WO2010028459A1

WO2010028459A1 - Direct reduction

Info

Publication number: WO2010028459A1
Application number: PCT/AU2009/001217
Authority: WO
Inventors: John David Winter
Original assignee: Austpac Resources N.L.
Priority date: 2008-09-15
Filing date: 2009-09-15
Publication date: 2010-03-18

Abstract

A process and apparatus for the direct reduction of metal oxides to metals. The process features the gasification and metallisation (410) within a single stage fluidised bed (412). Carbonaceous fuel (114) and metal oxide feedstock (118) may be fluidised with air (116), oxygen enrichment (417) and steam (419) at the base of the fluidised bed (412). In addition off-gas (120) from the gasification and metallisation maybe be adjusted (126, 130) for H₂O and CO₂ concentration and then fed into the fluidised bed (412) as syn gas (422) for the control of the gasification and metallisation process. Resultant metal product is recovered (132) at the top of the fluidised bed (412). Also disclosed are apparatuses and means for producing and supporting the fluidisation for the gasification and metallisation processes (410) to occur.

Description

DIRECT REDUCTION

BACKGROUND OF THE INVENTION

1. Field of the Invention

[0001] The present invention relates to a method and apparatus for the direct reduction of metal oxides to metals and in particular the direct reduction of iron (DRI).

2. Description of the Art

[0002] Direct reduction is used to reduce metal oxides to metals.

[0003] Prior art direction reduction processes and apparatus for metal oxides, and in particular iron oxide, have been proposed, but none of these prior art processes or apparatus provides an entirely satisfactory solution to the direct reduction of metal oxides in a single stage or the direct use of coal in the reduction process, nor to the ease of construction, capital cost and efficiency of direct reduction plant for metal oxides.

SUMMARY OF THE INVENTION

[0004] The present invention aims to provide an alternative direct reduction arrangement which overcomes or ameliorates the disadvantages of the prior art, or at least provides a useful choice.

[0005] hi one form the invention provides a process for the direct reduction of a metal oxide to a metal within a single stage fluidised bed reactor. The process includes the steps of: feeding to a single stage fluidised bed the metal oxide and reactants for production of a reductant gas, then production of the reductant gas, contacting the metal oxide with the reductant gas to reduce it to a metal. Preferably the process includes controlling a rate of reaction of the gasification and/or reduction by control of one or more of a CO₂ concentration or a H₂O concentration of the reactor where tHe CO₂ and/or H₂O concentrations may be controlled in a feed gas to the fluidised bed reactor. Preferably the feed gas for CO₂ and/or H₂O concentration control or adjustment is derived from a recycled off-gas discharged from the fluidised bed. Optionally a gasification rate of reaction may be controlled by introduction of a controlled amount of steam. Optionally further control may be had with the delivery to the fluidised bed of reactants such as fuel, air and oxygen.

[0006] Preferably the reactants for production of a reductant gas include a carbonaceous fuel and a reactant gas. Where the carbonaceous fuel may be selected from the group consisting of coal, oil, natural gas and hydrocarbon fuels, whilst the reactant gas may be air and/or oxygen.

[0007] Preferably the metal is iron or is predominantly iron and the metal oxide an iron oxide. Optionally the metal oxides may be obtained from iron ore, metallic oxide wastes or metal chloride wastes. Alternatively the metal oxide may be obtained from wastes such as mill scales, baghouse dusts, blast furnaces, electric arc furnaces, galvanising processes or zinc contaminated iron processes.

[0008] Preferably the production of the reductant gas is conducted in a base region of the fluidised bed with the reduction of the metal oxide being conducted in an intermediate region of the fluidised bed. Furthermore incompletely reduced metal oxide is retained in the intermediate region of the fluidised bed until reduced due to the reduced metal having a lower density than the metal oxide and any incompletely reduced metal oxide. Once the metal oxide is reduced to the metal product it is transported to an overflow region of the fluidised bed.

[0009] In a further form, the present invention provides a single stage fluid bed apparatus and method for direct reduction of a metal oxide to metal, said single stage including production of a reductant gas, and reduction of the metal oxide by the reductant gas.

[0010] Preferred aspects of the invention include:

• Production of the reductant gas within the fluid bed apparatus by contact of a fuel with the metal oxide and a sub-stoichiometric amount of oxygen;

• Recycling of discharged off-gas as a feed to the fluid bed apparatus, following treatment to modify the composition of the off-gas. In one preferred form, the recycled off-gas is treated to reduce the CO₂ and/or H₂O concentrations; • Controlling or modifying the reaction kinetics by adjusting the CO₂ and/or H₂O concentrations of the recycled off- gas.

