chemical processing and design

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Encyclopedia of

Chemical Processing and Design

Main entry under title:

Encyclopedia of chemical processing and design.

Includes bibliographic references.

1 Chemical engineering—Dictionaries 2 Chemistry,

Technical—Dictionaries I McKetta, John J.

II Cunningham, William Aaron.

ISBN: 0-8247-2621-9

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Steve Chum, Ph.D. Research Fellow, The Dow Chemical Company, Polyolefins

Research, Freeport, Texas: Structure, Properties, and Applications of Polyolefins Produced by Single-Site Catalyst Technology

Ray A Cocco, Ph.D. Senior Specialist, The Dow Chemical Company, Midland,

Michigan: Circulating Fluidized Bed Reactors: Basic Concepts and namics

Hydrody-David M Fishbach, P.E. Senior Consulting Engineer, Starfire Electronic

Devel-opment & Marketing, Ltd., Bloomfield Hills, Michigan: Nanophase Materials in Chemical Process

Avery N Goldstein, Ph.D. Research Director, Starfire Electronics

Develop-ment & Marketing, Ltd., Bloomfield Hills, Michigan: Nanophase Materials in Chemical Process

Manfred Grove Senior Partner, Intermacom A.G., Technology Consultants,

Auerich, Switzerland: Introduction to the Selective Catalytic Reduction nology

Tech-Dennis Hendershot Rohm and Haas Company, Bristol, Pennsylvania: mentals of Process Safety and Risk Management

Funda-Trevor A Kletz Process Safety Consultant, Cheshire, United Kingdom: mentals of Process Safety and Risk Management

Funda-Fu-Ming Lee, Ph.D. Director of Technology, GTC Technology Corporation,

Houston, Texas: Recent Development of Extractive Distillation: A Distillation ternative

Al-Jim Makris Director, Chemical Emergency Preparedness and Prevention Office,

U.S Environmental Protection Agency, Washington, D.C.: Process Safety and Risk Management Regulations: Impact on Process Safety

M Sam Mannan, Ph.D., P.E. Associate Professor of Chemical Engineering,Mary Kay O’Connor Process Safety Center, Texas A&M University, College Sta-

tion, Texas: Fundamentals of Process Safety and Risk Management; Process Safety and Risk Management Regulations: Impact on Process Industry

Rajen M Patel, Ph.D. Technical Leader, The Dow Chemical Company,

lefins Research, Freeport, Texas: Structures, Properties, and Applications of lefins Produced by Single-Site Catalyst Technology

Polyo-H James Overman Dow Chemical Company, Freeport, Texas: Process Safety and Risk Management Regulations: Impact on Process Industry

iii

Michael V Pishko, Ph.D. Assistant Professor, Department of Chemical

Engi-neering, Texas A&M University, College Station, Texas: Recent Advances in materials

Bio-Alan W Weimer, Ph.D., P.E. Professor of Chemical Engineering, University

of Colorado, Boulder, Colorado: Effect of Pressure and Temperature in Bubbling Fluidized Beds

Contributors to Volume 69 iii

Circulating Fluidized Bed Reactors: Basic Concepts and Hydrodynamics

Effect of Pressure and Temperature in Bubbling Fluidized Beds

Fundamentals of Process Safety and Risk Management

M Sam Mannan, Dennis Hendershot, and Trevor A Kletz 49 Introduction to the Selective Catalytic Reduction Technology

Nanophase Materials in Chemical Process

Avery N Goldstein and David M Fishbach 150 Process Safety and Risk Management Regulations: Impact on

To convert from To Multiply by

vii

To convert from To Multiply by

Cost escalation via inflation bears critically on estimates of plant costs Historicalcosts of process plants are updated by means of an escalation factor Several pub-lished cost indexes are widely used in the chemical process industries:

Nelson Cost Indexes (Oil and Gas J.), quarterly

Marshall and Swift (M&S) Equipment Cost Index, updated monthly

CE Plant Cost Index (Chemical Engineering), updated monthly ENR Construction Cost Index (Engineering News-Record), updated weekly Vatavuk Air Pollution Control Cost Indexes (VAPCCI) (Chemical Engineering),

updated quarterlyAll of these indexes were developed with various elements such as materialavailability and labor productivity taken into account However, the proportionallotted to each element differs with each index The differences in overall results

of each index are due to uneven price changes for each element In other words,

TABLE 1 Chemical Engineering and Marshall and Swift Plant and Equipment Cost

TABLE 2 Nelson-Farrar Inflation Petroleum Refinery Construction Indexes since 1946

