb raw oil  

10, 585 BTU 25, 190 lb/hr coke  = 489.4 BTU/lb raw oil + ∆Hreaction lb coke 412, 923 lb/hr raw oil ∆Hreaction = 119 BTU/lb raw oil

157048 Process Calculations Page 36

FCC Unit Heat Balance Example Refiner: Location: Combustion Air Correction to a Dry Basis (a) Ambient Temperature = 62 °F (b) Relative Humidity = 97 % (c) Sat. Vapor Pressure = 10^[6.40375-(3165.36/(°F+392.565))] = 0.275586177 psia (d) lb H2O/lb dry air = [0.00622*(b)*(c)]/[14.7-0.01*(b)*(c)] = 0.011520532 lb/lb (e) Total Air to Regen = 380,200 lb/hr (f) Total Dry Air = (e)/[1+(d)] = 375,870lb/hr (h) H2O in Air = (e) – (f) = 4,330 lb/hr

Date: Flue Gas Composition Adjustment: GC, mol-% Corrected for Ar* (i) CO = 0.00 CO = 0.00 15.50 CO2 = 15.50 (j) CO2 = 3.47 O2 - (0.012*N2) = 2.50 (k) O2 = 81.03 N2 + O2 - Cor O2 = 82.00 (L) N2 = 0.00 SO2 = 0.00 (m) SO2 = 0.00 (n) NO2 = 0.00 NO2 = Total 100.00 100.00 *Correction required only for GC not for Orsat analysis.

Temperatures: Rg=Regenerator, Rx=Reactor (o) Flue Gas Line = 1368 °F (p) Rg Avg Cyclone Outlet = 1368 °F (q) Rg Avg Dilute = 1375 °F = 1371 °F (r) Rg Avg Dense (RgT) (s) Avg Hottest Rg Temp = 1375 °F Oxygen Balance: (y) (L) * (2 * 21 ) / 79 – 2(n) - 2(m) - 2(k) - 2(j) - (i) (z) 2.016*(y) =12.01[ (i) + (j) ] + 32.06*(m) + 46.01*(n) (1a) Hydrogen = 2.016 * (y) * 100 / (z) (1b) Air = 28.966 * 100 * (L) / [ 79 * (z) ] (1c) Coke = Dry Air / [ Air / Coke ] = (f) / (1b) (1d) Fresh Feed S.G. = 0.926 (1e) Coke Yield = (1c) * 100 / (1d)

(t) (u) (v) (w)

= = = = = = =

Air to Rg Rx Temp Hot Rg T – Air T Rg Dense T – Rx T

7.6 201.5 7.6 14.9 25,187 412,923 6.10

= = = =

399 970 976 401

°F °F °F °F

mol H2 / 100 mol Flue Gas lb coke / 100 mol Flue Gas wt-% Hydrogen in Coke Air/Coke, lb/lb lb/hr of Coke lb/hr Feed wt-% Coke (of Fresh Feed)

Combustion of Coke Basis: Average Hottest Regenerator Temperature: (1f) Hc(CO) = 46,216 + 1.47 * (s) = 48,237 BTU/lb-mol (1g) Hc(CO2) = 169,135 + 0.5 * (s) = 169,823 BTU/lb-mol (1h) Hc(H2O) = 104,546 + 1.585 * (s) = 106,725 BTU/lb-mol H Comb = [ (1f) * (i) + (1g) * (j) + (1h) * (y) ] / (z) = 17,091 BTU/lb-mol (1i) (1j) Correction Factor = 1133 – 134.64 * (1a) = 109 BTU/lb-mol H Combustion Coke = (1i) + (1j) = 17,200 BTU/lb-mol (1k) Regenerator Heat Balance Basis: 1 lb of Coke (1m) H coke = 0.4 BTU/lb-mol °F * (w) H air = (1b) * 0.26 BTU/lb °F * (v) (1n) H H2O = (h) / (1c) * 0.485 BTU/lb °F * (v) (1o) H Radiant Losses (1p) (1q) Cat Cooler Heat Duty H removed = Cooler Duty / Coke = (1q) * 10^6 / (1c) (1r) (1s) Rg Heat = (1k) – (1m) – (1n) – (1o) – (1p) – (1r) (1t) Rg Eff = Rg Heat * 100 / HcombCoke = (1s) * 100 / (1k) (1u) Cat/Oil = (1s) * (1e) / 100 * (0.275 BTU/lb °F) * (w) (1v) Cat Circ = (Cat/Oil) * (FF) / 60 = (1u) * (1d) / 60 min/hr) Coke = 100 * Coke / (60 * Cat Circ) = 100 * (1c) / 60 * (1v) (1w)

= = = = = = = = = = =

160 3,787 81 250 58.7 2,331 10,590 61.6 5.86 40,313 1.04

BTU/lb coke BTU/lb coke BTU/lb coke BTU/lb coke MM-BTU/hr (calculated separately) BTU/lb coke BTU/lb coke % Regenerator Efficiency Catalyst-to-Oil Ratio Catalyst Circulation, lb/min Coke, wt-%

157048 Process Calculations Page 37

FCC Unit Heat Balance Example Continuation Refiner: Location:

Date:

Reactor Heat Calculations: Basis: 1 lb of Coke (2a) Rx Temp (RxT) (2b) Combined Feed Temp. (CFT) (2c) Fresh Feed (FF) (2d) FF Enthalpy @ CFT (H @ CFT)* (2e) FF Enthalpy @ RxT (H @ RxT)I (2f) Recycle (Recy) (2g) Recy Enthalpy @ CFT (E @ CFT)* (2h) Recy Enthalpy @ RxT (E @ RxT)* *See Following Pages for Enthalpy Calculation Method

= = = = = = = =

970 375 412,923 289 760 0 0 0

°F °F lb/hr BTU/lb BTU/lb lb/hr BTU/lb BTU/lb

(2i) (2j) (2k) (2m) (2n) (2o) (2p) (2q) (2r) (2s)

Inerts from Rg with catalyst (Inerts) Inerts Specific Heat (Cp) Lift Gas (LGas) Lift Gas Temp (LGasT) Lift Gas Specific Heat (LGas Cp) Steam Temp (StmT) Steam Specific Heat (Stm Cp) Stripping Steam Lift Steam Feed Steam

= = = = = = = = = =

2,748 0.275 3,250 110 0.5 380 0.485 5,000 12,900 1,800

lb/hr BTU/lb °F lb/hr °F BTU/lb °F °F BTU/lb °F lb/hr lb/hr lb/hr

(2t) (2u) (2v) (2w) (2y) (2z)

Heat consumed by FF = (FF) * (H @ RxT – H @ CFT) / (FF) Heat consumed by Recycle = (Recycle) * (E @ CFT – E @ RxT) / FF Heat Consumed by Steam = (Total Stm) * (Stm Cp) * (RxT-StmT) / FF Heat Consumed by Lift Gas = (LGas) * (LGas Cp) * (RxT-StmT) / FF Heat from Inerts = (Inerts) * (Inerts Cp) * (RxT-RgT) / FF Reactor Heat Loss

= = = = = =

471 0 13.7 3.4 -0.7 2

BTU/lb FF BTU/lb FF BTU/lb FF BTU/lb FF BTU/lb FF BTU/lb FF

(3a)

Rx Heat = (H FF) + (H Recy) + (H Stm) + (H Lgas) + (H Inrts) + (H Loss) + (H Rxn) = 489.3

BTU/lb FF + H Reaction

(Rg Heat BTU/lb Coke) * (Coke lb/hr) / (FF lb/hr) = Rx Heat = BTU/lb + H Rxn (3b)

