winter2013.pdf

© Woodhead Publishing Limited, 2013
1005
23
Applications of fluidized bed technology in
processes other than combustion and
gasification
F. Winter and B. Schratzer, Vienna University of
Technology, Austria
DOI: 10.1533/9780857098801.5.1005
Abstract: In this chapter, the additional applications of fluidized bed
technology are discussed. Examples include uses in the environmental,
chemical, and process industries. Fluidized bed technology can be seen
to contribute significantly to the high efficiencies and low emissions of
these processes. This chapter aims to illustrate the versatility of fluidized
bed technology and to give a complete picture of the technology’s useful
characteristics, such as its excellent heat and mass transfer, good contacting
between gas and solid elements, and its utility in a wide range of operating
conditions. Examples of fluidized bed applications presented are: fluid
catalytic cracking (FCC), the reduction of iron ores, flue gas cleaning, the
production of melamine and titanium oxide, the roasting of sulfide ores, and
the drying of coal.
Key words: fluid catalytic cracking (FCC), iron ore reduction, acid gas
control, melamine, fluidized bed chlorination, sulfide ore roasting, fluidized
bed drying.
23.1 Introduction
Fluidized bed applications are widespread in the environmental, chemical
and process industries. This chapter presents applications of fluidized bed
technology which are gaining importance because of their ability to bring
high efficiency and low emissions to various processes, their excellent heat
and mass transfer characteristics, good contact between gas and solid, and
a wide range of operating conditions. This chapter intends to demonstrate
the many versatile applications of fluidized bed technology.
Examples of the additional applications of fluidized bed technology
presented are: fluid catalytic cracking (FCC), the reduction of iron ores, flue
gas cleaning, the production of melamine and titanium oxide, the roasting of
sulfide ores, and the drying of coal. A general overview of each process is
presented, with a particular emphasis on the utilization of the fluidized bed
reactor within the process, in order to demonstrate the role of fluidized bed
reactors in comparison to other types of reactors. Moreover, process-specific

1006 Fluidized bed technologies for near-zero emission combustion
details about the design and operational modes of fluidized bed reactors are
given. In summary, the primary characteristics of the selected fluidized bed
applications are described.
23.2 Petroleum refining and chemical production
23.2.1 Fluid catalytic cracking (FCC)
Since the development of the continuous fluid catalytic cracking (FCC) process
in the late 1930s in the run up to the Second World War, it has become one of
the most important processes in the petroleum refining industry, the technology
now being utilized in most oil refineries across the globe. Furthermore, it is
considered to be one of the largest applications of catalysts.
Catalytic cracking has replaced thermal cracking because of its improved
performance. This includes a far higher cracking rate at a lower temperature
and pressure level, as well as an increase in the yield of suitable products
(Hocking, 2005). The capacity of all catalytic cracking units worldwide has
been speculated to be as slightly above 14 million barrels per year (Gary
et al., 2007). Through applying hydrocarbon cracking, it has been possible
to adapt the range of natural petroleum products to the prevailing market
and to cover the increasing demand for middle distillates. It is a simple
process to increase the efficiency of the refinery by transforming high-
boiling and high-molecular components of crude oil, which are only used in
low-grade applications, resulting in more valuable products such as naphtha
and butane or pentene. However, the catalytic cracking reactions are very
complicated. During these processes, operations such as hydrocracking,
isomerization, hydrogenation, or alkylation take place, to mention just a
few. Generally, hydrocarbon cracking requires a large amount of energy
for the breaking of long-chained molecules. This process is accompanied
by the formation of residual coke. Through the application of fluidized
bed technology it is possible to provide the required energy by a solid
circulation system between a two-region assembly. The heat required for
this can be generated through combustion of the deposed residuals in the
regenerator, and it is this that is used to heat the reactor. This process has led
to FCC being the most frequently used process in cracking, because of the
advantageous utilization of the fluidized bed technique, in comparison to other
systems.
Figure 23.1 illustrates the location of the FCC unit in a conventional
petroleum refinery.
Hydrotreated vacuum gas oil, with an initial boiling point in the broad
range of 350–560°C at atmospheric pressure, is a typical feedstock (Bonifay
& Marcilly, 2001). In addition, residuals from the distillation column (FCC
fractionator) are recycled to increase the production yield. Since more

1007
Additional applications of fluidized bed technology
heavy and sour crude oils have increasingly been produced in the last few
decades, hydrotreatment has become an increasingly important issue. The
main objective of hydrotreatment is the removal of undesirable elements
and compounds including sulfur, nitrogen, oxygen, and metals. From an
environmental point of view, in particular SO2 and NOx emissions, and
because of the toxic effect of some impurities on the catalyst, it is necessary
for this separation to take place. The quality of the cracking product greatly
depends on the technology used, the character of the feed, the conditions
of the operation, and the nature of the catalyst (Bonifay & Marcilly, 2001).
Nowadays there are various designs for the FCC unit on the market, but
practically all of these designs are based on a solid circulation system. These
typically consist of the following three sections: reaction, flue gas treatment,
and product fractionation. However, the reaction section is designed as a
two-region assembly with an additional stripper. The principle of the reaction
section is illustrated in Fig. 23.2.
The regenerated catalyst circulating between both regions of the unit
provides the required energy for the endothermic cracking reaction. An
injected vaporized feed takes over the function of the fluidization agent
for this particular catalyst. By contact with the hot powdered catalyst, the
feed is almost completely converted into smaller molecules. The operating
conditions in the FCC reactor are 1–3 bar in pressure (Bonifay & Marcilly,
2001) with a temperature range of 480–550°C (Yates, 1983). After sufficient
residence time of a few seconds, the gaseous product is separated from the
catalyst by a cyclone. Subsequently the product gas leaving the reaction zone
is separated in the FCC fractionator. During the process of cleavage, the
Butanes, Pentenes
Gasoline
…
Gas (LPG)
Vacuum residue
Atmospheric
residue
Vacuum
distillation
Atmospheric
distillation
Vacuum
gas oil
Crude
oil
Hydro-
treating
FCC unit
(fluidized
bed)
Gas (LPG)
Naphtha (Gasoline)
Kerosine
Gas oil (Diesel)
23.1 Location of the fluid catalytic cracking (FCC) unit inside a typical
oil refinery.