Nelson

TABLE 2 Continued

Nelson

Table 1 compares the CE Plant Index with the M&S Equipment Cost Index.Table 2 shows the Nelson-Farrar Inflation Petroleum Refinery Construction In-dexes since 1946 It is recommended that the CE Index be used for updating totalplant costs and the M&S Index or Nelson-Farrar Index for updating equipmentcosts The Nelson-Farrar Indexes are better suited for petroleum refinery materials,labor, equipment, and general refinery inflation

Since

Here, A ⫽ the size of units for which the cost is known, expressed in terms of

capacity, throughput, or volume; B⫽ the size of unit for which a cost is required,

expressed in the units of A; n ⫽ 0.6 (i.e., the six-tenths exponent); C A⫽ actual

cost of unit A; and C B ⫽ the cost of B being sought for the same time period as cost C A

To approximate a current cost, multiply the old cost by the ratio of the currentindex value to the index at the date of the old cost:

million-lb/yr capacity in 1974, find the cost of plant B at a throughput of 300 million lb/yr on the same basis for 1986 Let the sizing exponent, n, be equal to

0.6

From Table 1, the CE Index for 1986 was 318.4, and for 1974 it was 165.4.Via Eq (3),

TABLE 3 Vatavuk Air Pollution Control Cost Indexes (VAPCCI) First Quarter 1994

⫽ 100.0 (index values have been rounded to the nearest tenth)

1994 For details, see the Vatavuk Air Pollution Control Cost Indexes article involume 61

Editor’s note: For a more thorough explanation of updating costs, see the cle, ‘‘Tower Cost Updating’’ in volume 58

arti-john j mcketta

Chemical Processing and Design

69

Basic Concepts and Hydrodynamics

Introduction

Circulating fluidized beds (CFBs) consist of two basic designs, as shown in Fig 1.One design involves a fast-fluidized bed where high gas velocities convey a sub-stantial amount of solids to one or more cyclones The separated particles are fedback to the fluidized bed using a standpipe The second basic design uses a riser

to convey solids to one or more cyclones The separated particles are fed to anoptional fluidized bed and then back to the riser Solids flow rates can be controlledusing nonmechanical L- and J-valves or using a mechanical slide valve

The large-scale commercial realization of CFBs occurred in the early 1940s,although some coal gasification was done in a fluidized bed as early as 1926 [1].With the increased demand for gasoline during World War II, major efforts wereunderway to develop reactors to crack petroleum feedstocks into usable fuels moreproductively than the moving bed or snake reactors (i.e., the Houndry Process)used at that time The result was a fluidized catalyst cracker (FCC), where highcatalyst circulation rates allowed a balance between the exothermic burning ofcoke on the catalyst in the regenerator and the endothermic hydrocracking of petro-leum in the reactor The continuous circulation or regeneration of catalyst providedfresh catalyst for petroleum cracking and thereby resulted in high sustainable pro-ductivities With the addition of a stripping section after the reactor, even higheryields were obtained The addition of steam, CO2, or other inerts would removethe product from and around catalyst particles flowing toward the regenerator.Today, the evolution of the FCC unit has results into several basic designs, asshown in Fig 2

In 1960, circulating fluidized beds contributed to another breakthrough processfor the petroleum and chemical industry Standard Oil of Ohio (SOHIO) developed

a fluidized-bed reactor for the ammoxidation of propene to acrylonitrile Previoustechnology was done in tube-and-shell fixed-bed reactors However, the high heat

of reaction of 160 kcal/mol limited the economic feasibility of those units Thehigh heat transfer characteristic of fluidized-bed reactors made them ideal for theproduction of acrylonitrile Today, nearly all large-scale acrylonitrile plants arebased on the SOHIO design, with capacities up to 180,000 tons per year [4].The greatest challenge in developing the SOHIO process was in the manage-ment of backmixing The inherent hydrodynamics of fluidized beds, where solidsand, to a lesser extent, gas circulate from the top of the bed to the bottom, then

to the top again, would have a deleterious effect on acrylonitrile selectivity Toovercome backmixing, SOHIO developed sieve trays to compartmentalize the gasflow in the fluidized-bed reactor to resemble a more plug-flow characteristic [5]

In 1979, SOHIO redesigned the acrylonitrile reactor to more of a ‘‘tube-and-shell’’fluidized-bed unit [6], as shown in Fig 3

1

FIG 1 Basic design of circulating fluidized beds.

FIG 2 Typical FCC units based on the designs of (a) Standard Oil Development, (b) UOP, (c) logg, and (d) Exxon (Adapted from Refs 2 and 3.)

Kel-FIG 3 Two-dimensional schematic of the SOHIO acrylonitrile processes (Adapted from Ref 5.)