H Rxn = [(Rg Heat BTU/lb Coke) * (Coke lb/hr) / (FF lb/hr)] – BTU/lb =

157

H Rxn BTU/ lb FF

157048 Process Calculations Page 38

uop K Factor from °API and Engler Distillation Data:

ASTM Distillation: Vol% Temperature, °F 10 660 30 781 50 887 70 1015 90 1075

Specific Gravity:

0.9258

1. Calculate the volumetric average boiling point as the average of the 10, 30, 50, 70 and 90 vol-% temperatures. VABP = (T10% + T30% + T50% + T70% + T90%) / 5 VABP = (660 + 781 + 887 + 1015 + 1075) / 5 =

883.6

2. Calculate the Engler slope as °F per percent (°F/%) by subtracting the 10 vol-% temperature from the 90 vol-% temperature, and dividing the difference by 80. Slope = (T90% – T10%)/80

=

5.1875

3. Calculate the Cubic Average Boiling Point (CABP): CABP = VABP * A + B Where:

A = (0.000297 * Slope + 0.001438)*Slope + 1.0 A = 1.01545 B = (-0.581 * Slope – 1.339)*Slope B = -22.5809 4. Calculate UOP K: UOP K 

3

CABP  459.69 SG

= 11.89

157048 Process Calculations Page 39

Enthalpy of Heavy Petroleum Fractions The Following equations can be used to calculate the Enthalpies for the feed and recycle streams on the FCC unit. Liquid Enthalpy Equation:

Equation source is API Procedure 4.7.B4

A1 =

(-1171.26 + (23.722 + 24.907 * SG) * UOP K)

A1 =

A1 + (1149.82 – 46.535 * UOP K) / SG

A1=

A1 / 1,000

A2 =

(1 + 0.82463 * UOP K) * (56.086 – 13.817 / SG) / 1,000,000

A3 =

- (1 + 0.82463 * UOP K) * (9.6757 – 2.3653 / SG) / 1E+09

The enthalpy of liquid FCC feedstock at the riser inlet conditions is: Hin = A1 * (TE – 259.67) + A2 * (TE² – 259.67²) + A3 * (TE³ – 259.67³) Where:

TE

=

Combined Feed Temperature, (°F + 459.67)

SG

=

Fresh Feed Specific Gravity

UOP K

=

Fresh Feed UOP K

Vapor Enthalpy Equation: Equation source is a curve fit from UOP PD Chart PD-189 F1 =

3.0186E-04 * SG + 3.9975E-06 * UOP K * (UOP K – 13.8584)

F2 =

0.67036000 * SG + 0.00675130 * UOP K * (UOP K – 24.7770)

F3 =

85.52390000 * SG – 4.73260000 * UOP K * (UOP K – 21.9249) – 459.6742

Enthalpy of feed in fully vaporized condition is: Hout = F1 * (T²) – F2 * T + F3 Where: T

=

Reactor temperature, °F

157048 Process Calculations Page 40

Heat of Combustion of Coke, BTU/lb-Mole Table: Temperature, °F

77

1,100

1,200

1,250

1,300

1,350

1,400

CO

47,565

47,847

47,980

48,050

48,123

48,199

48,274

CO2

169,332

169,677

169,735

169,760

169,784

169,808

169,835

H2O Vapor

104,129

106,279

106,448

106,529

106,610

106,687

106,765

Equations: Hc(CO) = Hc(CO2) = Hc(H2O) = Where:

46,216 + 1.47 * (T) 169,135 + 0.5 * (T) 104,546 + 1.585 * (T)

T is in °F

157048 Process Calculations Page 41

MECHANICAL EVALUATION The Mechanical Evaluation Test should cover all facets of the FCC Unit, including the reactor, regenerator, main column, and the gas concentration unit. All major pieces of equipment should be part of this test, including vessels, pumps, compressors, heat exchangers, and piping hydraulics. The procedure for this test is long and involved, but the information can be very useful for the Refiner and for UOP. If the Refiner wants to revamp the unit, it is important to determine maximum throughput and actual equipment limitations. Most of the information will be collected only once, although parts of it (such as exchanger surveys) can be repeated to follow fouling or other potential problems. The lists and data sheets included in this section can be used as guidelines in collecting the required data, although the Refiner may have to modify certain parts for his particular unit. It is important to finish collecting the data within as short a time as possible. A single survey is generally satisfactory and it is no necessary to use long term average data. The Unit should be operating smoothly to get realistic and good quality data. Label the data collected and prepare a report in an orderly fashion.

157048 Process Calculations Page 42

GENERAL INFORMATION LIST This list is only a guideline. Please modify or expand as required. 1.

Ambient Air Conditions a. Temperature b. Relative Humidity c. Barometric Pressure d. Wind Velocity and Direction (show on rough plot plan)

2.

General Description of Unit a. Process Flow Diagram, including flow meter and control valve locations b. Plot Plan showing Layout of Vessel and Equipment c. Operational Mode (partial or complete CO Combustion)

3.

Units used (USA, Imperial, Metric) and Standard Conditions (0°C, 760 mm; 60°F, 14.7 psia)

4.

Limiting Factors a. Environmental Constraints (CO emissions, special specifications) b. Utility Limitation (shortage of steam or electricity, etc.)

5.

Performance Data a. Accurate Flow and Weight Balance including sample analyses

product

157048 Process Calculations Page 43

HYDRAULIC AND PROCESS SURVEY LIST 1.

Single gauge pressure survey of every point in reactor-regenerator circuit, including air into regenerator and flue gas out to point of discharge.

2.

Slide valve positions and hydraulic oil pressure at each valve.

3.

Air blower suction and discharge pressures, total and net (to regenerator) flow rates, relative humidity and temperature of air, with manufacturer’s data and performance curve for comparison.

4.

Electrical or steam consumption for blower driver.

5.

Power recovery units should add flue gas temperatures and pressures around expander, butterfly valve positions, electrical power consumed or generated and single gauge pressure survey of third stage separator.

6.

For electrostatic precipitator, or other flue gas treaters, take temperatures in and out, power consumption, and amount and size distribution of particulates removed.

7.

Complete flue gas sample before and after flue gas treater.

8.

Catalyst consumption and losses.

9.

Single gauge pressure survey of main column and gas concentration section (use one low pressure and one high pressure gauge, depending on location, for better accuracy). Include feed flow rate and temperature, reflux flow rate and temperature, reboiler heat input, overhead temperature and pressure and enough other data to calculate a heat and weight balance around the column.

10.

Samples of main column overhead receiver gas and liquid; and gas, hydrocarbons, and water samples from high pressure receiver for phase equilibrium studies.

157048 Process Calculations Page 44

11.

Compressor suction and discharge pressures, flow rates, composition, and temperatures, with manufacturer’s performance curve for comparison. Include amounts and compositions of liquids drained from knockout drums.

12.

Compressor driver type and power consumption.

13.

All pump suction and discharge pressures, flow rates, liquid compositions or boiling ranges, and power consumption of driver, with manufacturer’s performance data for comparison.

14.

Data on streams not usually measured, such as LCO to the sponge absorber, including flow rates, temperatures, composition or boiling range, and single gauge pressure survey of circuit.

15.

Pressures, temperatures, and flow rates of flushing oil to instruments and pump seals and glands.

16.

Utility consumption/product data: Steam (all pressures) Air (plant and instrument) Purges to instruments, packing glands, and expansion joints (specify air, steam, nitrogen, or fuel gas), with single gauge pressure survey of utility lines at purge points). Cooling water Boiler feed water Steam condensate Utility water Treating chemicals for boiler feed water (type and amount) Inhibitor and anti-corrosion chemicals used

157048 Process Calculations Page 45

EXCHANGERS INFORMATION LIST 1.

