Jones D.S.J., Pujado P.R. Handbook of Petroleum Processing

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242 CHAPTER 6

1500

1000

500

1975

1980

1985

1990

1995

2000

2005

Start-Up Year

Resid Capacity (1000’s B/D)

2000

Figure 6.2. Growth in resid cracking worldwide.

While the oleﬁns in gasoline have become a concern in the gasoline pool, the light

oleﬁns for petrochemicals are a valuable product, often exceeding the revenue ob-

tained for transportation fuels. As a consequence, distinctive processes for making

much larger amounts of propylene than a normal FCC unit have been developed. The

Deep Catalytic Cracking Process (1991) was the ﬁrst commercial scale process that

was designed to maximize propylene. Specially formulated catalysts, more severe

process conditions and equipment made to handle the unique product distribution

are all components of this technology. Table 6.3 compares the yields from the DCC

technology with those obtained in a normal FCCU. Commercial yields have veriﬁed

the laboratory data.

Table 6.2. Typical resid cracking yields

Unit A Unit B

Feed properties

API gravity 0.922 0.963

Conradson carbon 3.8 8.0

Ni + V 4 7.5

Wt% Wt%

Conversion 70.0 70.0

Dry gas 2.1 3.4

LPG 15.3 14.4

Gasoline 46.2 43.1

LCO 16.3 16.0

Decant oil 13.7 14.0

Coke 6.5 9.1

FLUID CATALYTIC CRACKING 243

Table 6.3. Deep oil cracking versus FCCU

Process FCC DCC type 1 DCC type 2

Yields, wt%

Dry gas 3.5 11.9 4.0

LPG 17.6 42.2 34.5

+ Gasoline 55.1 27.2 41.6

LCO 10.2 6.6 9.8

DO 9.3 6.1 5.8

Coke 4.3 6.0 4.3

Ethylene 1.1 6.1 1.6

Propylene 4.9 21.0 14.3

Butylenes 8.1 14.3 14.7

Feedstock: Chinese Waxy VGO

Even more severe cracking conditions are used in the Catalytic Pyrolysis Process

(2002), where the desire is to produce all petrochemical products, i.e. ethylene, propy-

lene, butenes, and aromatics. This process is really a substitute for a steam cracking

furnace in an ethylene plant. It allows the operator to use cheaper feedstocks and

vary the ratio of ethylene to propylene over a wider range than is possible with only

thermal cracking.

A diagram of the reactor-regenerator of a modern ﬂuid cracking unit is shown in

Figure 6.3. Hot regenerated catalyst contacts the oil near the base of the reactor

riser. Virtually all of the cracking takes place in the feed riser which connects to a

Riser Termination Device

Regenerator

Two Stage Cyclones

Product Vapors

Reactor

Single Stage Cyclones

Reactor

Stripper

Spent Cat. Standpipe

Steam Rings

Withdrawal Well

Reactor Riser

Recycle Nozzles

Feed Nozzles

Regenerator

Air Distributor

Air Ring

Regen Cat. Standpipe

Slide Valve

Figure 6.3. Reactor-regenerator of modern FCCU.

244 CHAPTER 6

catalyst/vapor separator. The hot catalyst is discharged into the catalyst stripper while

the vapor is routed through secondary cyclones to remove any remaining catalyst and

then to the main fractionator and gas plant for separation of the products.

The spent catalyst enters a multi-stage stripper where the absorbed hydrocarbons are

displaced with steam and leave with the product from the overhead of the reactor.

This prevents unwanted hydrocarbons from entering the regenerator, consuming air

and possibly causing excessive catalyst deactivation. Staging is accomplished by

the use of bafﬂes or packing and is similar in concept to a multi-tray distillation

tower.

Spent-stripped catalyst enters the regenerator where the coke is burned off the catalyst

to restore its activity. The heat generated in the combustion process can supply all

of the needed heat for the process. If excess heat is produced it can be removed by

external catalyst coolers. The regenerated catalyst ﬂows back to the base of the riser

where the cycle is completed. Typical catalytic cracking units undergo 100–400 such

cycles a day.

There have been many advances in the ﬂuid catalytic cracking process. The high

capacities of a unit (typically a third of the crude oil a fuels reﬁnery runs goes to a

standard gas oil FCC unit), and its positive inﬂuence on overall reﬁnery economics

has made it a prime target for innovation. A list of most of the major innovations in the

process is given in Table 6.4. These include catalyst, equipment and process changes

that have occurred on a continual basis over the 60 years that FCCUs have been

operating. By being able to both improve its performance and evolve its functions,

the catalytic cracking process has remained a staple in the modern reﬁnery.