During the late 1970s and early 1980s oil crisis, circulating fluidized beds foundapplications in coal combustion The high-heat-transfer capabilities of these re-actors resulted in lower operating temperatures, thereby reducing NOxand SO2

emissions In addition, the high gas velocities resulted in significant turbulence,which provided uniform temperatures in the combustor With the surplus of oilstarting in the late 1980s, fluidized-bed combustors became economically less at-tractive As of the early 1990s, only Dynergy (via the Destec process) and Lurgiand Ahlstrom are practicing this technology [7]

Today, circulating fluidized beds are used in a wide array of chemical cesses, as shown in Table 1 With fluidized beds having the unique distinction of

pro-excellent heat transfer and continuous in situ regeneration, the economic

attrac-tiveness of processing thermally sensitive chemicals or using catalysts that require

TABLE 1 Some Fluidized and Circulating Fluidized Bed Reactor Processes

frequent regeneration are more realized Once the obstacles of backmixing, masstransfer, and attrition have been addressed, these reactors often set the standards

in reactor design

Basic Concepts

As the gas velocity through a bed of solids increases, the bed undergoes severalregimes, as shown in Fig 4 At first, the gas velocity is insufficient to fluidizethe particles and the bed remains fixed With increasing velocity and under idealconditions, the fixed bed expands smoothly and uniformly Particles move in alimited fluidlike fashion and the bed pressure drop becomes constant At this point,the bed is commonly referred to as undergoing minimum fluidization

Further increases in gas velocity results in further bed expansion and particlesappear to freely move throughout the bed The gas permeates through the bedwithout the formation of bubbles This regime is referred to as smooth fluidizationand is only observed for Geldart Group A powders (see Appendix A) These pow-ders require noticeably higher gas velocities to promote the formation of gas bub-bles after minimum fluidization In contrast, Group B powders begin bubblingshortly after minimum fluidization Group C powders, being cohesive, may evenshow signs of bubbling prior to minimum fluidization; however, this is usually theresult of channeling

The onset of bubbles in the fluidized bed is commonly referred to as bubblingfluidization Here, gas bubbles form at or near the distributor and grow to a maxi-mum bubble size as they propagate through the bed The top of the fluidized bed

is still well defined, as it was in the minimum and smooth fluidization regimes.The pressure drop across the bed is still constant, on average, but starts exhibitinglarge, but regular, fluctuations with time

As the gas velocity continues to increase, the top of the bed becomes lessdefined Large amounts of particles are ejected into the freeboard region abovethe bed Concurrently, sizeable regions of voidage and particle clusters are seen

FIG 4 Various fluidization regimes with increasing superficial gas velocity.

in the bed itself For Group A and B powders, this transition from the bubblingfluidized-bed regime is called the onset of turbulent fluidization Group C and Dpowders may show a slugging behavior prior to the turbulent fluidization regime.During fast fluidization, the gas velocity is sufficient enough that the surface

of the bed can no longer be discerned Particle density is still higher at the bottom

of the unit compared to the top, suggesting that some sort of bed exists Particleclusters and streamers are readily observed and, in some cases, a core–annulusradial variation in particle density begins to take shape Particle entrainment is highand the total disengagement height may be well beyond the physical dimensions ofthe fluidized-bed unit To overcome the losses of particles due to entrainment,cyclones may be used to capture entrained particles and recirculate them back intothe bed

At very high gas velocities, nearly all the particles are entrained from the bed.This regime is commonly referred to as pneumatic conveying In this regime, axialvariation in particle density is no longer observed, except maybe in entrance andexit regions Radial variation in particle density can vary dramatically and rangefrom a core–annulus profile to a uniform profile For dense systems, clusters andstreamers are readily observed

Thus, for gas–solid systems, increases in the gas velocity results in dramaticand sometimes sharp transitions in the hydrodynamics In the design of fluidizedbeds, it is crucial that one knows the fluidized regime that will exist at operatingconditions The simple transition from one regime to another can have significantimpacts on reaction, heat transfer, attrition, and entrainment rates For circulatingfluidized beds, where several regimes may exist in a single unit (i.e., from con-veying in the riser to a bubbling fluidized-bed regime in the regenerator), knowl-edge of the fluidization regimes is paramount

In order to gain better understanding of these regimes, the methodology used

to determine the onset of each fluidization regime is discussed in the followingsections Keep in mind that most of the correlations are empirical and may notfully represent every system With the cost of these units running in the tens ofmillions of dollars for large-scale plants, experimental validation of the expectedregimes is critical when designing these processes

Minimum and Smooth Fluidization

As a gas permeates through a fixed or packed bed, the pressure drop can be scribed by the Ergun equation [8]:

referred to as the minimum fluidization velocity or u Assuming that the weight

of the particles in the fluidized bed corresponds to∆P/L, the Ergun equation can

where the particle Reynolds number at minimum fluidization is a simple function

of the Archimedes number and two constants (K1/2K2and 1/K1)