Flow through exchangers on both sides (gas and liquid), composition or boiling range, and mass flow.

2.

Temperatures in and out of both sides, also between shells and bundles.

3.

Pressures in and out of both sides, also between shells and bundles.

4.

Any material bypassed around exchangers (give rough sketch).

5.

If air coolers: air temperatures in and out, air velocity out, motor amps, note any belt slippage, variable pitch position, louver position, etc.

6.

In preparing data, submit overall heat transfer coefficient and specifics on exchangers.

157048 Process Calculations Page 46

FIRED HEATERS INFORMATION LIST 1.

Process flow (volume and mass, composition, molecular weight and boiling range).

2.

Single gauge pressure survey for both process and fuel system.

3.

Fuel type (gas or oil) and analysis (composition, sulfur, gravity, etc.), pressure and temperature of fuel at heater.

4.

Fuel consumption.

5.

Steam or air pressure for fuel oil atomization.

6.

Temperatures throughout the heater, such as firebox, convection points, stack, air preheat, and all process points.

7.

Draft in firebox and stack.

8.

Design information: type of furnace, materials of construction, and number, layout and materials of tubes; including dimensions of furnace.

9.

Burner data: rating, design, number. Note any unusual problems such as plugged or inoperative burners.

10.

Refiner should obtain sufficient data to calculate heat flux from both process and fire side, heat release, heater efficiency and steam balance.

157048 Process Calculations Page 47

157048 Process Calculations Page 48

Reactor-Regenerator Pressure Survey Refiner: Location: To Main Column

Date:

J

Time: By:

I 19 18 17 16 15 14 12

13

Steam Generator

Orifice Chamber

Flue Gas SV

11

20

H

10 ESP

Flue Gas to Stack

9

G 8

Slide Valves % Open Regenerated Recirculating Spent Flue Gas A Flue Gas B

6

7

F

Cat Cooler Slidevalve

E

22

Feed

5 D

4

Process Flow

C

DFAH

Feed Rate Air Rate

Main Air Blower

2

A

B

3

1 Atmospheric Air

Pressure Survey: _____________Units of Pressure 1 2 3 4 5 6 7 8

11 12 13 14 15 16 17 18

A B C D E F G H

157048 Process Calculations Page 49

157048 Process Calculations Page 50

157048 Process Calculations Page 51

MAIN COLUMN SUMMARY – BOTTOMS

page ____________________________ date ____________________________ Item No.: _____________________________ by ____________________________ Service: __________________________________________________________________ Type of Operation: _________________________________________________________ No. of Trays: _________________ Reflux Ratio:_________________________________ Type of Trays: _____________________________________________________________ Main Column Bottoms

Circulating HCO LCO

Circ.

Quench

CSO

Other

Mass Flow

_______

_______

_______

_______

_______ ________

Temperature Out Return Pressure

_______ _______ _______

_______ _______ _______

_______ _______ _______

_______ _______ _______

_______ ________ _______ ________ _______ ________

Distillation IBP 5% 10% 20% 30% 40% 50% 60% 70% 80% 90% 95% EP API or S.G.

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

BS & W

_______

_______

_______

_______

_______ ________

Steam to Stripper

_______

_______

_______

_______

_______ ________

________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________

(Sketch system showing flows, P, T, Q on separate page) __________________________ Weight balance _______________________ Heat balance _________________________ Deviations from UOP Specifications: ___________________________________________ ________________________________________________________________________

157048 Process Calculations Page 52

MAIN COLUMN SUMMARY — CYCLE OIL PRODUCTS AND OVHD.

page _______________________________ date _______________________________ Item No.: ____________________________ by _______________________________ Service: _________________________________________________________________ Type of Operation: _________________________________________________________ No. of Trays: ________________ Reflux Ratio: ________________________________ Type of Trays: ____________________________________________________________ HCO Product

LCO Product

Naphtha Product

Reflux

Net Ovhd. Liquid

Ovhd. Gas

Mass Flow,

_______

_______ ________ _______

_______

_______

Temperature

_______

_______ ________ _______

_______

_______

Pressure Composition, ______ % H2 N2 H2S H2O C1 C2 C3/C3= iC4 nC4/C4= iC5 nC5 C6+ Avg. Mol. Wt. Gravity Distillation

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

Steam to Stripper

_______

_______ ________ _______

_______

_______

Flash Point

_______

_______ ________ _______

_______

_______

IBP 5% 10% 20% 30% 40% 50% 60% 70% 80% 90% 95% EP

157048 Process Calculations Page 53

COLUMN SUMMARY

page _______________________________ date _______________________________ Item No.: _____________________________ by _______________________________ Service: __________________________________________________________________ Type of Operation: _________________________________________________________ No. of Trays: _________________ Reflux Ratio:_________________________________ Type of Trays: _____________________________________________________________ Net Off Ovhd. Feed Reflux Gas Btms. Liquid Other Mass Flow

_______

_______

_______

_______

_______ ________

Temperature Pressure

_______ _______

_______ _______

_______ _______

_______ _______

_______ ________ _______ ________

Composition, ______ % _______ _______ H2 N2 _______ H2S _______ H2O _______ C1 _______ _______ C2 C3 _______ iC4 _______ _______ nC4 iC5 _______ _______ nC5 C6+ _______ Avg. Mol. Wt. _______ Gravity _______ Distillation _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________

IBP 5% 10% 20% 30% 40% 50% 60% 70% 80% 90% 95% EP

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

(Sketch system showing flows, P, T, Q on separate page) __________________________ Weight balance _______________________ Heat balance _________________________ Deviations from UOP Specifications: ___________________________________________

157048 Process Calculations Page 54

ABSORBER SUMMARY

page _______________________________ date _______________________________ Item No.: ____________________________ by _______________________________ Service: _________________________________________________________________ Type of Operation: _________________________________________________________ No. of Trays: ________________ Reflux Ratio: ________________________________ Type of Trays: ____________________________________________________________ Gas In

Liquid In

Gas Out

Liquid Out

Pumparound Upper Lower

Mass Fl Temperature

_______ _______

_______ ________ _______ _______ ________ _______

_______ _______

_______ _______

Pressure

_______

_______ ________ _______

_______

_______

Composition, ______ % H2 N2 H2S H2O C1 C2 C3 iC4 nC4 iC5 nC5 C6+ Avg. Mol. Wt. Gravity Distillation, ° _______ IBP 5% 10% 20% 30% 40% 50% 60% 70% 80% 90% 95% EP

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________ ________

_______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______ _______

(Sketch system showing flows, P, T, Q on separate page) __________________________ Weight balance ______________________ Heat balance ________________________ Deviations from UOP Specifications: ___________________________________________

157048 Process Calculations Page 55

CENTRIFUGAL COMPRESSOR DATA page _______________________________ date _______________________________ Item No.: _____________________________ by _______________________________ Service: __________________________________________________________________ Manufacturer: _____________________________________________________________ Type, Model: ______________________________________________________________ No. of Stages: _____________________________________________________________ OPERATING CONDITIONS/PERFORMANCE Flow Rate: ____________ Suction Pressure: ____________ psig Discharge Pressure: ____________ psig Differential Head: ____________ Polytropic : ____________ Operating Speed: ____________ rpm

Suction Temperature: Discharge Temperature: Power: MW:

________ ________ ________ ________

°F °F hp

Type of Seal: _________________________________________ Lube/Seal Oil System: ________________________________________ Buffer Gas: (yes/no) Buffer Gas Rate: ______________ SCFH Automatic Surge Control: (yes/no) DRIVER Motor Manufacturer: _________________________________________ Rating: ____________________ Service Factor: ______________ Insulation Class: ____________________ Voltage/phase/cycle: Turbine Manufacturer:_________________________________________ Speed: ______________ Steam Supply: _______ psig Steam Rate: ______________ Steam Exhaust: _______ psig Gear Manufacturer: Rating: Type:

______ °F ______ °F

_________________________________________ ____________________ Service Factor: _____________ ____________________ Power Loss: _____________

Deviations from UOP Specification: ____________________________________________ ________________________________________________________________________ ________________________________________________________________________ ________________________________________________________________________

157048 Process Calculations Page 56

RECIPROCATING COMPRESSOR DATA page ___________________________ date ___________________________ by ___________________________

Item No.: ______________________________ Service: _______________________________ Manufacturer: ___________________________ Cylinder Lubrication: ____________ Type, Model: ___________________________ Clearance Pockets: (yes/no) No. of Stages, No. of Cylinders: ___________ Sparing Description: ____________ OPERATING CONDITIONS/PERFORMANCE Flow Rate: ____________ Suction Temperature: _________ °F Suction Pressure: ____________ psig Discharge Temperature: _________ °F Discharge Pressure: ____________ psig HP/stage: _________ hp MW: ____________

Operating Speed: ____________ rpm Cylinder Diameters: _________ Piston Speed: ____________ ft/s # of Suction/Discharge Valves: _________ Actual Rod Loadings, T/C: ________________________________________ lbf Max Allowable Rod Loadings, T/C: ________________________________________ lbf DRIVER Motor Manufacturer: Rating: Insulation Class:

________________________________________ ____________________ Service Factor: _____________ ____________________ Voltage/phase/cycle:

Turbine Manufacturer: ________________________________________ Speed: _______________ Steam Supply: _______ psig Steam Rate: _______________ Steam Exhaust: _______ psig Gear Manufacturer: Rating: Type:

______ °F ______ °F

________________________________________ ____________________ Service Factor: _____________ ____________________ Power Loss: _____________

Deviations from UOP Specification: ___________________________________________ ________________________________________________________________________ ________________________________________________________________________ ________________________________________________________________________

157048 Process Calculations Page 57

CONTROL VALVE SUMMARY page ___________________________ date ___________________________ Item No.: ______________________________

by ___________________________

Service: __________________________________________________________________ Description of Valve: _____________________

Design CV: ______________________

Mfgr. and Catalog No.: ______________________________________________________ Positioner? _______________________________________________________________

Actual

Design

Percent open (valve position)

____________

Flow rate:

______________________

____________

__________

Upstream pressure:

______________________

____________

__________

Downstream pressure: ______________________

____________

__________

Flowing temperature:

____________

__________

______________________

Deviations from UOP Specification: ____________________________________________ ________________________________________________________________________ ________________________________________________________________________ ________________________________________________________________________

157048 Process Calculations Page 58

AIR FIN COOLER SURVEY

page _______________________________ date _______________________________ Item No.: ____________________________ by _______________________________ Service: _________________________________________________________________ Manufacturer: _____________________________________________________________ Type, Model: _____________________________________________________________ No. of Bundles: _______________________ No. of Passes: _____________________ No. of Tubes per Pass: _________________ Fans/bundle: ______________________ Tube Size _______________ ID x _______________ Gauge x _____________ Length Piping Geometry: ______________________ Type*: ____________________________ Overall Heat Transfer Coefficient: _____________________________________________ Inlet Outlet Air

In Out No. fans on__________________________ Louver position_______________________ Mass flow Q (calc.) Composition, ____ % H2 N2 H2S H2O C1 C2 C3 iC4 nC4 iC5 nC5 C6+ Avg. Mol. Wt. Relative Humidity

Pressure ______________ ______________

Temperature _____________ _____________

______________ _____________ ______________ _____________ Pitch control ________________________ Air ______________ ______________

______________

Process _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________ _____________

157048 Process Calculations Page 59

Gravity Distillation, ° ______ IBP 10% 30% 50% 70% 90% EP

_____________ _____________ _____________ _____________ _____________ _____________ _____________

Deviations from UOP Specification: ____________________________________________ ________________________________________________________________________ ________________________________________________________________________ *Include sketch of piping geometry if different from UOP standard practice types.

157048 Process Calculations Page 60

FLOW METER SUMMARY page ___________________________ date ___________________________ Item No.: ______________________________

by ___________________________

Service: _________________________________________________________________ Type of Fluid: ___________________________

Normal Units of Flow: ______________

______________________________________ Type of Meter: ____________________________________________________________

Meter Reading: ___________________________________________________________ Pressure

________________

Temperature

________________

Sp. Gr.**

________________

Meter Factor

________________

Corrected Flow Rate

________________

Mass Flow Rate

________________

Avg. mol. wt.

________________

Molar Flow Rate

________________

**Sketch piping layout, showing distances in nominal pipe IDs.

157048 Process Calculations Page 61

HEAT EXCHANGER SURVEY

page _______________________________ date _______________________________ Item No.: _____________________________ by _______________________________ Service: __________________________________________________________________ Manufacturer: _____________________________________________________________ Type, Model: ______________________________________________________________ No. of Bundles: ____________________________________________________________ No. of Passes/Bundle: __________________ Tubes per Pass: ____________________ Tube Size ______________ ID x _______________ Gauge x ______________ Length Heat Exchange Surface Area/Bundle: __________________________________________ Piping Geometry (sketch if necessary): _________________________________________ Length of Service: __________________________________________________________ Design Heat Transfer Coefficient: ______________________________________________ Shell Side

Inlet

Stream A

Outlet Tube Side

Inlet Outlet

B

Q (calc.) Shell side Q (calc.) Tube side

______________ ______________

Composition, ______ % H2 N2 H2S H2O C1 C2 C3 iC4 nC4 iC5 nC5 C6+ Mass Flow Avg. Mol. Wt.

A ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________

Pressure ______________

Temperature _____________

______________

_____________

______________ ______________

_____________ _____________

B ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________ ______________

157048 Process Calculations Page 62

Gravity Distillation, ° _______ IBP 10% 30% 50% 70% 90% EP

______________ ______________ ______________ ______________ ______________ ______________ ______________

______________ ______________ ______________ ______________ ______________ ______________ ______________

Deviations from UOP Specification: ___________________________________________ ________________________________________________________________________ ________________________________________________________________________

157048 Process Calculations Page 63

HEATER SURVEY page _______________________________ date _______________________________ Item No.: _____________________________ by _______________________________ Service: __________________________________________________________________ Manufacturer: _____________________________________________________________ Type, Model: ______________________________________________________________ No. of Passes:_________________________ Tubes per Pass: ____________________ Tube Size ______________ ID x _______________ Wall x ________________ Length Geometry (Process): ________________________________________________________ Geometry (Flue Gas): _______________________________________________________

Radiant

Inlet

Stream A

Outlet

Pressure ______________

Temperature _____________

______________

_____________

Convection I

Inlet Outlet

B

______________ ______________

_____________ _____________

Convection II

Inlet Outlet

C

______________ ______________

_____________ _____________

Convection III

Inlet Outlet

D

______________ ______________

_____________ _____________

Fuel Gas

E

______________

_____________

Fuel Oil

F

______________

_____________

Flue Gas Under Convection I

G

______________

_____________

Flue Gas Under Convection II

H

______________

_____________

Flue Gas Under Convection III

I

______________

_____________

Flue Gas Under Stack Damper

J

______________

_____________

Flue Gas Above Floor

K

______________

_____________

157048 Process Calculations Page 64

HEATER SURVEY page _______________________________ date _______________________________ by _______________________________