Fluidization

The basic circulating, ﬂuid bed reactor system that is widely used today in many

other applications got its start in ﬂuid catalytic cracking. It was the development of

the standpipe that made the process possible. The ﬁne powder acts as a ﬂuid when

contacted with a gas and it is the standpipe that allows the catalyst to be circulated from

a vessel at lower pressure to one of higher pressure, thus completing the circulating

catalyst loop.

In Figure 6.4, the head gain from the ﬂuidized catalyst is given as a function of the bed

density. The latter depends on the catalyst properties, the amount of aeration being

used and the rate at which catalyst is being added to the system.

The change in density of a bed of cracking catalyst with gas velocity is shown in Fig-

ure 6.5. At very low velocities the bed is stationary but soon reaches a point where

FLUID CATALYTIC CRACKING 245

Table 6.4. Major advances in ﬂuid catalytic cracking

1942 First FCC unit on stream (Exxon)

1947 Stacked conﬁguration (UOP)

Compact with small inventory

1948 Spray dried catalyst (improved ﬂuidization)

1952 Synthetic high alumina

1955 Reactor riser cracking (shell)

1959 Semi-synthetic catalyst (addition of clay)

1960 Improved metallurgy

(Higher regenerator temperatures)

1961 Heavy oil cracking (Phillips-Kellogg)

1964 Zeolitic catalysts introduced (Mobil)

1972 Complete CO combustion process (Amoco)

1974 Combustion promoters (Mobil)

1975 Metals passivation (Phillips)

1981 Two independent regenerators for unlimited

Regeneration temperature (Total)

1982 High performance feed injectors

1982 Dense bed catalyst coolers (Ashland/UOP)

1987 Vapor quench (Amoco)

1987 Mix temperature control (Total)

1988 Close-coupled cyclones (Mobil)

1991 Deep catalytic cracking (RIPP/Sinopec)

1996 Enhanced stripping designs

2002 Catalytic pyrolysis (RIPP/Sinopec)

the upward force of the gas balances the weight of the catalyst (minimum ﬂuidiza-

tion velocity) and the catalyst becomes suspended in the gas. More gas expands the

ﬂuidized bed to a point where the next bit of gas enters and forms bubbles. This

is the minimum bubbling velocity and the two phases are referred to as the emul-

sion and bubble phases, respectively. Once bubbles are formed all of the added gas

0.0

01020

40 50 60

1.0

2.0

3.0

4.0

5.0

STATIC HEAD GAIN - PSI

CATALYST DENSITY – LBS / FT

Figure 6.4. Static head gain per 10



of standpipe versus catalyst density.

246 CHAPTER 6

0.01 0.1

Catalyst Density (lb/ft

)

Bulk 55.0

Densities 51.2

lb/ft

45.6

Gas Velocity (ft/sec)

Figure 6.5. Fluidization curves FCC catalysts.

contributes to this phase. At even moderate velocities the catalyst is entrained with

the gas and must be captured and returned to the bed. This allows operation of the

catalyst beds at much higher superﬁcial velocities than would otherwise be possible.

Typical velocities in the reactor-regenerator system are given in Table 6.5.

Bubbles play a key part in the operation of a ﬂuid bed system. They are the ‘engine’

that stirs the bed of catalyst making the high heat and mass transfer rates possible. In

the standpipes bubbles can have a profound effect on the smoothness of catalyst ﬂow

and care must be exercised to avoid over-aerating as well as under-aerating the catalyst.

A formula for adding the proper amount of air to a standpipe, Q, was presented by

Zenz and is shown below:

Q = 2000





−



−



−



(1)

Table 6.5. Superﬁcial velocities

for FCC (ft/sec)

Minimum ﬂuidization 0.01

Minimum bubbling 0.10

Bubbling bed 0.3–2

Turbulent bed 2–4

Fast-ﬂuidized bed 4–8

Pneumatic conveying 12

FLUID CATALYTIC CRACKING 247

Figure 6.6. Advanced FCC process control scheme. Reprinted with permission from HYDROCARBON

Typically 70% of the above calculated aeration rate is about optimum and it should be

added at even intervals, around the standpipe, about six to eight feet apart. Fine-tuning

is done once the unit is in operation and a single gauge pressure survey can be taken

to aid in the operation of the system. It is critical that the catalyst be properly ﬂuidized

entering the standpipe since reﬂuidizing catalyst is very difﬁcult. Over-aeration is a

more common problem than under aeration in commercial operations.

Process control

While there are a number of ways to conﬁgure the controls of the cat cracker, Fig-

ure 6.6 shows the typical scheme used in the industry for a slide valve controlled unit.