Many correlations for the minimum fluidization velocity are based on Eq (7)

TABLE 2 K1and K2Values for Eq (7)

Source: Adapted from Ref 7.

for the constants K1/2K2and 1/K1 These constants are presented in Table 2 for

a wide range of studies For typical Geldart Group A powder, the constants ofWen and Yu are most often used However, these correlations are specific to agroup of particles with common characteristics and may not represent a less-than-ideal particle morphology and texture

The minimum fluidization velocity can be experimentally determined by suring the pressure drop across a bed of particles with increasing superficial gasvelocity For smooth, round, and noncohesive particles, the pressure drop increaseslinearly with gas velocity until the minimum fluidization velocity is reached Withfurther increases in the gas velocity, the pressure drop remains constant Hence,the minimum fluidization velocity is the intersection of the linearly increasing linewith the constant-pressure-drop line

mea-Figure 5–8 demonstrate the results of such an experiment mea-Figure 5 is thepressure-drop curve for alumina particles with a mean particle diameter of 60µm

in a 4.5-in.-inner diameter fluidized bed unit The minimum fluidization velocityfor these particles was determined to be 6.5 cm/min When measuring the mini-mum fluidization velocity, less scatter in the data is obtained from larger or higherbeds The scatter in Fig 5 suggests that perhaps a higher bed should have beenused The diameter of the fluidized bed used in this type of experiment is also

FIG 5 Minimum fluidization curve for smooth and round alumina particles, d ⫽ 60 µm.

FIG 6 Minimum fluidization curve for smooth and round alumina particles, d p,ave⫽ 60 µm, with increasing and decreasing gas velocities.

critical to obtaining accurate data For Geldart Group A powders, bed diametersless than 3 in can result in experimental data that are influenced by frictionaleffects at the wall For the coarser Group B powders, the minimum diameter ismuch larger

Figure 6 shows two pressure drop versus superficial gas velocity curves forthe same particles used in Fig 5 The black data points are the pressure drop withincreasing gas velocity and the gray data points are the subsequent pressure-dropmeasurements with decreasing gas velocities For round, smooth, and noncohesiveparticles, the two curves should overlap each other, as shown in Fig 6 However,for irregular, rough, or cohesive particles, a hysterisis effect is typically observed.This is obvious in Fig 7 for rough and irregularly shaped alumina particles with

FIG 7 Minimum fluidization curve for rough alumina particles, d ⫽ 92 µm.

FIG 8 Minimum fluidization curve for cohesive iron catalyst particles, d p,ave⫽ 68 µm.

a mean particle diameter of 92µm This resulted in higher solid shear forces duringfluidization such that the pressure drop is dependent on previous conditions or ispath dependent Figure 8 shows the pressure-drop curve for a catalyst supported

on alumina with a mean particle diameter of 68µm Here, it is almost impossible

to detect the minimum fluidization velocity High cohesive forces result in a ized bed prone to channeling Each peak or spike in Fig 8 is the result of anotherchannel achieving fluidization while the majority of the bed remains fixed Thisbehavior is typically of Group C powders

fluid-Bubbling Fluidization

As discussed earlier, beds with Group A powders pass from minimum fluidization

to smooth fluidization to bubbling fluidization with increasing gas velocity Group

B and D powders exhibit bubbling fluidization at the onset of minimum tion Oddly enough, gas bubbles in all fluidized beds behave similarly to gas bub-bles in low-viscosity liquids [7] Large gas bubbles are typically spherical on topand flattened or even inverted on the bottom; smaller bubbles tend to be completelyspherical As in liquid systems, gas bubbles in fluidized beds can coalesce intolarger bubbles or split into smaller bubbles, depending on bed conditions Also,

fluidiza-as gfluidiza-as bubbles approach the top of a fluidized bed, they collapse such that solidsare propagated into the freeboard region Higher pressures or temperatures result

in a decrease in the maximum bubble size (due to changes in the gas physicalproperties) and tend to make fluidization smoother [1]

The minimum velocity for bubble formation is referred to as the minimum

bubbling velocity or Umb For Geldart Group A and C powders, Abrahamsen andGeldart [15] proposed that the minimum bubbling velocity can be calculated from