Stream A B Mass Flow, ________ _____ _____ Composition, ______ % _____ _____ H2 _____ _____ _____ _____ N2 _____ _____ O2 CO _____ _____ _____ _____ CO2 _____ _____ H2S _____ _____ SO2 _____ _____ C1 _____ _____ C2 _____ _____ C3 _____ _____ iC4 _____ _____ nC4 _____ _____ iC5 _____ _____ nC5 _____ _____ C6-205°C (400°F) 205°C (400°F)+ _____ _____ Avg. Mol. Wt. _____ _____ Gravity _____ _____ Viscosity _____ _____ Total Sulfur, _______ _____ _____ Metals, ___________ _____ _____ Q (calc.) Absorbed _____ _____ Q (calc.) Released Heater Gross Efficiency Excess Air, % Tube Skin Temps:,° _____ Burner Pressure ______________________

C _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____

D _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ ____

E _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____

F _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____ _____

_____

_____

G,H, I,J,K ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______ ______

% of Rating ____________________

Provide sketch showing piping and controls for process piping. Deviations from UOP Specification: ___________________________________________ ________________________________________________________________________ ________________________________________________________________________

157048 Process Calculations Page 65

CENTRIFUGAL PUMP SURVEY page _______________________________ date _______________________________ Item No.: _____________________________ by _______________________________ Service: __________________________________________________________________ Manufacturer: _____________________________________________________________ Type, Model: ______________________________________________________________ No., Size and Style (Mfgrs. Designation) ________________________________________ ________________________________________________________________________ Suction

Pressure ______________

Discharge

______________

Temperature _____________

Other Information Rated Flow (STP) _____________ Seal Type? Single, Tandem, Double, Bellow Sp. Gr. _____________ Spillback? Yes/No Viscosity _____________ NPSHR? _________________________ Static Suction Head _____________ Suction Specific Speed: ________________ Speed _____________ Differential Head (flowing condition) _________________________________________ Driver Type: ___________________________________________________________ Manufacturer: ___________________________________________________________ No., Size, Rating and Style (Mfgrs. designation): __________________________________ Rating: _________________ Insulation Class: _________________ Service Factor: _________________ Voltage/Phase/Cycle: _________________ Motor: Power consumption Speed Turbine: Steam consumption Steam supply

______________ ______________ ______________

Steam exhaust Speed

Pressure ______________

Temperature ______________

______________

______________

______________

Supply copy of Mfgrs. pump curve and plot operating point. Deviations from UOP Specification: ____________________________________________ ________________________________________________________________________ ________________________________________________________________________

157048 Process Calculations Page 66

157048 Process Calculations Page 67

SUPPLEMENTAL CALCULATIONS REGENERATOR VELOCITIES The following procedure shows how to calculate the superficial velocity in the combustor, upper regenerator, and cyclones. This section presents two methods to calculate the velocities in the regenerator. The first method is more precise but requires more information and it is more laborious than the second method.

Method A 1.

Required Information

The FCC Unit is in total combustion mode for this case and with no catalyst cooler. The process conditions are: Temperatures:

Average Dense Average Dilute Average Cyclones Air to Regenerator Ambient Relative Humidity

1371°F 1375°F 1375°F 399°F 62°F 97%

Pressures:

Regenerator Combustor Cyclones

32 psig 34 psig 31 psig

Areas:

Combustor Cross sectional Regenerator Cross sectional First Stage Cyclones Second Stage Cyclones

300 ft2 452 ft2 24.5 ft2 21.3 ft2

157048 Process Calculations Page 68

Flow Rates:

Flue Gas:

Air to Regenerator 83,615 scfm = (scfm x 28.76 lb/mol x 60 min/hr)/(379.5 scf/mol) = 380,200 lb/hr CO CO2 O2 N2 + Ar SO2 NO2

2.

= = = = = =

0 15.50 2.5 82.0 0 0

Combustion Air Correction to a Dry Basis

A psychometric chart is used to determine the moisture content of the air. At atmospheric conditions of 62°F and a relative humidity of 97%, the moisture content is: Moisture Content =

0.01152 lb H2O lb dry air

Wet Air = 380,200 lb/hr Dry Air = 380, 200 lb/hr wet air 

1 lb dry air (1 + 0.01152) lb wet air

Water in Air = 380,200 lb/hr - 375,870 lb/hr = 4,330 lb/hr

= 375, 870 lb/hr

157048 Process Calculations Page 69

3.

Calculate Flue Gas Rate

The flue gas rate can be calculated from the regenerator air rate. These two streams are related by the inert N2 + Ar content which remains constant through the catalyst regeneration. Since, moles =

Weight Molecular Weight

then, Water in Air = (4,330 lb/hr)/(18 MW) = 241 mol/hr Dry Air = (375,870 lb/hr)/(28.966 lb/mol) = 12,976 lb mol/hr mol/hr (N2 + Ar) in dry air = mol/hr (N2 + Ar) in flue gas 79 mol inerts lb mol FG 82 mol inerts 12, 976 lb mol  =  100 mol air hr 100 mol FG hr

Flue Gas (FG) = 12,501 lb mol/hr

4.

Calculate the Water Produced by the Hydrogen Content of Coke

The overall reaction occurring in the regenerator is: C + H2 + S + N + O2 = CO + SO2 + NO2 + H2O + O2

157048 Process Calculations Page 70

The water produced by the hydrogen (H2) content of the coke can be calculated from an O2 balance: O2 in regeneration air = excess O2 in flue gas + + O2 reacted to CO (0.5 mol O2/mol CO) + O2 reacted to CO2 (1 mol O2/mol CO2) + O2 reacted to H2O (0.5 mol O2/mol H2O) + O2 reacted to SO2 (1 mol O2/mol SO2) + O2 reacted to NO2 1 mol O2/mol NO2 where: O2 in regen. air =

12, 976 lb mol dry air 21 mol O 2 2,725 lb mol  = of O2 hr hr 100 mol air

Excess O2 in FG =

O2 reacted to CO =

12, 501 lb mol FG 2.5 mol O2 312 lb mol  = of O2 hr hr 100 mol FG

12, 501 lb mol FG 0 mol CO 0.5 mol O2   = 0 lb mol/hr O2 hr 100 mol FG mol CO

O2 reacted to CO2 =

12, 501 lb mol FG 15.5 mol CO 2 1 mol O2   =1,938 lb mol/hr of O2 hr 100 mol FG mol CO2

O2 reacted to SO2 =

12, 501 lb mol FG 0 mol SO2 1 mol O2   = 0 lb mol/hr of O2 hr 100 mol FG mol SO2

O2 reacted to NO2 =

12, 501 lb mol FG 0 mol NO2 1 mol O2   = 0 lb mol/hr of O2 hr 100 mol FG mol NO 2

157048 Process Calculations Page 71

O2 reacted to H2O (by difference) is: O2 Reacted to H2O = 2,725 - 312 - 0 - 1,938 - 0 - 0 lb mol/hr O2 = 475 lb mol/hr of O2 Since H2 + 1/2O2 = H2O Then The water produced by Hydrogen and Oxygen in the regenerator is: H2 O Produced by O2 =

5.

