WO1989005282A1

WO1989005282A1 - Cyclical reductive and oxidative decomposition of calcium sulfate in two-stage fluidized bed reactor

Info

Publication number: WO1989005282A1
Application number: PCT/US1988/004360
Authority: WO
Inventors: Thomas D. Wheelock
Original assignee: Iowa State University Research Foundation, Inc.
Priority date: 1987-12-07
Filing date: 1988-12-06
Publication date: 1989-06-15
Also published as: AU2826589A

Abstract

Calcium sulfate is converted to calcium oxide and sulfur dioxide in a reactor providing two superposed fluidized beds arranged so that the fluidizing gas from the lower bed passes through the upper bed, and the process includes the steps of preheating the feed in the upper fluidized bed, passing the feed from the upper to the lower bed, and converting it in the lower bed by cyclically subjecting the lower bed to reducing and oxidizing conditions.

Description

CYCLICAL REDUCTIVE AND OXIDATIVE DECOMPOSITION OF CALCIUM SULFATE IN TWO-STAGE FLUIDIZED BED REACTOR

FIELD OF INVENTION

The field of this invention is processes and appara¬ tus for decomposition of calcium sulfate to calcium oxide and sulfur dioxide. The invention is particularly concerned with the operation of fluidized beds for this purpose using a combina¬ tion of reductive and oxidizing conditions.

BACKGROUND OF INVENTION

Prior patents relating to processes and apparatus^* for decomposition of calcium sulfate (CaS0₄) include the Whee- lock and Boyland U.S. Patents 3,087,790 and 3,607,045, and Whee- lock Patent 4,102,989. These patents illustrate the prior art of using fluidized beds for conversion of CaS0₄ to CaO and S0₂. The Wheelock and Boylen patents disclose reaction condi¬ tions by which the fluidized bed is maintained under carefully controlled reducing conditions for the entire conversion.

The Wheelock patent describes an alternative pro¬ cess in which the lower portion of the fluidized bed is a reduc¬ ing zone and the upper portion is an oxidizing zone providing a sequence of reducing and oxidizing conditions. Stoiσhiometric excess of fuel gas over the primary combustion air in the lower zone produces the reducing CO or CO and H₂. Secondary air is introduced into the upper portion of the bed in stoichiometric oxygen excess for completing the combustion of the CO and H₂ and leaving unreacted 0₂. This two-zone combination reduced CaS contamination of the CaO product. Substantially all of the sulfur in the feed is discharged as S0₂ in the off-gas. While the process of Wheelock patent 4,102,989 achieves excellent results with respect to the conversion of the calcium sulfate and the production of a pure CaO product, commer¬ cial utilization requires use of a relatively high cost reac¬ tor. Providing for the introduction of secondary air, especial¬ ly in large diameter fluidized beds, adds appreciably to the equipment cost. Alternatives to the two-zone design have there¬ fore been investigated.

As discussed in the introductory portion of Whee¬ lock Patent 4,102,989, in certain experiments of Walter M. Boy- land there may have been an incidental fluctuation between oxi¬ dizing and reducing conditions during the conversion of calcium sulfate in a batch-type fluidized bed reactor. The cause of the apparent fluctuation would have been due to incidental or uncon¬ trolled fluctuation in either the air. ξlow rate or natural gas ' flow rate.

Subsequently, Dr. T. D. Wheelock investigated the decomposition of another form of calcium sulfate known as phospho-gypsum. The results of that investigation were summa¬ rized in a Report of Engineering Research Institute, Iowa State University, Ames, Iowa, prepared by T. D. Wheelock (ISU-ERI-AMES-83412, August, 1982). In one series of laboratory tests a combustion tube furnace was employed with a stream of gas flowing over the phosphogypsum. The gas- stream was made al¬ ternately reducing and oxidizing. Temperatures of 1150 to 1200 C were used with reaction times of from 30 to 90 minutes. The data obtained indicated that such alternating conditions might be feasible, although the laboratory tests did not simulate a fluidized bed reactor. More recently. Dr. T. D. Wheelock and Dr. C. W. Fan studied the disposal of wastewater treatment sludges by inciner¬ ation. A major component of the sludges was gypsum crystals (CaS0₄ * 2H₂0) , which were decomposed to CaO by reductive calcination. It was therefore hoped to convert the sludge to a useful material. (See the Report of Engineering Research Institute, Iowa State University, Project 1548, ISU-ERI-AMES-86401, July 31, 1985).