If the bed diameter is small and the bed is sufficiently high, the fluidized bed willslug before entering the turbulent fluidization regime For Group A, B, and Dpowders, slugging is basically the result of a bubble diameter that exceeds abouttwo-thirds the bed diameter The wall stabilizes the bubble such that almost theentire bed is translated up to the top of the bed or even higher Group C powdersmay also exhibit slugging behavior even in large-diameter beds due to the cohesiveforces Thus, the larger and more cohesive the particles or the smaller the beddiameter, the higher the probability of a bed exhibiting slugging For these cases,

the minimum slugging velocity, Ums, can be estimated using the expression ofStewart and Davidson [17], where

tion, umt, can be calculated using the expression of Bi and Grace [20]:

referred to as the minimum transport velocity or umr Schnitzlein and Weinstein

[22], however, were unable to determine umr using Yerushalmi’s method Their

observations suggested that the value of umrwas strongly dependent on the locationand the distance separating the two pressure taps used to measure the pressuredrop

Bi and Grace [20] measured the entrainment rate versus superficial velocity for

a wide range of fluidized-bed systems They noted that the onset of fast fluidizationcorrelated to the point where significant particle entrainment was observed ForGroups A and B powders, this minimum transport velocity can be estimated withthe expression

umr⫽ 1.53 Ar0.45冢 µ

where 2⬍ Ar ⬍ 4 ⫻ 106 For the larger Group D powders, Eq 16 may underpredict

umr compared to using the terminal velocity as the minimum transport velocity.Under these conditions, the minimum transport velocity should be set to equal theterminal velocity [19]

Pneumatic Conveying

As the velocity continues to increase in a fast-fluidized bed, the axial transitionbetween the dense and lean regions disappears This transition marks the onset ofthe pneumatic conveying regime and the superficial velocity corresponding to this

point is called the minimum conveying velocity or umc The onset of pneumaticconveying is readily measured by starting with a high superficial gas velocity anddecreasing its value while holding the entrainment rate constant The superficialgas velocity where the suspension collapse (i.e., choking) is observed corresponds

to the minimum conveying velocity [19] Yang [23] proposed that this velocitycan be estimated with the equation

FIG 9 Flow regime map for various powders Slugging for Group A and B powders depend on vessel diameter Group C powders tend to slug and Group D powders almost always exhibit slugging.

ization corresponds with the onset of bubble formation For Group C powders,smooth and bubbling fluidization are often replaced by channeling due to largesolid stresses or large cohesive forces on these particles Group C powders alsohave a higher tendency to exhibit slugging prior to turbulent fluidization Group

D powders often exhibit slugging during the onset of fluidization The large particlesizes and/or high densities of Group D powder result in the formation of largemaximum bubble diameters that often exceed the diameters of commercial-scalefluidized beds In contrast, slugging for Group A, B, and, to a lesser extent, Cpowders may not be observed in large commercial units where the effective beddiameter often exceeds the maximum stable bubble size

In an effort to develop a design guide for fluidized unit operations, variousattempts have been made to quantify the flow regimes for gas–solid flow Reh[24] first made the attempt for a unified flow regime map by comparing the parti-cle’s Reynolds number, Rep, to the inverse of the drag coefficient, 1/CD Furthermodification to Reh’s map were made by Werther [25] Li and Kwauk [26] andAvidan and Yerushalmi [27] took a different approach to developing a unified flowregime map by comparing the relationship of the superficial gas velocity to solidsvoidage This approach was further modified by Rhodes [28] Yet, another ap-proach was taken by Leung [29], Klinzing [30], and Yang [23], who comparedthe superficial gas velocity to the solid flux

In each of these cases, however, the transition from the fast-fluidized bed to

a pneumatic transport regime was not well defined Grace [31] resolved this lem by generating a unified flow regime map based on the dimensionless variables

prob-d * p and u* of Zenz and Othmer [32] By comparing the dimensionless particle diameter, d * p ⫽ Ar1/3, with the dimensionless gas velocity, u*⫽ Rep/Ar1/3, Gracewas able to discern the transition to pneumatic conveying Kunii and Levenspiel[7] further modified Grace’s work by including the observations of van Deemter[33], Horio et al [34], and Catipovic et al [35] Figure 10 shows the result of thiseffort It is interesting to note that Kunii and Levenspiel’s [7] demarcation fromfast fluidization to pneumatic conveying is not identical to that first proposed byGrace Kunii and Levenspiel describe this transition as not being well defined

With the aid of Fig 10 and the relatively simple calculation of d * p and u*, the

flow behavior of most gas–solid systems can be readily obtained Keep in mind,however, that the boundary for each regime shown in Fig 10 should not be consid-ered as sharp transitions, but more of a gray area between two adjacent regimes

Pressure Balance in a CFB

In designing a circulating fluidized bed (CFB), special attention needs to be given

to the pressure-loop calculations This is especially true in the proper design ofstandpipes and mechanical or nonmechanical valves (i.e., L-valves, J-valves, slidevalves) The bed height in a standpipe counterbalances the dynamic pressure dropoccurring in the riser or fast-fluidized-bed section Mechanical or nonmechanicalvalves provide better response for this counterbalancing For instance, an increase

in the gas flow rate in a riser would require a constriction in a slide valve or lessaeration in a L- or J-valve to prevent backflow This correction results in a higherstandpipe bed height and pressure drop The higher pressure in the standpipe coun-