• Controlling the feed rates and/or composition of the feed components to the fluid bed apparatus to modify reaction kinetics of the reduction;

• Overall co-current travel of reaction components through the fluid bed apparatus;

• The fluid bed apparatus configuration includes a fluid bed reaction chamber which increases in transverse cross-section from a feed end to a discharge end;

• A feed distributor configuration for feeding input gases to the reaction.

[0011] In yet another form the invention provides a single stage fluidised bed reactor for direct reduction of a metal oxide to a metal product. The reactor comprises: a support for a fluidised bed, a metal oxide feed to the fluidised bed, one or more reactant feeds for feeding reactants for production of a reductant gas within the fluidised bed. The production and presentation of reductant gas to the metal oxide produces the metal product within the fluidised bed. Metal product may then be discharged from an upper portion of the fluidised bed.

[0012] Preferably the fluidised bed in the fluidised bed reactor increases in cross-sectional area from the base portion to the upper portion by 20% to 40%. Also preferably a taper angle of a wall of the fluidised bed reactor, between the base portion and the upper portion of the fluidised bed, is in the range of 1° to 15°.

[0013] Preferably the reactor includes a distributor for the feeding of the reactant gases to the fluidised bed and for supporting and maintaining the fluidisation. Optionally the distributor feeds a recycled off-gas feed and a reactant gas feed. Where the recycled off-gas is originally obtained as discharge from the fluidised bed. Preferably the reactor also includes a CO₂ removal apparatus for at least partial CO₂ removal from the off-gas being recycled. The reactor may also include a H₂O removal apparatus for at least partial H₂O removal from the off-gas being recycled.

[0014] Optionally the reactant gas feed is introduced to the fluidised bed above the level of the recycled off-gas feed into the fluidised bed. Preferably the reactant gas feed is directed generally downwardly toward the off-gas feed which is directed generally horizontally.

[0015] Preferably the distributor may comprises: a reactant gas feed with a conical protuberance that at its apex has a reactant gas aperture formed with a corresponding conical cap to the conical protuberance and an off-gas feed comprising one or more tuyeres (other suitable feed syn-gas apertures) located in an annular arrangement about the base of the conical protuberance.

[0016] Further forms of the invention are as set out in the appended claims and as apparent from the description.

DISCLOSURE OF THE INVENTION

BRIEF DESCRIPTION OF THE DRAWINGS

[0017] The description is made with reference to the accompanying drawings; of which:

[0018] FIG. 1 is a flowchart of a process according to one embodiment of the invention;

[0019] FIG. 2 is a schematic cross-section elevation of a fluid bed reactor according to a further embodiment of the invention; and

[0020] FIG. 3 is a detailed schematic of the bottom of the fluidised bed reactor of FIG 2.

[0021] FIG 4 is an alternate embodiment of FIG 1.

[0022] FIG 5 is an alternate embodiment of FIG 3.

[0023] FIG 6 is a plan view of the transverse section along the lines 6-6 of FIG 5.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0024] FIG 1 is a flowchart of a continuous process to directly reduce a metal oxide to a metal, according to a first embodiment. While the process is shown here as a stand-alone process, it may also be combined with other operations as part of another process, for example a steel making facility, or processing of metal chloride solutions such as described in the Applicant's patent application WO 06/133500 "Processing of Metal Chloride Solutions and Method and Apparatus for Producing Direct Reduced Iron", the contents of which are incorporated herein by reference.

[0025] In the following description the process and apparatus is exemplified as the reduction of iron oxides to iron as direct reduced iron (DRI). However it will apparent to a person skilled in the art that the process and apparatus described here may be applied to other metals and their oxides, for example chrome and nickel.

[0026] In FIG. 1 a gasification and metallisation vessel ("G&M vessel")

110 contains a single, continuous fluidised bed 112. Within the fluidised bed 112 gasification to produce a reductant gas with direct reduction metallisation of the metal oxide by the reductant gas are both undertaken.

Gasification

[0027] Gasification to produce a reductant gas may be achieved within the G&M vessel 110 using a feed of either a solid, liquid or gaseous carbonaceous or hydrocarbon fuel as appropriate to local availability and cost structures. Suitable fuels that may be converted to a reducing gas (predominantly consisting of hydrogen and carbon monoxide) include coal, oil or natural gas. Common terms that may be used for reductant gases of this general type include synthetic gas ("syn gas"), water gas, producer gas and the like. The term "syn gas" is used herein.