catalyst is partially deactivated by the coke formed, which is deposited on its
surface. The spent catalyst flows into the stripper, where the removal of any
adsorbed volatiles takes place (Yates, 1983). In the following regeneration
zone, the coke deposition is burned off to provide the large demand for
heat from the endothermic cracking reactions. Using feedstock of the low
residual coke content, additional energy has to be supplied by preheating,
since the combustion of the coke does not generate sufficient heat to cover
the energy demand. Combustion of the residual coke generates emissions of
CO, CO2, SOx, NOx, and dust. These have to be removed in an additional
step for environmental reasons. In practice, the effort for flue gas treatment
varies greatly and may include recovery systems for heat and mechanical
energy to optimize the process plant. The regenerated catalyst, with a residual
coke content of 0.01–0.4 wt%, is fed back to the reaction region (Gary et al.,
2007).
Needless to say, the catalyst has a very important function inside the whole
FCC process. Modern catalysts are complex composites, commonly built up
from highly active zeolites (consisting mainly of alumina silicates), matrix
components, and additives. The mean particle size of the catalyst powder
is typically 60–70 mm with a range approximately 20–100 mm (Bonifay
& Marcilly, 2001). The catalyst is a Geldart Group A solid. Today, there
are various FCC catalysts in industrial use. One of the desirable physical
properties of the FCC catalyst is a low attrition rate, which greatly improves
particulate emission rates. The attrition resistance characterizes the catalysts’
mechanical strength. The progress in catalyst design and the associated increase
in activity and selectivity results in a decrease in the required residence time
for the cracking reactions. Since the regenerated catalyst holds the highest
temperatures at the regenerator inlet, the main occurrence of the reaction
is at the inlet region. Based on these effects, the modern FCC designs (so-
Stripper
Regenerator
Reactor
Feed
Regenerated
catalyst
Lift-gas Air
Exhaust gas
Product
Steam
Spent
catalyst
23.2 Schematic of the reaction section of the FCC unit, including
fluidized bed reactor, regenerator and the stripper, connected via the
solid circulation system of the catalyst (bed material).

1009
called riser-cracker-FCC units) have no need of the formerly-used capacious
reaction volumes. This has led, since the 1960s, to further development in
established designs of the modern riser cracker type, which is depicted in a
simplified schematic in Fig. 23.3.
In this new design, the reaction takes place exclusively in the riser
(transport line), and the earlier used reactor deals with the separation in
a smaller construction type. The reaction takes place in the vertical riser,
which is 0.8–1.3 m in diameter and 30–40 m in height (Jimenez-Garcia
et al., 2011). The superficial velocity in the riser is in the range of 15–20
m/s, with a residence time of about two seconds and temperatures which
range between 480 and 570°C (Bonifay & Marcilly, 2001). The dimension
of the regenerator is about 7 m in diameter and approximately 25 m in height
(Reichhold et al., 1999). The superficial velocity in the regenerator is about
0.3–0.8 m/s, which is much smaller than in the riser (Bonifay & Marcilly,
2001).
The temperatures in the output flow range from 650 to 815°C (Gary
et al., 2007), and the residence time of the catalyst in the regenerator is
less than 10 minutes (Bonifay & Marcilly, 2001). To avoid deactivation,
regenerator temperatures must not overheat the catalyst (Gary et al., 2007).
Feed
Air
Regenerated
catalyst
Air grid
Regenerator
Bubbling
fluidized bed
Flue gas
Product vapors
Two-stage cyclones
Stripper
Spent
catalyst
Steam
Moving
fluidized
bed
Riser
High expanded
fluidized bed
23.3 Schematic of a fluidized bed riser cracking FCC unit with internal
two-stage cyclones for solid vapor separation.

The separation system in both parts is fitted with cyclones for the final
deposition of particulate matter. The fluidization operation regime of each
process is different. In particular, the predominant process of the regenerator
is a bubbling fluidized bed. A high-expanded fluidized bed is formed in the
riser and a moving fluidized bed prevails in the reactor. Together they form
the circulating fluidized bed system.
23.2.2 Production of melamine
Melamine primarily provides a base material for resins, for example
melamine-formaldehyde (MF) and melamine-urea-formaldehyde (MUF)
resins, which belong to the group of aminoplastic resins. The most common
use of melamine-based resins is as a wood adhesive. Moreover, melamine
can be used as a fire retardant additive, for example in paints, plastics, and
paper. Today’s worldwide annual melamine production is estimated to be in
the low single-digit range of millions of tons. Melamine was first produced
by Justus Liebig in 1834 and the first commercial plants were launched in
the late 1930s (Ullmann, 1990). Since the 1960s, different processes for
trimerization of dicyandiamid (2-cyanoguanidine) to obtain melamine have
been developed. However, the utilization of this raw material has been
replaced by urea. Nowadays, melamine is only manufactured out of urea
according to the following reaction [23.1], which represents the endothermic
global reaction of the synthesis:
6(NH2)2CO Æ C3N3(NH2)3 + 6NH3 + 3CO2 DH > 0 [23.1]
In the operating pressure and the application of catalysts, two process routes
can be distinguished: the low-pressure processes with catalyst utilization,
and high-pressure processes without catalysts.
In the following section, the low-pressure process of Borealis Agrolinz
Melamine GmbH is described as an example of the application of fluidized
beds in chemical industries. Figure 23.4 illustrates a simplified flow chart of
this two-stage process, which has an output of about 20,000 tons of melamine
a year.
Internal NH3 system
NH3 NH3
Off gas
Exhaust
gas
treatment
CO2 (for urea
production)
Solid
melamine
Molten
urea
Decomposer
(fluidized
bed)
Converter
(fixed
bed)
Melamine
recovery
Purification
23.4 A schematic flow chart of a low pressure melamine production
plant including the highlighted fluidized bed area of the process.

1011
There are two parallel decomposers, which are fluidized bed reactors,
implemented using a bubble cap tray. Each fluidized bed reactor contains
several tons of silica sand as bed material with an average particle size of
approximately 800 mm in diameter. The operation of the decomposers is
through a bubbling fluidized bed. A baffle-plate on the top of the reactor is
installed and the bed material is preprocessed by rounding the sharp edges of
the silica sand. Particle elutriation due to abrasion can be kept as low as several
tons per year. The ammonia gas, which functions as a fluidization agent, and
molten urea with high purity, is fed inside the fluidized bed through a lance,
and reacts to isocyanic acid in parts by thermal decomposition (see reaction
[23.2]) at a temperature of almost 400°C and a low excess pressure:
6(NH2)2CO Æ 6HNCO + 6NH3 DH >0 [23.2]
The required heat for this endothermic partial reaction is provided by hot
molten salt through several internal heat exchangers. Figure 23.5 shows
the fluidized sand-bed reactor in more detail. The usual dimensions of such
reactors are 10 m in height and about 5 m in diameter. In operation mode,
the height of the expended bed is about 4 m.
Gaseous products leave the decomposer at the top and therefore it is very
important to hold the particle entrainment to a low level. Next, the gas flow
reaches the converter, where it is sent to one of the six stacked fixed-bed
reactors. Since for adequate conversion rates large quantities of reactants are
Baffle-plate
NH3
Isocyanic acid, NH3
(to converter)
Fluidized bed
Heat exchanger
Hot molten salt
Cold molten salt
Bubble cap tray
Urea lance
Molten urea
23.5 Principle of the urea decomposer in detail including the key
built-in components.