FIG 10 Flow regime diagram proposed by Kunii and Levenspeil (From Ref 7.)

terbalances the increased pressure in the riser Counterbalancing can also happenwithout the aid of mechanical or nonmechanical valves, but the response time isslow and the probability for backflow is much greater

Figures 11–13 give qualitative examples of the pressure-loop behavior in theCFBs shown in Fig 1 The pressure drop in the riser or fluidized-bed section mustequal the pressure drop in the cyclone, fluidized bed, standpipe, and control valvesections For example, in Fig 11, the pressure loop is defined as

[(P4⫺ P3)]fluid bed⫹ [(P5⫺ P4)]freeboard

⫽ [(P6⫺ P7)⫹ (P5 ⫺ P6)]cyclone (18)

⫹ [(P8⫺ P9)⫹ (P7 ⫺ P8)]standpipe

or

Assuming that particle frictional (both particle–particle and particle–wall) and ticle acceleration effects are negligible, the pressure drop in the fluidized bed can

par-be approximated by

where L is the height of the fluidized bed Similarly, the freeboard can be described

FIG 11 Pressure loop for a CFB of Design I in Fig 1.

using

∆P|freeboard⫽ ρs g冮L u

L b

[1 ⫺ ε(z)] dz ⬇ ρ s g(1 ⫺ ε)(L u ⫺ L b) (21)

where L uis the height of the unit Because the voidage profile,ε(z), in the freeboard

region is rarely known, Eq (21) can be approximated by using the average voidage

in the freeboardε

For the right-hand side of Eq (18), the pressure drop in the cyclone can bedetermined using the expression described by Muschelknautz and Greif [36],where

The first term of this expression corresponds to the friction losses at the wall due

to gas The second terms accounts for momentum losses in the vortex Under theassumption that the flow over the wall of the cyclone is similar to flow over a flatplate, the pressure drop due to wall friction in a cyclone can be calculated using

∆P|wall friction⫽ λpρg Asurface

1.8Q g

(u0u b)0.67 in SI units (23)

The pressure drop in the vortex depends on the average velocity in the vortex tubeand the tangential velocity in the vortex tube or exit This pressure drop can be

v0⫽ V ˙ g

For the standpipe, the pressure drop is determined in a method similar to that usedfor the fluidized bed under the same assumptions The pressure drop in the leanphase of the standpipe can be estimated as

∆P|standpipe, lean phase⫽ ρs g冮L s , 0

L s,b

[(1⫺ ε)(z)] dz ⬇ ρ s g(1 ⫺ ε)(L s , 0 ⫺ L s , b) (26)

where L s,0 and L s,bare the height of the standpipe and the dense-phase bed in thestandpipe, respectively For the dense-phase region of the standpipe, the pressuredrop is approximated as

For the CFB in Fig 12, the pressure loop has the expression

FIG 12 Pressure loop for CFB with a fluidized bed as shown in Design II of Fig 1.

FIG 13 Pressure loop for CFB with no fluidized bed as shown in Design III of Fig 1.

A similar set of expressions can be obtained for the CFB in Fig 13 Compared

to the pressure-loop calculations used for the system in Fig 11, the only newexpressions needed to complete the pressure-loop calculations in Figs 12 and 13are the pressure drop across the L-valve Yang and Knowlton [37] noted that the

pressure drop across an L-valve is similar to that used by Jones and Davidson [38]for a mechanical valve where

2ρs(1⫺ εmf)冢 W s

C D A0冣2

(32)

For mechanical valves, the valve opening area, A o, may vary during operation but

is always known For an L-valve, however, the opening area, A o, depends on theamount of aeration and is not known Yang and Knowlton [37] found an empiricalexpression for this opening area of the L-valve based on 158 experimental datapoints Their results showed that the opening area can be expressed as

Gas–Solid Hydrodynamics

In general, the hydrodynamics of a riser can be divided into macro-scale and scale flow behavior Macro-scale behavior is mostly concern with solids concentra-tion profiles and solids velocity on a large scale Risers typically exhibit a wideand diverse range of axial and radial solids profiles that are highly dependent onoperating and design conditions For instance, the design of the entrance and exitregions of a riser can influence the solid profile throughout the riser Hence, it isimportant to understand this macro-scale behavior in order to provide the correctriser design for the desired solids concentration and velocity profiles

meso-To make matters even more complicated, fast-fluidized beds and risers haveshown evidence of dynamic meso-scale flow behavior in the form of particle ag-glomeration called clusters and streams The size and frequency of these clusters

and streams are also highly dependent on operation and design conditions Boththe cluster size and frequency have a substantial impact on both catalytic reactionrates and heat transfer.