475 lb mol 2 mol H2O 950 lb mol  = hr O2 mol O2 hr H2

Calculate the Wet Flue Gas Rate

The total moles per hour of wet flue gas are: Wet Flue Gas = Dry Flue Gas + Water from Air + Water from H2 in Coke = 12,502 mol/hr + 241 mol/hr + 949 mol/hr = 13,692 mol/hr The actual cubic feet per second (ACFS) of the flue gas can be calculated by using the Ideal Gas equation of state PV = nRT Where R Prg Pcomb Pcycl Tcomb Tdense Tdilute Tcycl

then

= = = = = = = =

V = nRT/P

10.7 (psia x ft3)/(mol x °R) 32 psig + 14.7 = 46.7 psia 34 psig + 14.7 = 48.7 psia 31 psig + 14.7 = 45.7 psia 1,275°F + 460 = 1,735 °R 1,371°F + 460 = 1,831 °R 1,375°F + 460 = 1,835 °R 1,368°F + 460 = 1,828 °R

157048 Process Calculations Page 72

ACFS @Tcomb = 13,692 mol/hr x 10.73 x 1,735/(48.7x3600 sec/hr) = 1,454 ft3/sec ACFS @Tdens

= 13,692 mol/hr x 10.73 x 1,831/(46.7x3600 sec/hr) = 1,596 ft3/sec

ACFS @Tdilute = 13,692 mol/hr x 10.73 x 1,835/(46.7x3600 sec/hr) = 1,599 ft3/sec ACFS @Tcyc

6.

= 13,692 mol/hr x 10.73 x 1,828/(45.7x3600 sec/hr) = 1,632 ft3/sec

Calculate Superficial Velocities

Regenerator Superficial Velocity = (ACFS @ Tdens, ft3/sec) / (Rg Cross Sect Area, ft2)

= 3.5 ft/sec

First stage Cyclones Superficial Velocity = (ACFS @ Tdilute ft3/sec) / (Total Inlet Area, ft2)

= 65.2 ft/sec

Second Stage Cyclones Superficial Velocity = (ACFS @ Tcyc, ft3/sec) / (Total Inlet Area, ft2)

= 76.6 ft/sec

Combustor Superficial Velocity = (ACFS @ Tcomb, ft3/sec) / (Comb Cross Sec Area, ft2)

= 4.8 ft/sec

157048 Process Calculations Page 73

Method B 1.

Required Information

This method does not require the flue gas analysis. The process conditions are: Temperatures: Pressures: Area: Flow Rates:

2.

Average Dense Regenerator Regenerator Cross sectional Air to Regenerator

1371°F 32 psig 452 ft2 83,615 scfm

Calculate the Actual Cubic Feet per Second

The volumetric flue gas rate can be calculated by using the Ideal Gas equation of state PV = nRT

then

R = PV/nT

For air we have

R = P1V1/n1T1

For the Regenerator Air

R = P2V2/n2T2

Combining the last two equations

P1V1/n1T1 = P2V2/n2T2

Or

V2 = P1V1T2 x n2 T1P2 n1

Where: P1 T1 V1 P2 T2

= = = = =

0 pisg + 14.7 60 °F + 460 83,615 scfm/(60 s/m) 32 pisg + 14.7 1,371°F + 460

= = = = =

14.7 psia 520 °R 1,393.6 ft/s 46.7 psia 1,831 °R

157048 Process Calculations Page 74

n2/n1

=

1.04 Assumed. This factor is due to the combustion of Hydrogen to in the coke to water.

Then V2

=

(14.7 psia)(1,393.5 ft/s)(1,831°R)(1.04) = 1,606 ft3/s (520°R)(46.7 psia)

3.

Calculate Regenerator Superficial Velocity

The superficial velocity is calculated by dividing the volumetric flow rate by the cross sectional area: 1,606 ft3/s = 3.6 ft/s 2 452 ft The molar expansion factor n2/n1 can be approximated if the flue analysis is available by using the following equation:

Rg Velocity

=

n2/n1

=

2 - (79/N2%)

Where N2% is the Nitrogen percent form the flue gas analysis.

157048 Process Calculations Page 75

REGENERATOR AIR DISTRIBUTOR PRESSURE DROP The pressure drop across the regenerator air distributor can be calculated by the following formula:

P =

q2 C2 A2 (2g) 144

where: ∆P q r C A g

= = = = = =

pressure drop, psi air rate at flowing conditions, ft3/sec density of air at flowing conditions, lb/ft3 orifice coefficient, 0.60-0.80 total cross sectional area of holes, ft2 acceleration due to gravity, 32.2 ft/sec2

The perforated grid air distributor is designed for a pressure differential of about 0.71.2 psi. This will give good air distribution for fluidization without causing catalyst attrition. If the pressure differential is too high, the high velocity can cause attrition. If the pressure differential is too low, less than 0.5 psi, it can cause poor distribution of air and distributor erosion problems.

157048 Process Calculations Page 76

REACTOR STRIPPER DENSITY The following procedure shows how to calculate the density in Reactor Stripper. The density indicator is a differential type instrument and for this case the range is 0-145 inches of water.

1.

Calculate the Distance Between Taps

Lower Instrument Tap Elevation: 59' 8 7/8" or 59.7396' Upper Instrument Tap Elevation: 74' 4 1/8" or 74.3438' Distance Between Taps: 74' 4 1/8" - 59' 8 7/8" = 14' 7 1/4" or 175.25"

2.

Calculate the Density Instrument Readout: 75% Instrument Span: 0-145 inches H2O 75% x 145 inches H2O = 108.75 inches H2O 108.75 in H 2 O 

1 12 in lb/in 2 144 in 2 lb    = 38.7 3 2 175.25 in ft 27.705 in H 2 O ft ft

157048 Process Calculations Page 77

Notes: i)

The Spent Catalyst Stripper Density is expected to range from 30 to 45 lb/ft3.

ii)

In cases when the pressure taps are under the stripper baffles the distance between taps should be replaced by the distance between the bottom edges of the baffles.

157048 Process Calculations Page 78

REACTOR STRIPPER LEVEL This procedure shows how to calculate the level in the reactor. The level controller is a differential type instrument and for this case the range is 0-300 inches of water.

1.

Calculate the Distance Between the Taps

Lower Instrument Tap Elevation: 62' 3 7/8" or 62.3229' Upper Instrument Tap Location: 127' 0" - 3' 3" - 1' 6" = 122' 3" or 122.2500' Distance Between Upper Tap and Lower Tap: 122' 6" - 62' 3 7/8" = 60' 2 1/8" or 60.1771'

2.

Data Required Instrument Readout: Instrument Span:

50% 0-300 inches H2O

Stripper Density of 36.25 lb/ft3 (From Stripper Density Calculation) Assume Reactor Vapor Space Density of 1 lb/ft3 Distance Between Upper Tap and Lower Tap = 60.1771' Elevation of Lower Tap = 62.3229' Normal Reactor Catalyst Level = 84.500' Cyclone Dipleg Outlet = 79.5417' 300 in H2O x 50% = 150 in H2O

157048 Process Calculations Page 79

3.

Method A - This is rough method. lb/in 2 144 in 2 ft 3 150 in H 2 O    = 21.51 ft of Catalyst 27.705 in H 2 O 36.25 lb ft 2

Distance Relative to Normal Catalyst Level = (62.32 + 21.51) - 84.50 = - 0.67 (i.e. ~ 8" below normal level)

4. Method B - This method is more precise than Method A since consider the reactor vapor density. X + Y = 60.18 ft => Y = 60.18 - X Where:

X = Catalyst height. y = Reactor vapor height from catalyst bed to upper pressure tap. 60.18 ft = distance between pressure taps.

lb lb/in 2 144 in 2 X (36.25 lb)   + (60.18 - X)  1 3 150 in H 2 O  2 3 27.705 in H 2 O ft ft ft

779.64 lb/ft2 = 35.25 X lb/ft3 + 60.18 lb/ft2 Then X = 20.41 ft above Lower Level Tap Distance Relative to Normal Catalyst Level = (62.32 + 20.41) - 84.50 = - 1.77 (i.e. 1' 9 1/4" below normal level)

157048 Process Calculations Page 80

Notes: i)

The Spent Catalyst Stripper Density is expected to range from 30 to 45 Lbs/Ft3.

ii)

In cases when the lower pressure tap is under the stripper baffle the distance between taps should be replaced by the distance between the bottom edge of the baffle and the upper tap.