[0028] The gasification reactions involved are quite complex but have been exhaustively studied and reported in the literature. Examples of some of the more simple reactions involved that may be used here are given below:

C + O₂ → CO₂

2C + O₂ → 2CO

2C + 3H₂O → CO + CO₂ + 3H₂ CH₄ + 2O₂ → CO₂ + 2H₂O

CH₄ + H₂O → CO + 3H₂

[0029] The oxygen for these reactions may be supplied either as ambient or preheated air and with or without oxygen enrichment. The aim is to ensure that only sufficient oxygen is used to provide enough heat of combustion to maintain the endothermic reactions that generate the CO and H₂ components necessary for the metal oxide reduction reactions occurring simultaneously. For example, it has been found that satisfactory results may be obtained with a substoichiometric oxygen supply of between 30% to 50% of the full stoichiometric combustion requirement.

[0030] In FIG 1 the reactants for gasification are added to the single stage fluidised bed 112 as coal 114 in a preferred particle size range of 2 to 12 mm, air 116 pre-heated to less than 950⁰C or more preferably to a temperature range of 300° to 700⁰C or 400° to 700⁰C and recycled, modified syn gas 122 in a preferred temperature range of 800° to 1000⁰C. The reductant gas formed by the gasification reaction is supplemented with a feed of modified, recycled off-gas from the reactor discharge, as will be described in more detail later.

Direct Reduction

[0031] In the continuous process of FIG 1 the metal oxide 118 for reduction is introduced into the single stage fluidised bed 112. In the present example a Fe₂O₃ particulate feedstock 118 of a preferred particle size range of 0.5 to 4 mm - preferably pre-heated using waste heat to between 800 and 1000⁰C, and more preferably approximately 900⁰C, may be used. The Fe₂O₃ feedstock 118 may be iron ore, metallic oxide wastes such as mill scales and baghouse dusts, waste from zinc contaminated iron processes, wastes from blast and electric arc furnaces or derived from part of another process such as described in patent application WO 06/133500 "Processing of Metal Chloride Solutions and Method and Apparatus for Producing Direct Reduced Iron". Other feedstocks derived from application areas such as galvanising processes are described below with respect to FIG 4. [0032] Within the fluidised bed 112, the metal oxide 118 is fluidised by the gases and initially reduced by contact with the reductant gases, in the case of the present example iron (III) oxide, into an oxide of lower valence, Fe(II), without metallising. It is important that the initial reduction takes place rapidly so that the formation of FeO is predominant and the formation of the intermediate oxide Fe₃O₄ is minimised; as too high a proportion OfFe₃O₄ may lead to incipient fusion and/or a critical amount of accretion resulting in a consequent collapse of the fluidised bed 112 and/or the accretion OfFe₃O₄ onto the reactor vessel 110 internal surfaces. The design and construction to minimise incipient fusion and accretion within the G&M vessel 110 and the fluidised bed 112 is described below with respect to FIG 2.

[0033] The reactions taking place in the initial reduction within the fluidised bed 112 are:

3Fe₂O₃ + H₂ → 2Fe₃O₄ + H₂O 3Fe₂O₃ + C0 → 2Fe₃O₄ + CO₂

Fe₃O₄ + H₂ → 3FeO + H₂O Fe₃O₄ + CO → 3FeO + CO₂

[0034] The inventor has further noted that within the fluidised bed 112 that the presence of the FeO may also act as a catalyst for the gasification reactions described here.

[0035] The fluidised bed 112 for the reduction and gasification steps is maintained in a temperature range of 750° to 105O⁰C or in a preferred temperature range of 870 ° to 92O⁰C with a most preferred temperature of approximately 900⁰C. The pressure within the G&M vessel 110 may be approximately atmospheric or just above as required to sustain process flows, for example nominally about 20-40 kPa above the atmospheric pressure.

[0036] Continued contact of the FeO with the reductant gases as the mix of partially reduced iron particles and reductant gases passes through the reactor results in further reduction of the iron to the metallic state.

[0037] The reactions for the final reduction are:

FeO + CO → Fe + CO₂ FeO + H₂ → Fe + H₂O

[0038] The minor proportion of Fe₃O₄ in the reaction products from the first stage is also reduced by reaction with the CO, via formation of FeO: Fe₃O₄ + CO → 3FeO + CO₂

FeO + CO → Fe + CO₂

[0039] The inventor has also noted that in the overall reduction reactions described that the oxygen released from the Fe₂O₃ during reduction is immediately available for the gasification reactions since gasification and reduction reactions occur simultaneously in the feed region of the single stage, continuous fluidised bed 112. The construction of the fluidised bed 112 is described in further detail with respect to FIG 2 below.