required, a second ammonia flow into the converter is needed. This stems
from a limited fluidization velocity in the decomposer, and an increase in
the introduced urea into the decomposer leads to a strong entrainment of bed
material, which is not desirable. The superficial velocity in the decomposer is
limited to 0.2 m/s in order of magnitude. The bulk in the converter contains
aluminum oxide 5 ¥ 10–3
m in diameter as a catalyst with a filling height
per stack of almost 1 m. The operating temperature of the bed material
of around 340°C is stratified. On contact with the catalyst, isocyanic acid
forms gaseous melamine, according to reaction [23.3], which describes an
exothermic reaction usually cooled with heat transfer oil:
6HNCO Æ C3N3(NH2)3 + 3CO2 DH < 0 [23.3]
It is also feasible to supply additional ammonia gas to reach a desirable
content of melamine. Subsequently, the gas flow, which contains mainly
ammonia and carbon dioxide, is directed to the quenching unit to separate
the melamine from gaseous components in solid state by crystallization.
Next, in the purification unit, the separated melamine is dried and stored.
Furthermore, there are two more output flows, for the carbon dioxide and
ammonia gas, which are prepared in the exhaust gas treatment unit to be
reused in the process. Nowadays, there are several other low-pressure
processes in industrial use. As previously mentioned, the alternative to these
processes are high-pressure concepts. Unlike the low-pressure method, the
high-pressure processes operate in liquid phase at temperatures of 370°C
and higher. These operational conditions are mostly characterized by better
performance of the production units. Due to the common combination of
the melamine and urea production, the provided high pressure offgas is
beneficial in increasing the efficiency by recycling the CO2 and NH3. Due to
the reaction in the liquid phase, smaller reaction units are required. However,
the advantages are generally accompanied by higher capital costs for the
high-pressure processes caused by the construction type.
23.3 Production of metals and oxides
23.3.1 Production of iron for steelmaking
There are three main process routes for the production of iron as a pre-product
for steel in industrial use. These are: the blast furnace, smelting reduction,
and direct reduction. Nowadays, iron is primarily produced through the blast
furnace process. However, this production route is characterized by several
serious disadvantages, such as the dependence on high-quality coke and iron
oxide feedstock, environmental constraints, and the requirement of auxiliary
plants (Xu & Da-qiang, 2010). The steel industry requires about 5% of the
world’s total energy consumption. Because of this enormous energy need,

1013
a significant proportion of the world’s total greenhouse gas emissions are
attributed to this industry (Xu & Da-qiang, 2010). This vast consumption
of energy has led to the development of a number of new environmentally
friendly and more sustainable technologies. The expense of high-quality
feedstock, such as lump ore and coke, as well as the high cost of processing
raw materials, has resulted in the application of fluidized bed technology.
Fluidized bed technology is chosen because of its direct utilization of low-
grade input materials such as ore fines and coal. Because of this, auxiliary
pretreatment units can be avoided, which leads to a reduction of process
emissions. The major drivers of this technology are the rising cost and the
availability of the raw material, such as coke, pellets and lump ores (Orth
et al., 2008). Moreover, the application of the fluidized bed systems represents
low investment costs and flexible plant behavior. Today, there are different
processes based on the fluidized bed technology for the production of iron in
industrial use: examples include the Finex®, HIsmelt, and Dios processes.
In the following section, the Finex® process, a development of the South
Korean steel producer Posco and Siemens VAI Metals Technologies, is
highlighted as an example of a smelting reduction process. It is an advancement
of the Corex process, which uses a shaft furnace. In the Finex® process,
the shaft furnace is replaced by various fluidized bed reactors connected in
series as a pre-reduction unit for the subsequent smelting reduction process.
Figure 23.6 illustrates a simplified flow chart of the Finex® process.
The Finex®
process includes two main sections: the fluidized bed reactor
system and the melter-gasifier. The Finex®
process is usually based on four
fluidized bed reactors in series, which operate in a continuous countercurrent
mode. For the pre-reduction of ore fines in the reaction section, a reducing
agent is required. The degree of reduction increases from the fluidized bed
reactor system inlet (fine ore) to the last reactor (fluidized bed reactor 1),
while the content proportions of the reducing compounds in the fluidization
agent decrease. Therefore, the heterogeneous or homogeneous reactions such
as the water-gas shift reaction, the Boudouard reaction, and the formation
of methane, have to be considered. The following equilibrium reactions
[23.4–23.6) have to be considered (Weiss et al., 2009). The input material
hematite is reduced via the intermediate stages of magnetite and wüstite to
metallic iron. The reduction of iron ore fines takes place in a H2, H2O, CO,
CO2, and N2 mixture atmosphere.
3Fe2O3 + H2 ´ 2Fe3O4 + H2O ΔH < 0 and/or
3Fe2O3 + CO ´ 2Fe3O4 + CO2 DH < 0 [23.4]
Fe3O4 + H2 ´ 3FeO + H2O DH > 0 and/or
Fe3O4 + CO ´ 3FeO + CO2 DH > 0 [23.5]
FeO + H2 ´ Fe + H2O DH > 0 and/or
FeO + CO ´ Fe + CO2 DH < 0 [23.6]

Nitrogen
Air
Export gas
Dedusting
Depleted reduction gas Power
plant
Air
separation
unit
CO2
CO2
removal
Oxygen
Recycle
gas
DRI
Hot DRI
compactor
HCI
Melter-
gasifier
Briquetted
coal fines
Pulverized coal
Smelting products,
slag, pig iron
Reduction gas
(fluidization agent)
Air
Additives
Iron ore
fines Fluidized
bed 4
Fluidized
bed 3
Fluidized
bed 2
Fluidized
bed 1
23.6 Simplified flow chart of the Finex®
process consisting of four
fluidized bed reactors building up the reactor system.
In the second section, the melter-gasifier, the smelting to liquid iron as
well as the generation of the reducing gas take place. This process allows
the direct utilization of fine iron ores and coal instead of sinter and coke.
The fluidized bed reactor system is fed with iron ore fines of a particle size
smaller than 8 ¥ 10–3
m (Ahn et al., 2010) and additives such as limestone
or dolomite. To heat up the raw iron ores in the last fluidized bed reactor
(fluidized bed reactor 4) in the reaction section, the required energy is achieved
by partial combustion of the reduction gas through the injection of oxygen.
A schematic diagram of one of the fluidized bed reactors is shown in Fig.
23.7.
During the residence time of about one hour in the reaction section, the
pre-reduction of the ore fines takes place (Schuster et al., 2006). The product
of the gasification of the coal functions both as a reducing gas as well as being
a fluidization agent. Reduction tests at laboratory scale have determined a
specific reduction gas consumption of 1,000 Nm3
/t ore in order of magnitude
(Plaul et al., 2008). The gas, which is used for the pre-reduction of the iron