Fortunately, a clearer understanding of these macro-scale and meso-scale haviors are under intensive investigations Today, fast-fluidized beds and riserscan be designed with the solids concentration and radial profile in mind However,this can only be achieved if one has an understanding of gas–solid hydrodynamics

be-Macro-Scale Behavior: Solids Profile

Axial Profile of Solids in a Fast-Fluidized Bed and Riser

In 1971, Reh noted that there exists an axial gradient of solids concentration in ariser similar to that observed in a fast-fluidized bed Figure 14 illustrates thesesubtitle differences in the axial solids concentration profile commonly observedthroughout a fast-fluidized bed and riser As the gas velocity increases through afluidized bed, the boundary between the dense fluidized-bed region and the leanfreeboard region becomes indistinguishable Indeed, having a distinguishable bedheight is one of the indicators for fast fluidization (see the subsection Fast-Fluid-ized Regime) However, it was surprising that this axial gradient also exists in ariser (i.e., pneumatic conveying region) where higher gas velocities are used.This behavior was later quantified by Li and Kwauk [26], Weinstein et al [40],Hartge et al [41], and Rhodes et al [42], who found that the axial gradient of thesolids concentration exhibited an S-shaped curve Horio [43] suggested that this

S-shaped curve is restricted to units with a large L/D, riser length to diameter ratio.

Large-scale units, such as atmospheric fluidized-bed combustors, may not exhibitthis axial profile Unfortunately, data on large units are limited and the solids con-centration profiles in these units are still a subject of debate

FIG 14 Axial solids concentration profile in (a) a fast-fluidized bed and (b) a riser.

FIG 15 Length-normalized pressure drop in a riser with increasing solids flux.

Recent studies have shown that the design of the entrance and exit region in

a riser can have a substantial effect on the resulting performance As illustrated

in Fig 15, the pressure drop, normalized by the distance between pressure taps,

versus the riser length to diameter ratio, L/D, can be affected by the entrance design for 20 to 30 L/D’s Furthermore, this effect becomes more pronounced with higher

solids fluxes Using an x-ray absorption technique, Kostazos et al [44] were able

to further substantiate this effect by examining the radial profile in a riser at variedfeed ports Their results showed that an asymmetric position of the feed manifested

itself in the asymmetry of the radial profile at an axial position of up to 30 L/D’s.

Radial Profile of Solids in a Fast-Fluidized Bed and Riser

The radial profile of solids that exists in fast-fluidized beds and risers is even moresurprising At some point beyond the entrance region of the fast-fluidized bed orriser, particles segregate toward the wall to form a core–annulus profile, as illus-trated in Fig 16 Studies using kinetic sampling probes, aγ-ray densitometer, andfiber optic probes were able to resolve this core–annulus profile [1,45–48] Theirresults showed that the core consists of a lean concentration of solids moving upthe riser, whereas the annulus consists of a dense concentration of particles Atmoderate solids fluxes, particles in the annulus region actually exhibit a downwardvelocity, as shown in Fig 16 [46,48,49] Karri and Knowlton were able to quantifythis downflow as a function of radial profile by measuring the solid mass fluxes

in a 20-cm-diameter by 14-m-high riser [49] Figure 17 presents their results wheredownflow in the annulus regions was observed for solid mass fluxes of 49 and 93

FIG 16 Representation of the core–annulus profile in a riser where downflow is observed near the

walls (Adapted from Ref 48.)

kg/m2s Miller and Gidaspow [50] showed that the largest magnitude of annulardownflow flux at and near the wall was near the bottom of the riser At less than

2 m from the inlet of Miller’s 7.5-cm-diameter riser, downward fluxes were severaltimes the average feed flux [51]

The implications of this behavior can be substantial For many catalytic tions, backmixing near the feed region and, to a lesser extent, up throughout theriser can have a significant impact on productivity Fortunately, many of thesereactions require very high solids fluxes where downflow may be less of an issue.For example, Fig 17 shows that operating Karri and Knowlton’s riser [49] at orabove solid fluxes of 195 kg/m2s, results in a core–annulus profile where particles

reac-at the wall move in the same direction as those in the core region (positive solidsmass flux of⬃1 kg/m2s) In this case, backmixing was limited Similar findings

FIG 17 The effects of solids mass flux on the radial net solids mass flux profile (Adapted from

Ref 49.)