[0040] After final reduction the metal product 132, in this example Fe_(S)

132 (directly reduced iron, "DRI"), is discharged from the fluidised bed 112. The recovery of the Fβ(_S) 132 is described in detail below with respect to FIG. 2. The hot Fe(_s) discharged from the fluidised bed 112 may be then be pelletised or hot briquetted to form hot briquetted iron (HBI) as known per se in the art, as appropriate for its intended use. For example the HBI form may be most suitable for use as feed for electric arc furnaces.

[0041] Once the Fe(s) product has been formed into the desired shape it may be indirectly cooled under such conditions as to exclude air and so avoid any reoxidation of the product which could occur whilst the material is at an elevated temperature. Nominally, the metallised pellets will be cooled to less than 200° C or lower before contact with air is allowed.

[0042] The inventor has found that reaction rate of the reductions described here may be controlled by adjusting the rates and composition of the fuel 114, air 116 and/or syn gas 122 feeds to the reactor such that the off gas 120 composition from the fluidised bed 112 is according to the equation:

K = ([CO] + [H₂]) / ([CO] + [CO₂] + [H₂O] + [H₂]) > 0.7, Where [ ] denotes the partial pressure. However it will be readily appreciated that other equivalent expressions may be used with equivalent concentration or proportion terms to the example given here to partial pressure expression.

[0043] A value of K>0.7 is preferred in order for the reduction reactions to proceed at an appropriate speed within the fluidised bed 112. By way of example, preferably K is approximately 0.9, to obtain a relatively rapid reduction of the metal oxide and hence a low mean residence time in the order 20 minutes, up to 60 minutes at lower K~0.7, for completion of the reduction reaction.

[0044] The rate of reaction may thus be controlled by modifying the concentrations of CO₂ and H₂O, with the rate of reaction increasing as the concentrations of CO₂ and/or H₂O decrease. In one preferred form, these concentrations may be controlled by adjustment of the composition of the modified off-gas recycled back to the reactor as syn gas, and/or the amount of the recycled syn gas relative to the fuel and air feed.

[0045] FIG 1 illustrates one method for treatment of the off-gas 120 CO₂ and/or H₂O concentrations to produce a modified syn gas 122 for recycle feed back to the fluidised bed 112, in order to achieve the desired reaction rate for the reductions.

[0046] A part 121 of the off-gas 120 discharge from the top of the fluidised bed 112 is split off and fed through a heat exchanger 124 that is used to preheat the modified syn gas 122 that is supplied to the fluidised bed 112.

[0047] The remaining or excess portion 136 of the off gas 120 is fed to an afterburner 140 where it is combusted with combustion air 138. The hot combustion gases 142 then being fed to a heat exchanger 144 for further heating of the syn gas feed 122. The spent waste gases 146 from heat exchanger 144 may then be piped away for waste heat recovery.

[0048] The cooled off-gas 123 is fed into a H₂O removal system 126 to adjust the concentration of H₂O in the off-gas 128. The H₂O removal apparatus 126 may be of a type known per se, for example the use of cooling water 133 to condense (or quench) H₂O as a condensate 134 from the off gas 128 or other appropriate system or arrangement for dehumidifying the off-gas 120.

[0049] Preferably, in operation the H₂O removal apparatus is operated so as to lower the H₂O content from about 7 % in the off gas 120, 123 to about 1% v/v (volume / volume) in the syn gas feed 128, 122, such that preferably about 85% of the moisture present in the off-gas 120 is removed.

[0050] Associated with the H₂O removal system may be a removal system (not shown) for sulphur compounds. The sulphur removal system may be of a type known per se, for example a caustic (NaOH) scrubber, or other appropriate system or arrangement for removing sulphur from the off-gas.

[0051] The dehumidified off-gas 128 may then be fed into CO₂ removal apparatus 130 to reduce the concentration of CO₂ in the off-gas 128. The CO₂ removal apparatus 130 may be of a type known per se, for example a monoethanolamine (MEA) scrubber or other appropriate system or arrangement for reducing the CO₂ of the off-gas.

[0052] Preferably, in operation the CO₂ removal apparatus is operated so as to lower the CO₂ content from about 8% down to about 1% v/v, such that preferably about 90%, of the CO₂ present in the off-gas 120 is removed 131.

[0053] The CO₂ and H₂O concentrations of the reactor feed gases may be controlled so as to control the reaction rate/s described above. Means for controlling the CO₂ and H₂O concentrations include (i) varying the operation of the CO₂ and H₂O removal apparatuses 130,126 to modify the amount of CO₂ and H₂O removed, (ii) treating a variable proportion of the off-gas 120 using the CO₂ and H₂O removal apparatus, and then optionally blending the treated and untreated off-gases prior to or at the feed to the reactor; or (iii) varying the amount of recycled off-gas fed back to the reactor 110.