157048 Process Calculations Page 81

REACTOR RISER RESIDENCE TIME Residence time, the time that hydrocarbons spend in the riser, is another design variable utilized in controlling reaction severity. This variable is of particular importance in operation using high activity zeolitic catalyst. Typical residence time for current designs is two to three seconds. Conversion is proportional to residence time in that it increases with prolonged contact of catalyst and feedstock. Gasoline yields increase with residence time up to a point after which over-cracking may occur. This results in a loss of gasoline yields and a significant increase in conversion. The method used to calculate riser residence time is as follows: 

= VR/[(1/3)(VF) + (2/3) (VP)]



= Residence time in seconds

VR = Riser volume, ft3 VF

= Volume of vaporized feed, steam, lift gas, water, and inerts calculated at the average conditions at the point of feed injection, ft3/s.

VP

= Volume of vaporized products, steam, lift gas, water and inerts calculated at the average conditions at the point of feed injection, ft3/s.

It should be noted that prior to the residence time calculation, the average temperature and pressure at the point of feed injection must be estimated to obtain VF and VP.

157048 Process Calculations Page 82

The riser pressure at the point of feed injection can be approximated by assuming a 5 psig pressure drop, hence: Riser Pressure = Reactor dome pressure + 5 psi The average temperature at the point of feed injection is calculated as:

Tavg 

0.2753C / O TR   CPo To   L / O CPL TL   0.495W / O TW   S / O TS   W / O HWL   H Rx   H OL  0.2753C / O   CPo   0.495W / O   S / O   L / O CPL 

Where: C/O

=

Catalyst to Oil wt. ratio, calculated in reactor-regenerator heat balance section

S/O

=

Steam to Oil wt. ratio

W/O

=

Water to Oil wt. ratio

L/O

=

Lift gas to Oil wt. ratio

TR

=

Regenerator dense bed temperature, °F

TO

=

Oil feed temperature, °F

TL

=

Lift gas feed temperature, °F

TS

=

Steam feed temperature, °F

TW

=

Water feed temperature, °F

CPO

=

Specific heat of vaporized oil feed, Btu/lb/°F

157048 Process Calculations Page 83

CPL

=

Specific heat of lift gas, Btu/lb/°F

∆HOL

=

Latent heat of vaporization of oil feed at inlet temperature, To, Btu/lb

∆HWL

=

Latent heat of vaporization of water at inlet temperature, Tw, Btu/lb

∆HRX

=

Heat of reaction, Btu/lb, calculated in reactor-regenerator heat balance section

Tavg

=

Average temperature at point of feed injection, °F

157048 Process Calculations Page 84

HYDROGEN BALANCE This document describes a manual method for calculating the Hydrogen balance for an FCC Unit.

Data required: A normalized to 100% recovery product summary in wt-% A breakdown of the C4- components The distillation and API of each C5+ product The distillation and API of the feed API Technical Data Book Figure 2B1.1 "Characterizing Boiling Points of Petroleum Fractions" Figure 2 - UOP Chart 409B-12 "Hydrogen Content of Liquid Petroleum Hydrocarbons"

Description of the Calculation Method: Step 1. Determine the molecular weight of Hydrogen per molecule for each C4product, e.g. H2S, H2, C1, C2, C2=, etc. See column 3 of the attached example. Step 2. Determine the percentage Hydrogen in each C4- product component by dividing the molecular weight of Hydrogen per molecule by the molecular weight of each component. In the attached example this is column 3 divided by column 4 times 100. The result is given in column 5.

Step 3. Calculate the Volume Average Boiling Point (VABP) for each of the heavier products (C5+ gasoline, LCO, and MCB). See Figure 2B1.1 comments for procedure and definitions.

Step 4. Calculate the Engler Slope for each of the heavier products (C5+ gasoline, LCO, and MCB). See Figure 2B1.1 comments for procedure and definitions.

157048 Process Calculations Page 85

Step 5. Determine the Mean Average Boiling Point from API Figure 2B1.1. Step 6. Determine the Hydrogen content of the hydrocarbon liquid from Figure 2. See the lower half of column 5 in the attached example. Step 7. Multiply the wt% normalized yield pattern for each component by the percentage Hydrogen in each product component. In the attached example this is column 2 times column 5. The result is given in column 6. Step 8. Calculate the percentage of Feed Hydrogen in each component by dividing the Wt% Hydrogen in each product component by the Wt% H2 in the feed. In the attached example this is the value in column 6 divided by the Wt% H2 in the feed (13%).

157048 Process Calculations Page 86

Hydrogen Balance Summary 1

2

3

4

5

6

7

Mass Balance

wt-%

#H/Molecule

MW

%H2

H2 wt-%

% Feed H2

Results Feed:

100.00

Values from Figures 2B1.1 & Fig. 2

13.00

100.00

H2S

0.04

2.0158

34.08

0.06

0.0021

0.02

H2

0.23

2.0158

2.02

1.00

0.2280

1.75

C1

0.89

4.0361

16.04

0.25

0.2245

1.73

C2

0.81

6.0474

30.07

0.20

0.1622

1.25

C2=

0.91

4.0316

28.05

0.14

0.1303

1.00

C3

1.30

8.0632

44.10

0.18

0.2375

1.83

C3=

4.64

6.0474

42.08

0.14

0.6662

5.12

IC4

2.69

10.0790

58.12

0.17

0.4665

3.59

NC4

0.63

10.0790

58.12

0.17

0.1094

0.84

C4=

4.88

8.0632

56.11

0.14

0.7007

5.39

C5+ Gasoline

44.55

LCO

27.60

MCB Coke Total Products

13.80

6.15

47.29

Values to the right

11.30

3.12

23.99

5.22

were determined from Chart

10.20

0.53

4.10

5.62

2B1.1 & Fig 2.

4.44

0.25

1.92

39.74

12.98

99.82

100.00

See attached method.

Total

Laboratory Summary for Gasoline Distillation D-86

°F

°R

(°R)^1/3

SpGr

0.7286

IBP

96.8

556.8

---------

API

62.02 214.7

10%

131.9

591.9

8.3962

VABP

20%

145.4

605.4

8.4596

Engler Slope

2.40

30%

163.4

623.4

8.5426

Correction Factor

6.0

40%

182.3

642.3

8.6280

CABP, °R Uncorrected

672.8

50%

203.9

663.9

8.7237

CABP, °F Corrected

206.8

60%

230.9

690.9

8.8404

CABP, °R Corrected

666.8

70%

259.7

719.7

8.9616

UOP K

11.99

80%

291.2

751.2

9.0904

Total Sulfur, wt%

0.0032

90%

323.6

786.3

9.2193

Octane (F1 C)

92

EP

371.3

831.3

---------

RVP @ 378°C, kg/cm

Recovered Volume

99.5

%

0.5

%

Residue Volume

2

39.6

157048 Process Calculations Page 87

Hydrogen Content of Liquid Petroleum Hydrocarbons

157048 Process Calculations Page 88

157048 Process Calculations Page 89

Comments on Figure 2B1.1 Purpose: The various average boiling points which are used to characterize petroleum fractions are correlated in Figure 2B1.1 with the ASTM D86 distillation properties of the fraction. If these boiling points are required for mixtures (or portions of a mixture) for which the composition is known, using the defining equations (2-0.3) through (2-0.7) given in the introduction. Reliability: The reliability is unknown. Notation: The volumetric average boiling point of a petroleum fraction is the weighted average of the ASTM D86 distillation temperatures after 10, 30, 50, 70 and 90 percent by volume have been distilled.