The reactor pressure is controlled by the inlet pressure to the wet gas compressor

but the actual pressure at the top of the reactor in the dilute phase is this pressure

plus the pressure drop taken through the main fractionator overhead system, the main

fractionator and the reactor overhead system (cyclones, plenum, and overhead vapor

line). At steady state this pressure should be constant. The regenerator pressure is

controlled by the ﬂue gas slide valve and is maintained to give a constant differential

pressure between the reactor and regenerator. This is adjusted to provide optimum

catalyst circulation, adequate pressure drops across the control valves and to balance

the wet gas compressor and air blower.

The reactor temperature is controlled by the regenerated catalyst slide valve. More

hot catalyst is circulated to raise the reactor temperature or accommodate increased

248 CHAPTER 6

feedrates. The feed going to the unit is preheated by exchange with the hot products.

While heat exchange is adequate for feed temperatures ranging up to about 500

◦

F, a

ﬁred heater is included for higher temperatures.

The regenerator temperature is a function of the coke on the catalyst entering the

regenerator, its composition and the mode of coke burning. Full combustion or the

complete burning of the carbon to CO

is most common due to the necessity of

limiting the CO concentration in the ﬂue gas to 500 ppm or less to meet environmental

regulations. Partial combustion resulting in CO/CO

ratios ranging up to one are

practiced in units that do not have the needed metallurgy in the regenerator to operate

at high temperatures or wish to maximize regenerator coke burning capacity or limit

the heat produced in the regeneration process. Such units are equipped with CO

incinerators or boilers where the CO is converted to CO

before being discharged to

the atmosphere. Typical regenerator temperatures in these two operating modes are

1150–1250

◦

F (620–675

◦

C) and 1275–1350

◦

F (690–732

◦

C), respectively for partial

and complete combustion.

The spent catalyst slide valve controls the catalyst level in the stripper or reactor bed

if one is being used. No direct control of the regenerator catalyst bed level is made and

its level ﬂoats depending on the catalyst losses and the catalyst withdraw and addition

policies. Most units are capable of retaining most of the fresh catalyst added daily and

must periodically withdraw equilibrium catalyst to keep the regenerator level from

getting too high. The regenerator is used as the ‘ﬂoating’ vessel because it is larger

and it is desirable to withdraw regenerated rather than spent catalyst.

With the advent of computers and advanced process controllers, many reﬁnements

to the basic scheme are possible. The biggest beneﬁts come from operating closer to

several limits at one time. Better analysis of the feedstocks could allow feed-forward

control in the future. While these control systems can improve reﬁning proﬁtability

from 20 to 40 cents/barrel processed, they require more instrumentation, which must

be maintained to achieve the stated beneﬁts.

Reaction chemistry and mechanisms

Many studies have been performed elucidating the difference between thermal and

catalytic cracking. Two separate reaction mechanisms are attributed to the methods

of cracking, i.e., thermal cracking goes through free radicals and catalytic cracking

via carbenium ions. The latter are generally associated with the bronsted acid sites

on the catalyst. In Table 6.6, the major differences between the two mechanisms are

shown. Thermal cracking is minimized as much as possible in current FCC units by

the use of advanced equipment such as radial feed injectors, riser termination devices

and post riser quench. Catalyst selection is also critical.

FLUID CATALYTIC CRACKING 249

Table 6.6. Characteristics of cracking mechanism

Thermal—Free radical Catalytic—Carbenium ion

–C

Principle products

Little skeletal isomerization

Much skeletal isomerization

Cracking at beta position with little

preference for free radical type

Aromatics dealkylate at ring if

chain is at least 3 carbons long

Alpha oleﬁns primary product

cracking occurs at

tertiary > secondary > primary

carbeniums ions

Products contain an oleﬁn

Largest molecules crack fastest

unless steric hindrance controls

There are many reactions that occur during the cracking process. These are listed

below as primary or secondary reactions (Table 6.7). Most of the secondary reactions

are undesirable and are controlled through reactor and catalyst design. Numerous

cracking reactions occur with the large feed molecules before the desired products

are achieved. Typical feeds to a catalytic cracking unit contain molecules boiling above

the diesel end point (650–700

◦

F or 343–371

◦

C) and may boil as high as 1500

◦

F. The

consecutive reactions that occur are shown in equation (2):

ABCD

Gas Oil or Resid −→ Diesel −→ Gasoline −→ LPG + Coke

(2)