[0054] FIG 1 also illustrates an example of how the heat balance within the fluidised bed 112 varies in conjunction with (iii) above. As the proportion of off- gas 120 bled off to the recycle stream 121 is varied, the remaining or excess off gas 136 combusted with combustion air 138 in the afterburner 140 to produce hot combustion gases 142 also varies, and thus the amount of the hot combustion gases 142 to the second heat exchanger 144.

[0055] By way of example, the relative proportions of the various reactants, with air and syn gas expressed in m³/h at standard temperature and pressure (STP) and coal 114 and iron oxide feed are in units of kg/h, all per kg/h of coal feed, may be approximately: air 116 : recycled, modified off gas 122 : coal 114 : Fe₂O₃ feed 118

1.0-1.5 m³/h : 1.2-2.0 nrVh : 1.0 kg/h : 1.5-6.0 kg/h and more preferably for a stand-alone process for iron ore feed:

1.2 : 1.6 : 1.0 : 3.2

Or in an embodiment where the iron oxide is derived from the pyrohydrolysis step of WO 06/133500 as described earlier:

1.2 : 1.6 : 1.0 : 1.6

[0056] FIG. 2 illustrates a fluid bed reactor arrangement 110 according to a second embodiment of the invention.

[0057] The reactor 110 is upright and generally cylindrical, with a refractory lining 214 forming a generally frusto-conical reaction chamber defined by tapered side walls 210, which contain the fluidised bed 112.

[0058] The feed apparatus for the respective reactants feed to the bottom of the fluidised bed - the metal oxide feedstock 118 via a metal oxide inlet tube 216, coal or other fuel 114 via a fuel inlet tube 218, and air 116 and recycled and modified syn gas 122 via a distributor arrangement 222 which will be described in more detail later with reference to FIG. 3.

[0059] The reactor 110 further includes a bed drain discharge 224 communicating with the reaction chamber at the base of the fluid bed 112, for removal of large particles and purging of the reactor 110. There may also be an overflow removal port / metal product discharge 226 for removal of the DRI product 132, and an off-gas discharge 120 at the top of the reactor vessel 110.

[0060] The inlet tubes 216 and 218 for the solid reactants are preferably angled down through the side wall of the reaction chamber 110 to a location near the base of the fluidised bed 112, just above the distributor 222. The metal oxide 118 and coal 114 feeds may be fed by gravity or pneumatically assisted into the reaction chamber, or by any other suitable means. Preferably, the solid material feeds enter the reaction chamber adjacent the gas inlets via the distributor, in a manner to facilitate initial mixing of the reactants. [0061] The recycled, modified syn gas 122 and air 116 are fed via the distributor 222 into the bottom of the fluidised bed, causing fluidisation of the bed 112.

[0062] On an overall, macro level, the gaseous and solid reactants both move upwards through the reactor, although there will at any time be a proportion of particles within the fluidised bed which are moving counter to this. Thus the gasification and initial stage reduction reactions occur predominantly at the bottom / base portion of the fluidised bed, any transient formation OfFe₃O₄ occurring at an intermediate height within the bed, and the reduction of FeO to Fe metal occurring mostly in the upper portion of the bed.

[0063] The outward taper 212 of the reaction chamber side walls 210 is designed to provide an increase in the transverse (in the example of FIG. 2, horizontal) cross sectional area of the fluidised bed 112 between the bottom of the fluidised bed 112 adjacent the feeds 114, 118, 222 and the top of the chamber adjacent the metal discharge 132. The increase in transverse cross-sectional area being in the order of 20- 40%, preferably about 30%, to accommodate a reduction in particle density of the metal oxide feedstock as it is reduced to the metal. The taper angle 212 of the side wall may be dependent for example on the bulk flow rate up through the reactor and desired residence time, but may for example be in the range of 1-15 degrees, more preferably about 5-10 degrees.

[0064] As a result of the general reactor configuration and tapered fluidised bed 112 shape, the process achieves faster gas velocities at the base of the fluidised bed where the reactants are being fed and mixed initially, and then a lowering in mean gas velocity reduces as the reaction progresses and travels up through the fluidised bed, due to the increasing reactor cross-sectional area of the fluidised bed. Thus the appropriate fluidising velocities, varying with bed depth, for the various stages of the reactions corresponding to changes in particle density are maintained through the depth of the fluidised bed.

[0065] Waste heat from the process, for example from cooling of the

DRI pellets or combustion gas from heat exchanger 144, may be used for pre-heating feed streams for the process or for other process on-site, or for co-generation of electricity for site requirements or sale back to the grid. [0066] FIG. 3 is a detail of the bottom / base region of the fluidised bed reactor 112, showing the air 116 and recycled syn gas 122 inlets via distributor 222.