1015
ore fines, is provided by the melter-gasifier and is rich in CO and H2 besides
CO2, CH4, and others. It is ducted in countercurrent flow with the ore fines.
The operating pressure is about 3.5 bar, the temperatures are up to 850°C,
and the superficial gas velocities range between 1 and 4 m/s (Plaul et al.,
2008). Based on the broad particle size distribution of the fine ores, ranging
from 50 mm to 8 ¥ 103
mm, there is no specific operation regime inside the
fluidized bed reactors (Plaul et al., 2008). Depending on the particle size,
the operation field ranges from the state of moving bed, bubbling bed up to
circulating fluidized bed for the finest parts of the input material. For decreasing
particle entrainment, each fluidized bed reactor includes an internal cyclone
system. The application of a multi-stage reactor system has the advantage of
the minimum requirement of reducing gas and an improvement in reduction
kinetics. The modified operation conditions in each of the four fluidized bed
reactors tend to result in the attainment of an equilibrium status followed by
better performance. To increase the utilization rates of the reduction gas, it
is partly recycled after dedusting and treatment in a CO2 separation unit via
pressure swing absorption (PSA). The remaining proportion of unutilized
gas can be used as export gas, typically for the generation of electric power,
which is needed in the air separation unit of the overall installation. The
intermediate product is direct-reduced iron (DRI) with a reduction degree of
Internal cyclones
Depleted
reduction gas
Partially reduced iron
ore (to reactor 3)
Air
Reduction gas
(from reactor 3)
Baffle-plates
Iron ore
fines and
additives
23.7 Simplified isolated view of the fluidized bed reactor 4 including
all the relevant interface information.

about 85%, which is briquetted to hot compacted iron (HCI). The DRI has
a porous structure; because of this it is often called sponge iron. The HCI
is fed into the melter-gasifier. Additionally, non-coking coal is charged into
the melter-gasifier as briquetted coal at the top and as pulverized coal plus
oxygen at the bottom. At the top of the melter-gasifier, the reducing gas is
drawn off and recycled back to the fluidized bed reaction section. The final
product of the Finex®
process is liquid iron, whose quality is comparable to
that of liquid iron produced in a blast furnace. After years of development, in
2007 Posco put the first commercial plant applying the Finex®
principle with
an annual hot metal capacity of 1.5 million t into operation in Pohang (Plaul
et al., 2008). Because of its high performance, the operators are planning to
construct another plant with a capacity of 2.0 million t/a by 2013 (Siemens
VAI Metals Technologies, 2012).
23.3.2 Fluidized roasting process of sulfide ore
Most non-ferrous metals such as copper, lead, zinc, or cobalt deposit in the
earth’s crust naturally, mainly as sulfides (Ullmann, 1996b). Zinc sulfide
ores are the feedstock used for 90% of zinc produced.
The main target of the roasting process is the transformation of the naturally
occurring metal sulfides (MeS) into its oxidic form (MeO) and the removal
of the sulfur as sulfur dioxide, according to reaction [23.7].
2MeS + 3O2 Æ 2MeO + 2SO2 DH < 0 [23.7]
The roasting process converts the sulfides into the oxide through a combustion
reaction, but no removal of the gangue material takes place. In the following
section, the application of the fluidized bed technology for roasting using
zinc sulfide is described as an example of the process. Zinc represents the
24th most common element in the Earth’s crust. Annual mine production
of zinc was about 12.4 million tons in 2011; the leading zinc producer is
China followed by Australia and Peru (US Geological Survey, n.d.). The
mining of zinc is in fourth place of metals mined worldwide. As mentioned
previously, zinc is regarded as a chalcophilic element and mainly occurs as
sphalerite, known as zinc blended ZnS. Typical zinc ores contain 10–20%
zinc, in many cases in combination with iron, lead, cadmium, manganese,
or copper as well as traces of arsenic, mercury, silver, or gold.
Figure 23.8 illustrates today’s standard route to obtain primary zinc out
of sulfidic ore by hydrometallurgical production, where the metal production
takes place in low temperature conditions. Hydrometallurgical processes
utilize the solubility and wettability of the elements. This production route
typically consists of the following sections: ore beneficiation, roasting,
leaching, separation, electrolysis and melting, and casting.
In order to increase the zinc content in mining ore from about 10% to the

1017
range of 40–60%, upstream processes for ore beneficiation are necessary.
This typically includes size reduction processes in combination with density
separation and flotation units, with the aim of removing the gangue material
such as galena and pyrite. Subsequently the zinc concentrate, which still is
present as sulfide, has to be converted into the oxidic form, since the following
steps require oxides as input material. Therefore roasting has, since the 1940s,
become an important part of the metallurgic industries. The fundamentals
are based on knowledge of fluidization which was being developed in the
petroleum industry a few years before. A large number of roasting processes
were developed in the last few decades, and today fluidized bed and sinter
roasting are frequently used in industry. Due to the fine output material, the
fluidized bed roasters are preferred for hydrometallurgical production. The
leaching requires fine-grained feedstock for better performance. In contrast,
sinter roasting is more suitable for the subsequent pyrometallurgical processes,
because of the ability of the process to obtain porous material. The fluidized
bed roaster has clear advantages over a sintering belt or a multiple hearth
roaster, such as the high throughputs at low reactor areas installed and the
great deal of flexibility the process affords according to moisture content.
Due to a low excess air rate, because of the intimate gas–solid contact, sulfur
enriches in the roaster gas to around 10%. On this account, it is viable to
generate sulfuric acid (Enghag, 2004). Moreover, the process has practical
advantages in the high availability of the roasting unit since there are no
moving parts installed. It is also easy to control and stabilize (Ullmann,
Gas
cleaning
Sulfuric
acid plant
Dust Waste
H2SO4
Waste gas
including SO2
Sulfidic zinc
concentrate
Air
Roasting
(fluidized
bed)
Leaching Separation Electrolysis
Cathodic zinc
Melting
and
casting
Zinc ingot
Residue
(to treatment)
Cu, Co, Cd
(to treatment)
23.8 A simplified block flow diagram of the location of the fluidized
bed roasting stage in the entire process.