FIG 18 Representation of the core–annulus profile in a riser where upflow is observed near the

walls (Adapted from Ref 49.)

have also been reported by Issangya et al [52] A representation of this behavior

is presented in Fig 18

Particle segregation appears also to be influenced by the core–annulus profile

in a riser Karri and Knowlton [49] observed that in the presence of downflow,the particle size distribution in the annulus region was larger than that found inthe core In contrast, this effect appears to only occur for downflow operations.Particle segregation was not observed for core–annulus upflow profiles for eithervery high or low solid mass fluxes [49] Jones et al [53] examined this phenomenausing, the Laser Doppler Velocimetry (LDV) of particle-laden jets Their resultsshowed that eddies or recirculation zones were responsible for this particle segrega-tion Hence, the high shear and resulting recirculation zones generated from thesolids downflow near the wall may be responsible for the segregation effect ob-served by Karri and Knowlton [49] With upflow at the wall, the low shear maynot generate strong enough eddies to effect the particle size distribution across theriser diameter

There are also design features that can reduce backmixing in risers Bafflescan induce wakes and turbulence, which limit the core–annulus profile Of course,the added attrition caused by baffles needs to be factored into the design process.Another option is to use secondary feeds to produce a higher plug flow or uniformsolids velocity profile at the entrance region A core–annulus profile may stilldevelop further up the riser, but backmixing is less severe in this region

As with the axial profile, the design of the entrance and exit region can have

a substantial effect on the solid radial concentration profile Rhodes et al [42] used

a nonisokinetic sampling probe to examine the radial solids loading in a inner diameter by 7.2-m-high riser Their results showed that a side solids feed

0.09-m-resulted in a nonuniform radial distribution of solids beyond 40 L/D’s, as depicted

in Fig 19 In addition, Rhodes et al noted that the asymmetries in solids radialdistribution were more noticeable in the interphase between the dense and diluteregions Thus, depending on the design of the feed region, a nonuniform radialprofile may exist throughout many industrial risers

In a similar fashion, the exit configuration of a riser can have an impact on

the solids profile for several L/D’s below the exit region Brereton and Grace [54]

observed this effect for smooth and abrupt riser exits As shown in Fig 20, using

FIG 19 Illustration of the nonuniform solids radial profile in a riser due to solids feed on the side

of the riser (not drawn to scale) (Adapted from Ref 28.)

FIG 20 Effect of exit configuration on solids volume fraction for a 0.15-m-diameter by 9.3-m-high

riser with a superficial gas velocity of 7.1 m/s, initial solids flux of 73 kg/m 2 s, and

148-µm sand particles (From Ref 54).

FIG 21 Solids flux ratio with respect to radial position for a riser with a smooth, rounded exit at a

superficial gas velocity of 4.2 m/s, solids flux of 50 kg/m 2 s, and 80- µm sand particles (From Ref 55.)

a smooth, wide-radius bend to terminate the riser resulted in little deviation in theaxial solids concentration profile However, an abrupt bend, such as a square bend

or tee, resulted in backmixing, which affected the overall riser solids volume

frac-tion profile up to 20 L/D’s below the exit region.

Similar effects for solids fluxes are reflected in the data of Kruse and Werther[55] who compared normalized solid fluxes to radial solids loadings for a 0.4-m-diameter by 15.1-m-high riser, as shown in Figs 21 and 22 For smooth bends,substantially less downflow is observed compared to the abrupt exit configurations

In addition, the region of downflow for the abrupt exit configuration was overtwice the size of that observed for the smooth configuration

These results provide a good example of the importance of riser design forchemical production For combustors, where backmixing is tolerable and some-times even desired, asymmetric feed designs and abrupt exits are less critical How-

FIG 22 Solids flux ratio with respect to radial position for a riser with an abrupt, squared exit at a

superficial gas velocity of 4.2 m/s, solids flux of 71 kg/m 2 s, and 80- µm sand particles (From Ref 55.)

FIG 23 Illustration of riser entrance region for a more uniform solids loading profile.

ever, in chemical production such as in oxidation and chlorination, asymmetricsolids profiles and backmixing can seriously reduce selectivity and activity Fortu-nately, both the entrance and exit can be designed such that minimal asymmetricsolids profiles and backmixing ensues For the entrance region, care needs to betaken such that entering solids are well mixed with the entraining gas One suchdesign is shown in Fig 23 A fluidizing gas is used to distribute incoming solids,and one or more jets are used to entrain catalyst into the riser Similarly, the exitregion should have either a long radius bend or a disengagement section Typically,industrial risers have a stripper section at the top of the riser to not only strip gasbut also minimize exit effects on the riser, as shown in Fig 24

Meso-Scale Behavior: Clusters and StreamersRiser sections in circulating fluidized beds exhibit a core–annulus profile withdownward flow resulting in the formation of clusters and streamers of particles

FIG 24 Illustration of riser exit region (i.e., stripper) for a more uniform solids loading profile.