[0067] The distributor 222 is situated adjacent the bottom of the reactor

110, forming a plenum chamber 316 below, between the recycled syn gas inlet 122 and a series of syn gas apertures 312 in the distributor 222.

[0068] The distributor 222 is itself constructed so as to form an internal air plenum 318 which receives and distributes air 116 from an air or oxygen supply tube 220 to a series of air flow passages 310 in a top refractory layer 314 of the plate.

[0069] This construction allows the modified off gas / syn gas 122 and the air 116 feeds to be kept separate and contacted only within the fluidised bed, where they are contacted with the coal 114 and metal oxide 118 particles. The alternating rings or other patterns of the air apertures 310 and syn gas apertures 312 and relatively high gas inlet velocities facilitate mixing at the bottom of the fluidised bed 112 as well as the fluidising of the bed 112.

[0070] The number arrangement and size of the air and syn gas inlets in the distributor 222 may be varied according to geometry and feed materials of the particular reactor.

[0071] Alternative air/oxygen and syn gas inlet arrangements may be used, for example mixing nozzles may be used instead of or in addition to a distributor. An alternate embodiment is described below with respect to FIGs 5 and 6.

[0072] FIG 4 is an alternate embodiment of the continuous metal reduction process of FIG 1. Like numbers have been used to denote like features between FIGs 1 and 4. The alternate G&M vessel 410 features an adaption to mix oxygen 417 with the air 116 for injection 415 into the base region of the fluidized bed 412 as described in detail with respect to FIG 5. Oxygen 417 may be mixed with the air 116 to enrich the oxygen content delivered to the fluidized bed up to approximately 30%, or more preferably in the range of 25% to 50% v/v (volume/volume). The enrichment of the oxygen content has the benefit of lowering the nitrogen content in the process as well as improved reaction and fuel values with the fuel 114 and metal oxide 118. Alternatively 50% to 100% oxygen 417 concentration may be mixed with the air 116. At 100% oxygen concentration therefore, no air is fed to the fluidised bed.. At higher concentrations of oxygen 417 it may be desirable to include steam (water vapour) 419 with the gas mix injected 415. The addition of the steam 419 is to compensate for the increased temperatures that may result from the use of high oxygen 417 concentrations, the consequences of which are described below with respect to FIGs 5 and 6 below. The steam 419 may quench or otherwise reduce temperatures at the base of the fluidised bed 412 by the thermal mass of the steam and/or the steam (H₂O) entering into endothermic gasification reactions (e.g. the water shift reaction, given earlier) with the additional advantage of extra H₂ being generated within the fluidised bed 412.

[0073] The oxygen 417 and air 116 gas mix may be delivered with no preheating, that is an ambient temperature gas in contrast to the preheating used with respect to the embodiment of FIG 1. The steam 419 may be mixed into the gas injection stream 415 as appropriate to maintain the vapour phase of the steam. It will be readily appreciated that the gas mix injection 415 is as appropriate to maintain the fluidisation of the fluidised bed 412.

[0074] The process of FIG 4 also features a single heat exchanger and a bag house filter as well as other differences which are described in the following. Off- gas 120 from the alternate G&M vessel 410 is fed into a heat exchanger 424 which is used to pre-heat the modified off-gas / syn gas 422 prior to its injection into the fluidised bed 412. The cooled off-gas 423 from the heat exchanger 424 may then be temperature controlled to a constant temperature between 120° to 200⁰C by suitable direct injection of water 448 into the gas stream of the off-gas 423. The level of water injection 448 is always below water saturation for the off-gas 423 together with the temperature of the off-gas 423 being always above the dew point prior to entry into the bag house filter 450. The bag house filter 450 may be used to remove fine particulates 451 from the off-gas 423 The bag house 450 may be substituted with any other suitable filter as selected by a person skilled in the art, for example micro-pore ceramic filters. The fine particulates removed 451 may include carbon complexes and volatile metal oxides. The particulate metal oxides may result from volatile metals liberated in the fluidised bed 412 which are then oxidised in some manner to result in fine particulates within the off-gas 120. Zinc may be an example of one such volatile metal which may be generated when a feedstock 118 derived from zinc-iron-chloride solutions is used. Zinc-iron-chloride solutions may be typically associated with galvanising processes,

[0075] The filtered off-gas 452 may then be fed to the respective H₂O

126 and CO₂ 130 removal systems as described previously with respect to FIG 1. The H₂O and/or CO₂ modified off-gas / syn gas 422 may then be pre-heated by heat exchanger 424 as described above and then injected into the fluidised bed 412.