1996b). The fluidized roasting is also higher in capacity than alternative
reactors (Kunii & Levenspiel, 1991). The Outotec company offers roasters
with a grate size up to 123 m² with an annual zinc capacity of 300,000 t.
The fluidized bed roaster is shown in Fig. 23.9 in schematic detail.
The dimensions of a fluidized bed roaster for industrial use are up to
20 m in height (Baerns et al., 2006), about 12.5 m in diameter, and the
roaster operates at gas velocities in the range of 1–3 m/s (Yates, 1983). In
the fluidized bed reactor, zinc concentrates are converted in the presence of
air to zinc oxide, which is also known as calcine as well as sulfur dioxide,
according to reaction [23.8]:
2ZnS + 3O2 Æ 2ZnO + 2SO2 DH < 0 [23.8]
This happens at a temperature range of 800–1,000°C. The characteristic
particle size of the feed material is lower than 6 ¥ 10–3
m in diameter and
air is used as fluidization agent. For the start-up of the roaster, the fitting of
an additional heating system (typically a gas-firing unit) is necessary in order
to raise the roaster temperature to the desired range, to reach the autogenous
operation mode of the process. Ignition temperature is between 550 and
650°C depending on the particle size (Winnacker-Küchler, 2006). As soon
as the reaction is started, energy in the form of heat is released. In order
to keep the temperature in the reactor at a constant optimum level, parts of
this reaction heat must be cooled down by steam generation via immersed
cooling coils. Volatile compounds of the raw material, such as cadmium,
lead, and mercury are largely enriched in the gaseous flue gas and have to
be treated in another process step. Practically all modern roasting processes
in use today are a gas cleaning unit with a SO2 recovery and subsequent
acid sulfur production plant. Roasting gas dust has to be precipitated at
temperatures above 350–400°C. This guarantees the avoidance of H2SO4
Oxidic zinc
(to leaching)
Waste gas
including SO2
Sulfidic zinc
concentrate
Steam
Water
Air
23.9 A schematic diagram of the fluidized bed roaster.

1019
condensing. The precipitation of the dust takes place in a cyclone or an
electrostatic precipitator (Baerns et al., 2006). The oxidic output material
of the roaster is ducted to the leaching section. In the subsequent separation
unit, impurities such as copper, cobalt, and cadmium are conducted. For
further electrolysis, metals less electronegative than zinc have to be removed.
Finally, the products of the electrolysis (cathodic zinc) are melted and cast
to obtain zinc ingots. Today about 80% of the total production of metallic
zinc is gained by electrowinning.
23.3.3 TiO2 pigment production by chloride process
Titanium dioxide is commonly used as white pigment and it is counted among
one of the most consumed inorganic chemicals in the world (Kirk-Othmer,
2007). Due to its low ability for absorption of visible light as well as its
beneficial light-scattering effects, TiO2 is very suitable for white pigments
(Buxbaum & Pfaff, 2005). It is used in large amounts in paper production, the
plastic industry, and in paints (Yates, 1983). In 2003, about 4.2 million tons
(Buxbaum & Pfaff, 2005) were produced, mostly TiO2 pigments and small
quantities of titanium metals. TiO2 is the ninth most abundant element in the
Earth’s crust and occurs in nature in an almost pure form, with low contents
of impurities such as iron, chromium, or vanadium. Furthermore, TiO2 is
by far the most important titanium compound (Römpp, 1999). Deposits of
TiO2 are mainly found in mineral form occurring as sand or in the form of
large hard rocks. The polymorphs of TiO2 are rutile, anatase, and brookite,
but only the first two are of industrial importance.
Nowadays there are two processes in industrial use for the production of
pigment-grade titanium dioxide. These are the wet sulfate process and the
dry chloride process. The chloride process is preferred because it is able to
produce the product in the desired particle size and in continuous operation
mode, and because fewer waste products are generated. For the latter reason, it
is an eco-friendly process which is favored despite the method’s comparatively
higher capital costs. Because of the aforementioned factors, the formerly
predominant sulfate process has become significantly less common. The
advantage of the sulfate process, however, is its ability to use a feedstock
with a higher initial content of TiO2. Based on this, the principal input
material converted in the chloride process is natural or synthetic rutile with
a content of TiO2 of higher than 90 wt% (Kirk-Othmer, 2004). Therefore,
heavy mineral sand is eligible for this process, enriching the rutile content,
since primary rocks are too low in TiO2 contents.
Recoverable reserves are estimated at 45 million tons of natural rutile.
There have been developments to move towards the production of synthetic
rutile, due to the shortage of high-grade reserves (Buxbaum & Pfaff, 2005).
Metallurgical processing of ilmenite or titanomagnetites to remove the

undesirable iron fraction produces such synthetic rutiles. In the sulfate
process, lower quality raw materials such as titaniferous slags can be used
directly. There are a few different process designs for the chlorine process
in industrial use, although the principle of each is the same. The separation
of the titanium dioxide from the host material is done by converting it to
titanium chloride, which can be removed in the following procedure more
easily and takes place through exothermic reactions [23.9] and [23.10].
TiO2 (impure) + 2Cl2 + C Æ TiCl4 + CO2 DH < 0 [23.9]
TiO2 (impure) + 2Cl2 + 2C Æ TiCl4 + 2CO DH < 0 [23.10]
In the literature different operational temperatures appear. They range between
800 and 1100°C; whereas at a higher temperature level, reaction [23.10] is
favored over [23.9], the reverse is true at lower temperatures. Coke, coal, or
other forms of carbon can act as carbon-carrier. Chlorine is commonly used
as a reduction agent. Alternative reduction or chlorinating agents such as CO,
COCl2 CCl4, or sulfur chlorides were tested, but these were not considered
further (Ullmann, 1996a). The titanium chloride is treated by oxygen to
obtain pure titanium dioxide, which is a low exothermic chemical reaction
according to reaction [23.11].
TiCl4 + O2 Æ TiO2 + 2Cl2 [23.11]
In the following section, the principle of titanium dioxide production by the
chloride process is described. The Kronos company in Leverkusen, Germany,
produces 124,000 t/a TiO2 pigment via the chloride process (Buxbaum &
Pfaff, 2005). The production contains three main steps: ore preparation, the
titanium dioxide production process, and post treatment. Figure 23.10 depicts
a schematic overview of the production route of titanium dioxide.
Waste gas
treatment
Oxygen
Solids
CO, nitrogen,
oxygen, others
Chlorine
Solids Heat
Impure
rutile
Coke
TiO2
(after
treatment)
Separation
Oxidation
Condensation
and
purification
Chlorinator
(fluidized
bed)
Recycled chlorine
23.10 Schematic flow diagram of the chloride process including the
highlighted fluidized bed area, called the chlorinator.