T10  T30  T50  T70  T90  5

. The slope is calculated assuming a linear ASTM D86 distillation curve

 T90  T10   in degrees Fahrenheit per percent distilled.  90  10 

between the 10 and 90 percent points 

The relationships between the various average boiling points given in Figure 2B1.1 for petroleum fractions are analogous to those defined by equations (2-0.3) through (2-0.7) for mixtures of identifiable hydrocarbons. Special Comments: For ASTM D86 distillation temperatures above 475°F, use the following correction for cracking: log D  1.587  0.00473  T (2B1.1-1) Where: D = correction to be added to T, in degrees Fahrenheit T = observed distillation temperature, in degrees Fahrenheit If the available distillation data are not from ASTM Method D86, they must be converted by the methods of Chapter 3 to calculate the volumetric average boiling point. Literature Sources: This figure was developed by Smith and Watson, Ind. Eng. Chem. 29 1408 (1937). Equation (2B1.1-1) was given by S.T.Hadden, Gulf Research and Development Company, Pittsburgh, Pa., private communication (1964). Example: Determine the molal average boiling point, weighted average boiling point, cubic average boiling point, and mean average boiling point of a petroleum fraction having the following ASTM D86 distillation properties: Distillation, percent by volume 10 30 50 70 90 Temperature, degrees Fahrenheit: 149 230 282 325 371

VABP 

149  230  282  325  371  271 F 5

Slope 

371  149  2.78  F % 80

Using Figure 2B1.1, the average boiling points are calculated from the volumetric average boiling point:

MABP  271  30  241 F WABP  271  7  278 F

CABP  271  7  264 F MeABP  271  19  252 F

157048 Process Calculations Page 90

Calculation of Flow Meter Constant “K” for Liquid Flow:  Gf   Q max  N  S  D 2  Fa  Fc  hm    Gb    Where: Qmax N D d S Fa Fc Gf Gb hm Note:

= = = = = = = = = =

maximum flow rate at base conditions, (@ 60°F) constant based on flow units Process pipe inside diameter orifice diameter discharge coefficient, f(d/D) thermal expansion of plate Reynolds correction factor liquid specific gravity at flowing conditions gas specific gravity at base condition (@ 60°F) maximum differential pressure (design basis)

N, S, Fa, and Fc can be found in L.Spink’s “Principles and Practices of Flow Meter Engineering”

Example for Sponge Gas Meter Flow, bpsd N, for bpd D d B = d/D S Fa Viscosity, cS Re Fc hm Temp, °F Gb Gf

= = = = = = = = = = = = = =

34,000 194.3 7.981 in 5.034 in 0.6307 0.2672 1.001 for type 304 stainless steel plate 10 39,292; Re = 92.235*bpsd / (viscosity*D) 1.015 100 in H2O 173°F 0.9266 0.8854 Gf = Gb*VCF = 0.9260 * 0.9562 = 0.8854

 0.8854   Q max  194.3  0.2672  7.9812  1.001  1.015  100    0 . 9266  

Q max  34,119 BPSD The “K” constant can be calculated by using the following equation:

K

Q max* Gb   3,360 BPSD 10  Gf

Where: 10: maximum units meter reading, MR, then:

 Gf  ; Q  3,360  MR    Gb   

Q  _________ BPSD

157048 Process Calculations Page 91

Calculation of Flow Meter Constant “K” for Gas Flow: Pf Q max  N  S  D 2  Fa  Fc  Y hm  Tf  Gb  Z  Where: Qmax N D d S Fa Fc Y hm Pf Tf Gb Z Note:

= = = = = = = = = = = = =

maximum flow rate at base conditions, (@ 60°F) constant based on flow units Process pipe inside diameter orifice diameter discharge coefficient, f(d/D) thermal expansion of plate Reynolds correction factor upstream orifice expansion factor maximum differential pressure (design basis) absolute flowing pressure upstream of the orifice, psia absolute flowing temperature, °R gas specific gravity at base condition (@ 60°F) compressibility factor of gas

N, S, Fa, Fc, Y and Z can be found in L.Spink’s “Principles and Practices of Flow Meter Engineering”

Example for Sponge Gas Meter Flow, scfm N D d B = d/D S Fa Viscosity, cP Re Fc Y hm Pf Tf Cp/Cv Gb Z

= = = = = = = = = = = = = = = = =

8,375 = lb/hr = 26,100 128.78 for scfm and psi 6.065 in 3.5834 in 0.5908 0.2284 1.0003, for type 304 stainless steel plate 0.011 2,472,487; Re = 6.32(lb/hr)/(viscosity*D) 0.988 0.9944; Y = 1-(0.41-0.35B^4)(h/2)/(27.67Pf*(p/Cv)) 200 in H2O 187.7 psia 573°R 1.27 0.7054 0.965 f(Tr,Pr) Tr = Tf/Tc Pr = Pf/Pc

Q max  128.78  0.2284  6.065 2  1.0003  0.9881  200 

Q max  10,490scfm The “K” constant can be calculated by using the following equation:

 Gb   Q max    1,539scm K    Tf    10   Pf  Where: 10: maximum units meter reading, MR, then:

Q  1539  MR 

Pf  Gb ; Tf

Q  _________ scfm

187.7 573  0.7054  0.965

157048 Treating Page 1

FEED/PRODUCT TREATING INTRODUCTION FCC feeds contain a number of contaminants that affect yields, product quality, plant emissions and corrosion in the main column and gas concentration unit. These contaminants are handled by a combination of treating either the feed or products as well as unit design. This subject is of increasing importance to refiners with the ever tighter limits on emissions from the plant, especially SOx and NOx and on limits in liquid product sulfur levels. The industry trend towards processing resid feeds which typically have higher concentrations of sulfur, metals and carbon residue makes this issue even more important. In the United States new fuel specifications are will limit the sulfur concentration in both gasoline and high speed diesel fuels to less than 50 wppm. This is an issue critical to the FCC because in a typical refinery gasoline pool 98% of the total gasoline pool sulfur comes from the FCC naphtha even though the FCC naphtha makes up only 30-40% of the pool. Limits on fuel oil sulfur will also require a reduction in the MCB product sulfur. In some areas the limits on SOx emissions will be more restrictive requiring less than 300 ppm in the flue gas. In the coming years these restrictions will likely become even more stringent. FEED TREATING Light and heavy vacuum gas oils are the most common FCC feedstock with an increasing trend towards atmospheric resid. Also, there are economic incentives towards processing lower priced crudes which typically contain higher levels of contaminants. Hydrotreating is the most common and effective method of improving the FCC feed quality. Hydrotreating not only reduces the contaminant concentration but also improves yields. Hydrogen addition to the feed, especially to the large polynuclear aromatics, makes these molecules easier to crack resulting in higher conversion to desired products with less coke and light gas make. Table 1 shows the impact of hydrotreating on both the FCC feed properties and the FCC yields.

157048 Treating Page 2

Table 1 Feed Hydrotreating Benefits Feed Desulfurization Feed Properties Gravity, ºAPI Sulfur, wt% Nitrogen, wppm Carbon Residue, wt% Metals (Ni + V), wppm Yields, wt% H2S C2LPG Naphtha LCO MCB Coke Conversion, lv% Key Product Properties Naphtha RONC Naphtha MONC LCO Cetane Index Product Sulfur, wppm H2S Naphtha LCO MCB SOx , vppm in flue gas

Untreated

90%

98%

99%

20.5 2.6 880 0.4 5

23.5 0.25 500 0.25 2

24.8 0.06 450 0.1 1

26.0 0.02 400 0.1