The sludge treatment experiments were carried out in bench-scale and pilot plant size fluidized bed reactors, using natural gas and air. In most of the tests the fluidize beds were operated continuously using a two-zone fluidized bed reactor. However, in one series of runs, the entire fluidized bed was made alternately oxidizing and .reducing by turning the natural gas on and off in a periodic manner while the air flow continued. Cycle times of the reducing and oxidizing sequence were varied from 10 seconds to 100 seconds.

The results obtained failed to establish that this alternating mode of operation was preferable to the established reducing mode. If the cycle time was reduced from 100 seconds to 10 seconds, resulfurization of the CaO product did decrease slightly (from 98 to 95%). At the same time, however, the sulfide content of particles entrained in the product gas was adversely affected. The report indicates that the higher sulfide content of the free particles entrained in the gas during the cyclic mode of operation was probably due to the lack of oxygen in the off-gas during the reducing phase of each cycle_t that is, particles entrained during the reducing phase were not as well oxidized as those entrained during the oxidizing phase. Further, the total recovery of the CaO product (CaO) was somewhat lower for the cyclic mode of operation than for the standard mode, and the recovery tended to decrease when the cycle time was shortened. It was also found that with the cyclic mode of operation the temperature of the fluidized bed oscillated by as much as 12°C. The alternating mode of operation therefore remained a theoretical possibility whose practical value was far from established.

SUMMARY OF INVENTION

This present invention is believed to represent the first commercially practical process for decomposition of calci¬ um sulfate in a fluidized bed using cyclical reducing and oxidiz¬ ing conditions. For carrying out the process, a reactor is used which provides two superposed fluidized beds arranged so that the fluidizing gas from^'the lower bed passes through the^'upper bed. The CaS0₄ feed is introduced into the upper bed for preheating by contact with the gas from the lower bed. The preheated feed from the upper bed is passed to the lower bed and converted therein to CaO and S0₂ products by cyclically sub¬ jecting the lower bed to reducing and oxidizing conditions while maintaining a temperature effective for the conversion to these products. The cyclical conditions in the lower bed are produced by continuously introducing preheated air beneath the lower bed as the primary fluidizing gas. Increments of a hydrocarbon fuel, preferably a gas phase fuel, are discontinuously intro¬ duced into the primary fluidizing gas for combustion therein. The fuel increments are in stoichiometric excess of the oxygen available for their complete combustion so that pulses of reduc¬ ing gas are formed which are separated by pulses of oxidizing gas. By this process, very high conversions of the calci¬ um sulfate to CaO and S0₂ can be obtained, the CaO product is essentially free of sulfide (CaS) , and the sulfur is substantial¬ ly all in the off-gas as S0₂. By preheating the calcium sul¬ fate feed and also preferably preheating the fluidizing combus¬ tion air, process heat can be highly conserved and temperature fluctuations in the cyclical reaction bed can be minimized.

THE DRAWINGS

The accompanying drawings are illustrative of the process of this invention.

FIG. 1 is a diagrammatic elevational view of a two stage fluidized bed reactor adapted for carrying out the pro¬ cess; and

FIG. 2 is a schematic flowsheet showing how the two stage fluidized bed of FIG. 1 can be integrated with a .complete system.

DETAILED DESCRIPTION

For use in the method of this invention, the calci¬ um sulfate may be in anhydrous or hydrated form. Anhydrite mineral is composed principally of anhydrous calcium sulfate (CaS0₄) , which is usually present in admixture with a minor proportion of hydrated calcium sulfate known as gypsum (CaS0₄ • 2H₂0) . Gypsum ore can also be utilized, or gypsum ore mixed with the anhydrite mineral. Industrial waste materials composed mainly of calcium sulfate, either hydrated or anhydrous, can also be utilized. Unless the calcium sulfate is already in a state of fine subdivision, it is prepared for use in the process by crushing, grinding, and screening to produce a finely-divided feed of relatively uniform mesh size, such as -6 to +65 mesh (Tyler Standard Screen). If the particle size is too fine for fluidized bed use, the particles can be aggregated to produce a material suitable for treatment in a fluidized bed.

The conversion of calcium sulfate to CaO (quick¬ lime) and S0₂ (sulfur dioxide) in the presence of hydrogen (H₂), carbon monoxide (CO), and oxygen (0₂), can involve at least six different reactions. These reactions are shown below in equations 1 to 6.