[0076] Excess off-gas 436 may be dealt with in a similar manner to that of FIG 1. The excess off-gas 436 may be removed from the recycled off-gas stream after the bag house filtration 450. The excess off-gas 436 may then be combusted with combustion air 138 in an afterburner 140, the hot waste gas 142, 146 being then utilised in waste heat recovery as required. By way of example for the alternate embodiment of FIG 4, the relative proportions of the various reactants, with air, oxygen and syn gas expressed in m³/h at standard temperature and pressure (STP) and coal 114 and iron oxide feed are in units of kg/h, all per kg/h of coal feed, may be approximately:

Proportions: air 116 : O₂ 417 : recycled, modified off gas 422 : coal 114 : Fe₂O₃ feed 118

1.0-1.5 m³/h : 0.2-0.8 m³/h : 1.2-2.0 m³/h : 1.0 kg/h : 1.5-6.0 kg/h and more preferably for a stand-alone process for iron ore feed:

1.2 : O₂ as per below : 1.6 : 1.0 : 3.2

Associated Oxygen 417 Proportion (by way of example)

O₂ enrichment (to air) Low (30%) High (50%)

Air (m³/h) 116 0.9 0.8

O₂ (m³/h) 417 0.03 0.15

Syngas (m³/h) 422 0.2 0.08

Coal (kg/h) 114 1.0 2.5

Fe₂O₃ (kg/h) 118 1.5 6.0

Where the addition of steam 419 to the injected gas stream 415 is as per required. [0077] FIG 5 is a distributor 522 in an alternate embodiment of that shown in FIGs 2 and 3. Like numbers have been used to denote like features between FIG 5 and those of FIGs 2 and 3. Air 116 and/or oxygen 417 are supplied by the tube 220 that emerges from an apex of a cone 554 of the alternate distributor 522. A cone cap 556 at the apex of the cone 554 redirects the air 116 and/or oxygen 417 gas streams in the formed aperture 310 mostly downwards, as shown by the undulating arrows, into the base region of the fluidised bed and towards the syn gas tuyeres 558. The syngas tuyeres 558 contain apertures 312 which are orientated so as to project the syn gas / modified, recycled off-gas 122, 422 mostly horizontally, as shown by the respective undulating arrows. The syngas tuyeres 558 are attached to an annular plate 560 of the alternate distributor 522. Items such as the refractory layer 314 have been omitted for clarity.

[0078] FIG 6 is a plan view of the alternate distributor 522 along the sectional line of 6-6 in FIG 5. FIG 6 in particular illustrates the flow of the air 116, oxygen 417 and syn gas 122, 422 gas streams from the top of the distributor 522 and into the bottom or base region of the fluidised bed 112, 412, as shown by the respective undulating arrows.

[0079] The introduction of the air 116 and oxygen 117 higher in the fluidised bed, via the cone 554 and cone cap 556, than the syn gas 122, 422 minimises the formation and promotion of sintered, fused, agglomerates / accretions of "sticky" particles; for example Fe₃O₄, particular metal oxides and as described previously with respect to FIG 1. Such agglomerates / accretions and the like are prone to form where the fluidised bed is at a higher temperature, "hot spots", as may occur where the bed has a higher proportion of air and/or oxygen. The configuration of the alternate distributor 522 allows any agglomerates / accretions which may form in the vicinity of the air/oxygen apertures 310 to sink in the fluidised bed, due to their higher specific gravity, to the base of the fluidised bed where conditions are not favourable to further agglomeration / accretion and where reactions to reduction of the metal oxide may proceed more favourably.

[0080] One example of the start-up of the above process may be by: (i) Initial heating of a bed of iron ore within the G&M vessel 110, 410 by gas burner, then feeding in coal 114 and air 116 under full oxidative conditions to reach a desired high bed temperature.

(ii) The condition is then converted to gasification by supplying a bolus ("dumping") coal into the fluidised bed with a consequent reversal of temperature control to a lower value, (iii) Recycling of a proportion of the off gas 120 and the H₂O and CO₂ removal systems 126, 130 are started in order to produce the required composition of modified off gas 122, 422 to start and control the reduction reactions to metallisation of the iron oxide. [0081] An example of the shut down of the above process may be by reversal of the start-up process with the addition of water injection (not shown) into the fluidised bed. In addition in FIG 4 the modified off-gas 422 may bypass the heat exchanger 424 via the bypass line 452 so that cool modified/recycled off-gas 422 is

8 only presented to the fluidised bed shut down.