1021
When using naturally occurring raw material, the ore preparation commonly
begins with rutile-bearing heavy mineral sands. By various ore beneficiation
steps such as sieving, magnetic or electrostatic separation, an enrichment of
over 90 wt% is reached. Byproducts such as zirconium dioxide, as well as
ilmenite and leucoxene are gained. Subsequently, the enriched ore is ducted
into the chlorinator. The mean particle size of the input material rutile is
below 300 mm, classified as Geldart Group B (Moodley et al., 2012). Good
solid mixing characteristics, intermediate contact, no need of briquetting
of the raw material, and a continuous operation mode are all advantages
of fluidized bed applications. Hence, they are preferred over fixed bed
(Yates, 1983) and nowadays are used exclusively. Figure 23.11 illustrates a
schematic diagram of the fluidized bed chlorinator operating in a bubbling
regime.
Characteristic dimensions of the reactor are 3.7–6 m in diameter and
8–12 m in height (Lüderitz, 1984). The fluidized bed reactor is refractory-
lined with carbon steel (Yates, 1983). The bed material in use is a mixture
of feedstock and coke and the superficial velocity in the fluidized bed
reactor is about 0.15 m/s (Moodley et al., 2012). Due to their low volatility,
alkaline-earths such as magnesium and calcium chlorides accumulate in the
reactor. These impurities can cause blockages to the fluidized bed or the
equipment. In addition, zirconium silicates accumulate because they react to
their chloride form very slowly (Buxbaum & Pfaff, 2005). It is important to
use dry concentrates; otherwise, the formation of muriatic acid takes place.
By increasing the temperature during the ongoing chlorination reaction,
carbon monoxide is formed preferentially over carbon dioxide, which is a
less exothermic process. Hence, oxygen is injected into the fluidized bed
Chlorine
TiCl4, CO2 and CO
(to condensation)
Fluidized bed
of TiO2, C and
impurities
Impure rutile
(mainly TiO2)
and coke
23.11 A simplified schematic diagram of the fluidized bed chlorinator.

chlorinator, so that the temperature is kept at 800–1200°C (Winnacker-
Küchler, 2004).
The conversion rate of all input materials during chlorination is above
90%. The gaseous flow is cooled down to below 300°C (Buxbaum &
Pfaff, 2005) to separate the main part of the impurities by condensation.
By further decreasing the temperature in the following step to under 0°C,
the titanium tetrachloride condenses so that a separation of the gaseous
compounds such as carbon monoxide, carbon dioxide, oxygen, and nitrogen
is possible. The additional purification of crude TiCl4 is reached through a
simple distillation unit. In the oxidation section, the highly pure titanium
tetrachloride is converted by an excess (in a range of 110–150 w% of the
stoichiometric demand) of oxygen, from oxygen-enriched air to oxidic metal
as well as chlorine, according to reaction [23.11]. To obtain the required
temperature of up to 1400°C (Buxbaum & Pfaff, 2005), it is essential to
preheat the combustion feed. The pigment formed is removed from the gas
flow, consisting of chlorine, carbon dioxide, and oxygen, by a dry separation
unit. The gas mixture is recycled back to the fluidized bed chlorinator. In
the final post treatment section, the crude pigment is ground and washed,
followed by a final mill to obtain the desired particle size and distribution.
23.4 Coal preparation, power plants and waste
incineration
23.4.1 Fluidized bed drying
The fluidized bed drying process has been in operation since the middle
of the twentieth century. It is used in a wide range of industries. This is
because of low construction costs, easy operation, high thermal efficiency,
large capacity, and suitability for very delicate feedstocks. There are several
advantages to using steam for drying bulk material with high water contents.
The inert nature of steam disables the explosion risk of the dust, which results
in high reliability. The moisture of the raw material evaporates into the
steam, which operates as a fluidization agent. The drying process is energy
intensive. Therefore, from an economic point of view, a heat integration
concept plays an important role. By condensation of the obtained steam of
the humid raw material, a high amount of the energy demand required for
the drying process can be recovered. This procedure affects the global energy
balance in a positive way.
As an example of the technology’s practical application, the fluidized bed
drying of lignite is described. Due to the prevailing energy situation, lignite
has gained increasing importance. Lignite provides a significant proportion
of the total energy supply. Another advantage is the low mining costs of
lignite because of the possibility of open-cut mining (Aziz et al., 2011). For

1023
the treatment of low rank coal and its further utilization for combustion in
conventional steam generators, it is beneficial to integrate a pre-drying unit.
The drying of coal is beneficial due to transportation and its costs, storage,
energy density, and the increased efficiency of the power plant. Moreover,
the emissions of the power plant decrease (Aziz et al., 2011). Conventional
processes, combining milling and drying in one step, are working with
beater mills and operate at high temperatures and require large amounts
of energy (Hoehne et al., 2010). Decoupling of these two sub-processes
enables separate optimization; decoupling is applied in modern fluidized bed
drying units in combination with independent milling units. Higher thermal
efficiency and drying intensity and a more uniform temperature distribution
lead to a better temperature control of the process, through applying the
fluidized bed technology. Key benefits are that it causes less degradation
of the particles and a better gas–particle contact (Calban & Ersahan, 2003).
Further advantages of the fluidized bed drying are the large contact surface
areas between solid and gas, the good degree of solid mixing, and the uniform
temperature distribution across the bed.
The rapid transfer of heat and moisture between solid and gas shortens
the drying time. The heat and mass transfer is more intensive compared
to tube dryers (Hoehne et al., 2010) and compact designs are possible. As
already mentioned, the evaporated coal water condenses isothermally. The
high-pressure drop and non-uniform moisture contents are disadvantages,
which have to be expected (Calban & Ersahan, 2003). Besides the increase
in the adiabatic combustion temperature, the flue gas losses are decreased by
using low temperature heat for drying. On this basis, the efficiency in power
generation is raised, which results in climate protection based on reduction
of emissions due to a higher fuel utilization rate.
Until 2003, RWE Power, based in Germany, developed a modern process
for treating and drying lignite (WTA Technology®). WTA (Wirbelschicht-
Trocknung mit integrierter Abwärmenutzung) is a German abbreviation
which stands for fluidized bed drying with internal waste heat utilization.
Independent of the subsequent utilization of the dried lignite (combustion,
gasification, coking of lignite) there are small differences in the design of
the drying unit. A fine grain dryer with an integrated vapor condenser for
preheating of boiler feedwater is depicted in Fig. 23.12. This principle is
used at the coal dust-fired 1,000 MWe BoA power plant in Niederaußem,
Germany, with raw lignite inputs of 210 t/h. BoA (Braunkohlekraftwerk
mit optimierter Anlagentechnik) is the German abbreviation for lignite-fired
power plant with optimized installation engineering.
The raw lignite has a characteristic particle size of below 80 mm. Lumpy
raw lignite has a high water content of approx. 50–60% (mass fraction),
which reduces the lignite’s calorific value. This raw lignite is ground, for
example with a double rotor mill, to a particle size of about 2 mm and

smaller. Subsequently the milled wet lignite is fed via a rotary star valve
into the dryer unit, which is schematically illustrated in Fig. 23.13.
The dryer unit is arranged as a fluidized bed with low expansion. In
this design, auxiliary steam is required for covering the energy demand of
the drying unit. For evaporation, a tubular heat exchanger is used with a
steam pressure of 3–4 bar, which means that the auxiliary steam is slightly
Cooling Milling
Milling
Raw lignite
Dry lignite
Steam
Dryer
(fluidized
bed)
Recycled lignite
Water
Auxiliary steam
Dedusted steam
Water
Water
Vapor
conden-
sation
Precipitator
Dust
Preheated boiler
feedwater
23.12 Principle of the lignite drying unit with vapor condensation
including the highlighted fluidized bed drying unit.
Dried lignite
Water
Auxiliary steam
Steam
Steam
Spreader
Raw lignite
Low expanded fluidized bed
Tubular heat exchanger
Steam for fluidization
23.13 A schematic diagram of the fluidized bed dryer unit in detail
including all of the relevant built-in components.