(1) CaS0₄ + CP ^•*^■ CaO + C0₂ + S0₂

(2) CaS0₄ + Hi * CaO + H₂0 + S0₂

(3) CaS0₄ + 4 CO -* CaS + 4C0₂

(4) CaS0₄ + 4H₂ CaS + 4H₂0

(5) CaS + 3/20₂ ^■*" CaO + S0₂

(6) CaS + 20₂ ^÷ CaS0₄

Equations (1) and (2) represent the desired product producing reactions. These reactions are favored by mildly reducing conditions. More highly reducing conditions tend to favor reactions (3) and (4) . However, it is necessary to have a sufficient concentration of the reducing gas (CO and/or H₂) to provide the driving force for the decomposition reactions. It is therefore difficult to avoid some production of calcium sul¬ fide by reactions (3) or (4). However, the CaS can be reconvert¬ ed to CaO by reactions (5) and (6) . Reaction (5) is preferred to reaction (6), since reaction (δ) reforms calcium sulfate, which can be entrained as fine particles in the off-gas, result¬ ing in CaO product loss and reducing the yield of the S0₂ by-product.

Since all of the above-described reactions will occur to some extent in the process involving alternating or cyclic reducing and oxidizing conditions, it is necessary to provide a balance of process conditions which in the reducing phase maximizes reactions (1) and (2), and in the oxidizing phase favors reaction (5). It is believed that the process of the present invention provides such cyclical process, which not ' only achieves the desired reactions but which also controls the temperature of the fluidized bed in which the reactions are occurring, and which achieves a high conservation of process heat.

In practicing the method of this invention, preheat¬ ed air is used as the customary fluidizing gas and also func¬ tions as the gas for the combustion of the fuel. Any gas phase hydrocarbon fuel can be employed, such as natural gas, methane, propane, etc. Normally liquid hydrocarbon fuels can be vapor¬ ized to provide a gas phase fuel for combining with the fluidiz¬ ing air. Solid carbonation fuels can also be used, such as powdered coal or coke, but are not preferred. Precise control of the reducing and oxidizing cycles can be better obtained with gaseous hydrocarbon fuels.

The fluidizing combustion air is delivered to the reactor at a velocity or flow rate sufficient to achieve full fluidization of the reactor beds. The required procedures for accomplishing this are well known in the fluidized bed art and need not be described herein.

In practicing the method of this invention, preheat¬ ed air, which is delivered at a fluidizing velocity, is introduced continuously beneath the lower bed of the reactor. Increments of the gas phase hydrocarbon fuel, for example, are discontinuously introduced into the fluidizing air for combus¬ tion thereby. The fuel increments are in stoichiometric excess of the oxygen available for their combustion. The cyclical introduction of fuel increments thereby generates pulses of reducing gas separated by pulses of oxidizing gas.

A complete cycle is defined as consisting of one sequential reducing pulse and an oxidizing pulse. The cycle should be carried out in a very short time interval. A pre¬ ferred cycle time is of the order of from 40 to 100 seconds. It is important to control the length of the reducing phase of the cycle in relation to the oxidizing phase. Fuel increments are preferably introduced from 70 to 90% of the cycle.

The division of the cycle into reducing and oxidiz¬ ing phases depends on the fuel and air rates required to satisfy the material and energy balances as well as the air to fuel ratio chosen for the reducing phase of the cycle. It has been found experimentally that an air to methane fuel ratio during the reducing phase should be in the range of 4 to 7 and prefera¬ bly in the range of 5 to 6. Generally it is better to operate a fluidized bed reactor with a constant gas flow rate to minimize entrainment of dust in the off-gas. Therefore it is proposed here to increase the air rate during the oxidizing phase and decrease the air rate during the reducing phase in order to maintain a constant total gas flow rate. Taking all of these factors into consideration, the fraction of the operating cycle devoted to the reducing phase can be calculated by using the following relation:

t = (R + DM θ A & M

where t = reducing phase time, Θ = oxidizing phase time, A = average mole air/mole CaS0₄ fed, M = average mole CH₄mole CaS0₄ fed, and R = mole air/mole CH₄ during reducing phase.

For use in the above relation, average values for the cycle of A and M are determined by material and energy bal¬ ance calculations. Once t/Θ has been calculated, it can be used to find the air and fuel rates for the reducing phase of the cycle. Thus for the reducing phase the following quantities will apply:

mole air = A + M - M mole CaS0₄ t/θ

mole CH₄ = _M_ mole CaS0₄ t/θ

Furthermore, for the oxidizing phase the following expression gives the appropriate air rate: mole air = A + M mole CaSO,

For the preferred range of the air to methane fuel •ratio the fraction of cycle time devoted to the reducing phase will be in the range of about 50 to 60% for a single-stage fluid¬ ized bed reactor system without heat recovery whereas it will be in the range of about 70 to 80% for a two-stage fluidized bed system with heat recovery.