Table 6.7A. Primary cracking reactions for hydrocarbon types

Parafﬁn Parafﬁn + Oleﬁn

2n+2

−→ C

2p+2

+ C

Naphthene (cyclic parafﬁn) Oleﬁn + Oleﬁn

−→ C

+ C

Alkylaromatic Aromatic (base) + Oleﬁn

ArC

2n+1

−→ ArH + C

Oleﬁn Oleﬁn + Oleﬁn

−→ C

+ C

Aromatic −→ No reaction

Table 6.7B. Secondary reactions of oleﬁns

Naphthene + Oleﬁn −→ Aromatic + Parafﬁn

Hydrogen + Oleﬁn

−→ Parafﬁn

Normal Oleﬁn

−→ Iso-Oleﬁn

Oleﬁn + Oleﬁn

−→ Larger Oleﬁn

Aromatic + Oleﬁn −→ Alkylaromatic −→ Cyclization

Multi-ring aromatic (Coke) ←− Dehydrogenation ←−

250 CHAPTER 6

While this is an over simpliﬁcation of what occurs, it does give an overall view to

the cracking process. From the above equation it is clear that separate processes or

applications can center around either product or feed differences.

As such, there are three basic catalytic cracking applications, today. These are:

Application Feedstock Products

1. Gas oil cracking Vacuum gas oils Motor gasoline

LCO and LPG

2. Resid cracking Atmospheric resid

VGO + Vacuum resid

Motor gasoline

LCO and LPG

3. Cracking for

petrochemicals

Vacuum gas oils and

added resids.

Light oleﬁns—C

’s , C

’s ,

and C

’s

Plus aromatics

The driving forces behind these applications are the need for gasoline, the lack of

demand for bottom of the barrel and the increase in demand for light oleﬁns and aro-

matics. The technologies for each of these process conﬁgurations are reviewed in turn.

Gas oil cracking technology features

Reaction technologies

All of the reaction systems offered today consist of a feed injection system, reactor

riser, riser termination device, and sometimes vapor quench technology. Each licensor

approaches the design with different equipment and conﬁgurations.

The feed injection system is probably the most important part of the reaction system

since it provides the initial contacting between the oil and catalyst. A good system must

vaporize the feed quickly, quench the hot catalyst as fast as possible and provide plug

ﬂow of the hydrocarbons and catalyst. Other features are important but are beyond

this limited description.

The hot-regenerated catalyst generally comes to the feed riser either through a wye

section or a J-bend depending on the proximity of the regenerated catalyst standpipe

and the feed riser (see Figure 6.7). Feed nozzles are usually located from 10 to 30

ft downstream of the base of the riser. Since raising the nozzles increases the back

pressure on the regenerated catalyst slide valve, lift gas may be incorporated to control

the pressure balance and minimize the back-mixing of the catalyst. The lift gas is

normally light gas from the product recovery section and adds to the load in the main

fractionator and wet gas compressor. Steam is sometimes substituted either whole or

in part to reduce the processed gas but can cause some catalyst deactivation if the

regenerator temperature gets too high.

FLUID CATALYTIC CRACKING 251

J-Bend

Wye Section

Figure 6.7. Regenerated catalyst/feed contacting conﬁgurations.

The feed injection systems (Figure 6.8) employed by the various licensors all use

different nozzle designs. Oil pressure and dispersion steam rates vary by both licensor

and application. Typical oil pressure drops and dispersion steam rates are 30–150 psi

and 1–7 wt% and depend on the type of feedstock processed, i.e. vacuum gas oils or

residual feeds (BP > 1050

◦

C (565

◦

C)).

The feed injection systems being offered by the licensors are shown in Figure 6.8.

These are all two ﬂuid nozzles that mix an oil feed and steam and distribute them

across the riser cross-sectional area. Differences exist in the methods of contacting the

steam and oil, the amount of steam and pressures used and in the tip design. Reﬁners

are concerned about both performance and reliability when choosing a system.

Reactors are vertical pipes that are generally straight and are about 100 ft long. The

diameters at each end are controlled to give adequate lift to the catalyst. Velocities

of 20–30 ft/sec at the base and 50–65 ft/sec at the vapor outlet are typical with vapor

residence times of about 2 sec based on the vapor outlet velocity. Bends in the riser

pipe do not change the vapor residence time but increases the catalyst residence time

(or slip) and usually results in poorer oil/catalyst contacting.

A termination or separation system is used at the top of the riser. These have been

simple Tee’s, rough cut cyclones or specialized inertial separators. Most recently the

riser termination devices have been close-coupled cyclones, linear disengagers, riser

separator strippers, or vortex separators. These are pictured in Figure 6.9 and all can

be or are directly connected to the secondary cyclones.