1025
overheated. In the fluidized bed reactor the temperature is about 110°C, with
a system pressure of about 1.1 bar, in order that residual moisture contents of
approximately 12% can be obtained (RWE Power Aktiengesellschaft, 2009).
The milled raw lignite is fed and uniformly distributed to the drying unit
via a spreader, a specially designed feeding hopper. Since the fluidization
of the milled raw lignite is difficult to produce, dried lignite is recycled to
avoid problems resulting from the cohesive character of the humid lignite.
While the lignite passes downwards through the reactor, the fluidization agent
is flowing countercurrent. The mixture of the fluidization agent (recycled
dedusted steam) and the generated steam is removed at the top of the dryer.
A sufficient part of this mixture is recycled and utilized as a fluidization
agent after the electrostatic precipitator (ESP). The remaining steam, which
was generated in the drying unit, is utilized for boiler feedwater heating.
The condensing steam can further be used as process water to decrease the
utility streams, as well as for minimization of waste streams.
In operation, the fluidized bed of the drying unit is about 3.5 m in height,
and the height of the dryer is approximately 10 m. The average fluidizing
velocity depends on particle size and for fine grain application it is about
0.14 m/s (Klutz et al., 2010). After a set drying time, the deepening lignite
flows under the fluidization level and is ground off via a rotary star valve.
Subsequently, the dried lignite is cooled down and passes through a secondary
milling unit to obtain the desired particle size depending on its intended
application, which is combustion in this case. RWE have developed alternative
designs for vapor utilization. Besides the application mentioned previously,
other concepts of energy integration are possible. To avoid auxiliary steam
for heating, it may be feasible to utilize parts of the vapor flow generated
in the drying unit after mechanical vapor recompression.
23.4.2 Flue gas cleaning
This section covers the cleaning of flue gas in power stations and waste
incineration plants. Three flue gas cleaning devices can be identified (dry,
semi-dry and wet processes) with varying characteristics. A big advantage
of the dry flue gas cleaning device in comparison to the alternative process
of wet scrubber is the avoidance of liquid output streams, which would need
additional treatment. Moreover, dry systems are often beneficial because
of their reduced capital costs, the lower energy demand during operation,
and the comparatively simple construction (Theodore, 2008). Due to the
ever-tightening emission restrictions over the last few decades, the flue gas
cleaning systems of the combustion processes have in many cases expanded
to a multitude of units. This development can be credited with high separation
efficiency but also high expense and complexity. Driven by increasing
economic pressure in recent years, an effort towards optimization of the cost

efficiency of waste incineration has been raised, because of the liberalization
of the electricity market in Europe. This context has led to advances in the
established cleaning technologies developed during the 1990s. Several of
these processes now combine the benefits of dry cleaning devices and the
advantageous properties of fluid bed technology. These techniques are used
in the removal of acid gases, heavy metals, dioxins, furans, and particulates,
the NID (novel integrated desulphurization) process, Circoclean®
process,
and the Turbosorp®
process for industrial uses.
This section describes the principles of acid gas cleaning with applied
fluidized bed technology based on the Turbosorp® process. This is a dry or
semi-dry process for flue gas cleaning. The chemical reaction equation of
the semi-dry operation mode is given in the following reaction:
Ca(OH)2 + SO2 Æ CaSO3 + H2O [23.12]
The reaction of calcium hydroxide and sulfur dioxide has a strong dependency
on humid conditions. The removal of hydrochloric acid from flue gas is
accomplished according to:
Ca(OH)2 + 2HCl Æ CaCl2 + 2H2O [23.13]
The presence of carbon dioxide, which is always the case in combustion
processes, leads to the formation of calcium carbonate, then to calcium
sulphite:
Ca(OH)2 + CO2 Æ CaCO3 + H2O [23.14]
CaCO3 + SO2 Æ CaSO3 + CO2 [23.15]
In an atmosphere containing oxygen, calcium sulfate is formed, according
to:
CaSO3 + 1
2 O2 Æ CaSO4 [23.16]
The Turbosorp® process has become an important downstream unit in
several parts of the field of combustion. The process is applied as a flue gas
desulfurization unit and in flue gas cleaning after biomass boilers or waste
incinerators. The resulting ratio of HCl and sulfur dioxide varies depending
on the fuel and flue gas produced. Therefore, whilst the process differs in
design, the main principle is the same. Figure 23.14 illustrates a simplified
flow diagram of the Turbosorp® process. An example of the Turbosorp®
process is the waste incineration plant in Zorbau (near Leipzig, Germany),
which has been in operation since 2005, and which uses this design. Each
of the two trains provides approximately 54 MW of thermal energy with a
total capacity of municipal solid waste and commercial waste of 300,000
tons per year (Kedrowski et al., 2010).
Depending on the constituents which have to be removed from the flue

1027
gas, the Turbosorp®
process may also be used in combination with another
cleaning unit. There is also the opportunity to combine a Turbosorp®
system
with a wet scrubber arranged downstream. This might be advantageous to
extend the range of species that can be precipitated and to increase separation
efficiencies. To avoid wastewater, the contaminated water produced is
recycled to the Turbosorp®
unit (Brunner, 2009). The interconnection of a
Turbosorp®
process with a conventional wet scrubber uses the benefits of
both processes. Due to the recirculation of the wastewater of the scrubber
to the Turbosorp®
reactor, no liquid output stream is produced. Generally,
the Turbosorp®
flue gas cleaning system consists of the Turbosorp reactor,
a separation unit, a recirculation conveyor, and a water injection system.
The combination of these units and the additional utilization of the flue gas
itself as a fluidization agent forms the circulating fluidized bed system. The
cleaning system is illustrated in Fig. 23.15.
The flue gas from the combustion (raw gas) is fed directly into the lower
part of the reactor section. The inlet is formed by a venturi nozzle to ensure
the correct gas velocities, in order to avoid a collapse of the fluidized bed.
Water and the recycled flue gas cleaning products from the separation unit
enter the reaction section close to the nozzle. Additionally, there is a further
input of fresh powdered additives, of less than 50 mm diameter (Mickal,
2001), to increase the separation efficiency and to balance the quantities
of the withdrawal of bed material. The applied particles can be classified
as Geldart Group C particles. The dimension of an industrial Turbosorp®
reactor is about 7.5 m in diameter and up to 21 m in height (Winter et al.,
1999). Through the injection of water, the inlet temperature, which ranges up
Product
Sorbent
preparation
unit
CaO
H2O Ca(OH)
2
Solid
recirculation
Circulating fluidized bed system
H2O
NH4OH
Fuel
Air
Combustion
and SNCR
Turbosorp
reactor
Separator Clean gas
23.14 Simplified flow diagram of the semi-dry Turbosorp®
process
including the highlighted circulation fluidized bed system.