In achieving the benefits of the process of the present invention, it is important to have it carried out in a reactor providing two superposed fluidized beds arranged so that the fluidizing gas from the lower bed passes through the upper bed. A reactor of the kind required is diagrammatically illus¬ trated in FIG. 1. The upper fluidized bed is a solids preheating bed, while the lower fluidized bed is the reaction bed in which the conversion reactions are carried out.

As illustrated in FIG. 1, CaS0₄ feed is introduced into the upper fluidized bed and is preheated therein by contact with the gas passing from the lower bed through the gas diffuser to the upper bed. For the purpose of the present invention, the temperature of the upper bed should be controlled to achieve a relatively high preheating temperature. In general, it is desired to preheat the feed in the upper bed to a temperature in the range from 1100 to 1600°F. This temperature can be achieved since the conversion temperature maintained in the lower bed will be in excess of 2000°F, and a preferred reaction temperature range being from 2000 to 2100°F. The off-gas from the lower bed can therefore be at a temperature above 2000°F. The gas discharged from the upper bed will still contain considerable sensible heat, and may have a temperature in the range from about 1100 to 1600°F. This off-gas can therefore be used for preheating of the fluidizing air. After preheating in the upper bed, the feed is passed by a solids downcomer through an automatic flow control valve into the lower bed, as indicated in FIG. 1. The lower bed is subjected to the cyclical conditions described above. This can be accomplished by introducing pressurized air through an automatic control valve at a predetermined flow rate and volume, the air passing into the bottom^" of the reactor where it is dis¬ persed by a gas diffuser which provides a fluidizing low through the solid particles in the lower bed. The gaseous hydro¬ carbon fuel, such as methane (CH₄) is introduced on a discon¬ tinuous basis through an automatic valve, such as an on-and-off solenoid valve. For example, as shown, the methane fuel can be injected into the incoming air and combined therewith to provide a series of plug-flow pulses. The fuel and the relative propor¬ tions of fuel to the preheated air should be selected to avoid an explosive condition. Preferably, the fuel is burned in the lower fluidizing bed.

In steady state operation, the CaS0₄ feed may be continuously introduced into the upper bed and continuously transferred at a controlled rate to the lower bed, and the CaO product may be continuously removed from the upper portion of the lower bed. The incoming air provides the principal fluidiz¬ ing gas and also serves as the combustion air. This fluidizing combustion air is preferably preheated. This can be done by indirect heat exchange with the gas from the upper bed, as illus¬ trated in FIG. 2. Off-gas from the upper bed can be passed through a cyclone separator where dust particles (solids) are removed. The dust can be recycled by combining it with the feed before it is introduced into the upper bed. After removal of the dust in the cyclone separator, the product gas containing the S0₂ can be passed through an indirect heat exchanger in heat exchange relation with the air supply, which may be at ambient temperature. This air is preferably preheated to a temperature in the range of 1000 to 1500°F. DESIGN EXAMPLES

For the design study, four cases were considered based on feed with different amounts of water of crystallization and two types of fuel, methane and petroleum coke. The design basis for calculating the fuel and air requirements is shown in Table 1, and the fuel and air requirements of the different cases are indicated in Table 2, as well as the sulfur dioxide content of the off-gas. By feeding calcium sulfate hemihydrate it would be possible to produce a product gas stream containing 11-12% sulfur dioxide after drying. Feeding calcium sulfate dihydrate would increase the fuel and air requirements by 11% and 15%, respectively, and reduce the sulfur dioxide concentra¬ tion of the dried product gas stream by 11%.

Table 1. Design basis for calculating fuel and air requirements

Parameter value

Reactor feed 100% CaS0₄-l/2 H₂0 or 100% CaS0₄*2 H₂0

Reactor fuel 100% CH. (natural gas) or 100% C (petroleum coke)

Ambient temperature 77^«F

Reaction temperature 2100^»F

CaSO. desulfurization 98%

CaSO. recycle 5%

Excess air 5%

Feed particle size -8/+40 mesh

Overall gas to gas heat transfer coef. 10 ^Btu/hr. ^βF ft²

Ambient temperature 77^βF

Heat loss from lower part of reactor 400 ^Btu hr ft²

Heat loss from upper bed of reactor 200 ^Btu hr ft² 13

Table 2. Fuel and air re uirements for convertin

CaS0₄ ^*l/2 H₂0 CH₄ 0.99 7.48 9.4 12.3 CaS0₄-2 H₂0 CH₄ 1.11 8.61 7.4 10.9 CaS0₄-l/2 H₂0 Coke 1.94 7.24 10.6 11.3