to 220°C (Spiess-Knafl, 2000), decreases to approximately 140°C (Abrams
et al., 2010). It is critical to use a water spray to increase the absorption
capacity of desulfurization by raising the relative humidity. A mixture of
calcium hydroxide, calcium carbonate, solid products of the reaction of the
flue gas cleaning system, and ash from the combustion process is used as
bed material. Based on agglomeration during the flue gas cleaning process,
the fine particles grow to a maximum size of approximately 1 cm (Winter
et al., 1999).
The operational state of this process is a fast fluidized bed close to pneumatic
transport with superficial gas velocities of between 4 and 6 m/s (Kedrowski
et al., 2010). The characteristics of the circulating fluidized bed technology,
especially the close contact of the reacting phases and high mass transfer
rates of the bed, are among the benefits of this concept. In particular, the
stress on the bed material is in comparison to other desirable applications
(Reissner et al., 2003). The close contact of the solids and the flue gas, in a
combination of high recirculation ratios of about 99% of the separated bed
material, yields a high utilization rate of the used sorbent materials. By this
method, the generated solid output product decreases; this is desirable from
an environmental and economic point of view.
The output product can be applied in different ways. In the case of a flue
gas desulfurization, the product can be dumped into a landfill without further
treatment. Products accruing from flue gas treatment after biomass boilers or
waste incinerations cannot be dumped into landfill without further treatment,
since heavy metals can leach out (Reissner et al., 2003). After a residence
time of approximately 2–4 s (Kedrowski et al., 2010) in the reactor section,
Product
Clean gas
Solid recirculation system
Reactor
Water
Raw gas
Pre-separator
Fabric baghouse filter
Fast fluidization bed
Calcium
hydroxide
23.15 A simplified schematic diagram of a Turbosorp®
reactor in
combination with a fabric baghouse filter.

1029
the gas–solid mixture gets into the separation unit. This usually consists of
a mechanical pre-separator (such as a cyclone or a system of deflectors) and
a fabric baghouse filter or an electrostatic precipitator. In desulfurization,
either a fabric filter or an electrostatic precipitator can be used as the main
separation unit. For the flue gas cleaning of waste incineration, only a
combination of pre-separator and fabric baghouse filter can be employed.
For the reduction of heavy metals, dioxins, and furans, activated carbon or
lignite coke can be added.
The good contact of the solid–gaseous system in the fabric baghouse
filter causes additional reactions to occur in the filter cake, which decrease
pollutants in the output flow. The recirculation of the removed solids from
the separation unit takes place pneumatically (in a fluidized conveyor) or
mechanically (in a screw conveyor). The availability of Turbosorp®
units is
very high and the system is able to regulate the process temperature, pressure,
and emissions (Kedrowski et al., 2010). A comparison of the experiences
of the Turbosorp®
process and conventional spray absorption units has
illustrated that no caking occurs. Due to the high dust load in the reaction
chamber, frequent cleaning of the walls takes place. No slurry systems are
needed because, as previously mentioned, no wastewater is produced. The
low residence times also mean that reactor proportions are smaller (Abrams
et al., 2010).
23.5 Conclusion
This chapter has presented applications of fluidized bed technology in the
environmental, chemical, and process industries. It has shown the versatility
of this technology, which is characterized by excellent heat and mass transfer,
good contact between gas and solid materials, and a wide range of operating
conditions. As examples, fluid catalytic cracking (FCC), the reduction of
iron ores, flue gas cleaning, the production of melamine and titanium oxide,
the roasting of sulfide ores, and the drying of coal are presented. Process
temperatures may vary from low temperatures up to more than 1,000°C and a
wide range of materials can be handled (see Table 23.1 for more details).
In each section, a specific process was described. After a general picture
of the process, typical details about the design of the fluidized bed reactors
specific to that operation were given. These are summarized in Table 23.1.
These details are process-specific and show the variety of engineering
solutions, which include internal cyclones, combinations of fluidized bed
reactors and in-bed heat exchangers. Reducing and oxidizing or highly
corrosive atmospheres are used.

©
Woodhead
Publishing
Limited,
2013
Selected Fluidized Bed Applications
FCC Riser FCC Regenerator Finex®
Reduction Unit Turbosorp®
Riser
Geldart classification A B C
Bed material Catalyst, mainly active zeolites Iron ore fines Ca(OH)2, CaSO3 and other
reaction products
Diameter (particles) 20–100 µm 50 µm–8 mm Feed < 50 µm
Agglomerates 1–10 mm
Superficial velocity (m/s) 15–20 0.3–0.8 1–4 4–6
Temperature range (°C) 480–550 650–815 up to 850 140–220
Type Highly expanded FB Bubbling FB Bubbling FB CFB
Reactor unit height (m) 50 25 n.a. 21
Reactor unit diameter (m) 0.8–1.3 7 n.a. 7.5
Residence time range a few seconds a few minutes about an hour a few seconds
Selected Fluidized Bed Applications
Melamine decomposer Chlorination Roasting Lignite drying
Geldart classification B B B B
Bed material Silica sand Impure rutile and coke Sulfidic zinc concentrates Raw lignite
Diameter (particles) ≈ 800 µm < 300 µm < 6 mm < 2 mm
Superficial velocity (m/s) 0.2 0.15 1–3 0.14
Temperature range (°C) 400 800–1200 800–1000 110
Type Bubbling FB Bubbling FB Bubbling FB Bubbling FB
Reactor unit height (m) 10 8–12 20 10
Reactor unit diameter (m) 5 3.7–6 12.5 n.a.
Residence time range a few seconds minutes (estimation) minutes (estimation) minutes (estimation)
Table 23.1 Summary of the typical values of the primary characteristics of selected fluidized bed applications

1031
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on Flame Temperature and Flooding Limits in the FINEX Process: A Theoretical
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Aziz, M., Kansha, Y. & Tsutsumi, A., 2011. Self-heat Recuperative Fluidized Bed Drying
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Baerns, M., Behr, A., Brehm, A., Gmehling, J., Hofmann, H., Onken, U. & Renken, A.
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Enghag, P., 2004. Encyclopedia of the Elements. Weinheim: Wiley-VCH Verlag.
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23.7 Appendix: abbreviations
BoA Braunkohlekraftwerk mit optimierter Anlagentechnik
CFB circulating fluidized bed
DRI direct-reduced iron
ESP electrostatic precipitator
FB fluidized bed
FCC fluid catalytic cracking
HCI hot compacted iron
LPG liquefied petroleum gas
MF melamine-formaldehyde
MUF melamine-urea-formaldehyde
PSA pressure swing absorption
SNCR selective non-catalytic reduction
WTA Wirbelschicht-Trocknung mit integrierter Abwärmenutzung
DH enthalpy of formation

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