Caso₄*2 H₂o Coke 2.17 8.38 8.3 10.0

Table A shows conditions appropriate for cyclic operation of the two stage system represented by FIG. 2 using methane fuels and either calcium sulfate hemihydrate or calcium sulfate dihydrate feed. These conditions are based on the fuel and air requirements shown in Table 2 and the design basis shown in Table 1. Table A shows the air and CH₄ rates in moles/sec required to maintain a constant total gas flow and the fraction of the cycle devoted to the reducing and oxidizing phases, re¬ spectively, for different air to fuel ratios. It also shows the concentration of CH₄ in the gas mixture and the temperature of the preheated air stream (I) and the temperature of the air/CH₄ mixture stream (K). The concentration of CH₄ in the mixture is outside the rate of explosive limits. For example, Lange's Handbook of Chemistry and Physics, Sixth Edition, lists the lower and upper explosive limits of methane as 5.3% and 13.0% respectively. These values represent volume percent in air. Tables B and C show the relative magnitude of the streams identified in FIG. 2 for the cases where calcium sulfate hemihydrate and calcium sulfate dihydrate are used, respective¬ ly, together with methane fuel. Indicated values represent moles per mole of CaS0₄ fed. Also shown are calculated stream temperatures.

TABLE A.

Calculations for Cyclic Operation with Methane Fuel

Basis: Feed rate 1 mole CaS0./sec.

Total cycle time 100 seconds.

TABLE B

Relative magnitude on a mole basis and temperature of the streams shown in FIG. 2 for the CH 4. - Hemihvdrate Case

Component Stream

E F

CaSO. 1.000 1.050 0.05 0.05 0.020

CaO 0.980 SO, 0.980 0.980 0.980 b

CO- 0.993 0.993 0.993

H₂0 0.500 0.500 2.486 2.486 2.486

N, 5.909 5.909 5.909 5.909 5.909

0.074 0.074 0.074 1.571 1.571

CH, 0.993

Temperature (°F) .77' 77 1444 1444 77 609 2100 77 1344 1344

TABLE C

Relative magnitude on a mole basis and temperature of the streams shown in FIG. 2 for the CH₄ - Dihydrate Case

Stream

Component

A B c E G

CaSO, 1.000 1.050 0.050 0.050 0.020

CaO 0.980 σ.

SO, 0.980 0.980 0.980

CO„ 1.106 •1.106 1.106

H₂0 2.000 2.000 4.211 4.211 4.211

N„ 6.802 6.802 6.802 6.802 6.802

0.086 0.086 0.086 1.808 1.808

CH, 1.106

Temperature (•F) 77 77 1216 1216 77 612 2100 77 1116 77

Claims

CLAIMSI claim:

1. A process for converting calcium sulfate (CaS0₄) to a solid calcium oxide (CaO) product and a gaseous sulfur dioxide (S0₂) product, comprising carrying out said process in a reactor providing two superposed fluidized beds arranged so that the fluidizing gas from the lower bed passes through the upper bed, and said process including the steps of introducing the CaS0₄ feed into the upper fluidized bed for preheating therein by fluidizing contact with the gas from the lower bed, passing the preheated feed from said upper bed to the lower fluidized bed, and converting said feed in said lower bed to said CaO and S0₂ products by ^"cyclically subjecting the lower bed to reducing and oxidizing conditions at a temperature effective for conversion of the feed to said products, the cycli¬ cal conditions in the lower bed being produced by continuously introducing preheated air beneath the lower bed as the primary fluidizing gas, and discontinuously introducing fuel increments into said primary fluidizing gas for combustion thereby, said increments being in stσichiometric excess of the oxygen avail¬ able for their combustion so that pulses of reducing gas are separated by pulses of oxidizing gas, each cycle consisting of a sequential reducing pulse and an oxidizing pulse being carried out in 10 to 150 seconds with the fuel increment being intro¬ duced for at least 70% and up to 90% of the cycle, whereby en¬ hanced conversion of CaS0₄ to CaO is obtained.

2. The process of claim 1 in which said fluidizing combustion air is preheated by indirect heat exchange with gas discharged from said upper bed.

3. The process of claim 1 in which said feed is preheated in said upper bed to a temperature in the range from 1100 to 1600°F.

.

4. The process of claims 1, 2, or 3 in which said fluidizing combustion air is preheated to a temperature in the range from 1000 to 1500°F.

5. The process of claim 1 in which said cycle is carried out in from 40 to 100 seconds.

6. The process of claims 1, 2, 3, or 5 in which the air flow rate is increased during the oxidizing pulse and reduced during the reducing pulse to maintain a substantially constant total gas flow rate.