[0082] The present invention in its preferred forms thus provides a continuous, single stage process for direct reduction of iron or other metals, which it is believed will overcome or ameliorate at least some of the problems of the prior art, such as complexity and capital cost, and operational problems such as those caused by the formation of a 'sticky' Fe₃O₄ phase (or other agglomerating / accretion prone contaminants or metal oxide phases), higher pressures for operation, multiple stages and their inherent multiple material transfer systems..

[0083] In this specification, the word "comprising" is to be understood in its "open" sense, that is, in the sense of "including", and thus not limited to its "closed" sense, that is the sense of "consisting only of. A corresponding meaning is to be attributed to the corresponding words "comprise, comprised and comprises where they appear.

[0084] While particular embodiments of this invention have been described, it will be evident to those skilled in the art that the present invention may be embodied in other specific forms without departing from the essential characteristics thereof. The present embodiments and examples are therefore to be considered in all respects as illustrative and not restrictive, and all changes which come within the meaning and range of equivalency of the claims are therefore intended to be embraced therein. It will further be understood that any reference herein to known prior art does not, unless the contrary indication appears, constitute an admission that such prior art is commonly known by those skilled in the art to which the invention relates.

Claims

1. A process for direct reduction of a metal oxide to a metal, comprising: feeding to a single stage fluidised bed the metal oxide and reactants for production of a reductant gas; production of the reductant gas within the fluidised bed; contacting the metal oxide with the reductant gas within the fluidised bed; reduction of the metal oxide by the reductant gas to the metal.

2. A process according to claim 1 further including the step of : controlling a rate of reaction by control of one or more of a CO₂ concentration or a H₂O concentration.

3. A process according to claim 2 wherein said control of the CO₂ concentration or the H₂O concentration comprises controlling one or more of the CO₂ concentration or the H₂O concentration of a feed gas to the fluidised bed.

4. A process according to claim 3 wherein said feed gas is an at least partially treated recycled off-gas discharged from the fluidised bed.

5. A process according to claim 3 or 4 wherein one or more of the CO₂ concentration or the H₂O concentration in the feed gas is controlled such that an equilibrium constant defined by the equation

([CO] + [H₂]) / ([CO] + [CO₂] + [H₂O] + [H₂]) is greater than 0.7; where the concentration is a partial pressure.

6. A process according to claim 5 where the equilibrium constant is approximately 0.9.

7. A process according to any one of claims 4 to 6 where the CO₂ concentration in the feed gas is about 1% volume / volume.

8. A process according to any one of claims 4 to 6 where the H₂O concentration in the feed gas is about 1% volume / volume.

9. A process according to any preceding claim where the reactants for production of a reductant gas include a carbonaceous fuel and a reactant gas; where the carbonaceous fuel is selected from the group consisting of coal, oil, natural gas and hydrocarbon fuels.

10. A process according to claim 9 where the reactant gas is air or oxygen or air and oxygen.

11. A process according to claim 10 where the proportion of oxygen in the reactant gas is up to 30% volume / volume.

12. A process according to claim 10 where the proportion of oxygen in the reactant gas is in the range of 25% to 50% volume / volume.

13. A process according to claim 10 where the proportion of oxygen in the reactant gas is in the range of 50% to 100% volume / volume.

14. A process according to any preceding claim further including the step of: controlling a gasification rate of reaction by introduction of a controlled amount of steam.

15. A process according to any preceding claim where a temperature within the fluidised bed is within the range of 750° to 1050⁰C.

16. A process according to any preceding claim where a temperature within the fluidised bed is within the range of 870° to 92O⁰C.

17. A process according to any preceding claim where a temperature within the fluidised bed is approximately 900⁰C.

18. A process according to any preceding claim, wherein the metal is iron or is predominantly iron.

19. A process according to claim 18, wherein the metal oxide is an iron oxide.

20. A process according to any preceding claim, wherein the metal oxide is obtained from iron ore, metallic oxide wastes or metal chloride wastes.

21. A process according to any preceding claim, wherein the metal oxide is waste obtained from one or more of mill scales, baghouse dusts, blast furnaces, electric arc furnaces, galvanising processes or zinc contaminated iron processes.

22. A process according to any one of claims 18 to 20 where the proportions of: air : recycled, modified off gas : coal : Fe₂O₃ feed are

1.0 - 1.5 m³/h : 1.2 - 2.0 m³/h : 1.0 kg/h : 1.5 - 6.0 kg/h.

23. A process according to any one of claims 18 to 20 where the proportions of: air : oxygen : recycled, modified off gas : coal : Fe₂O₃ feed are

1.0 - 1.5 mVh : 0.2 - 0.8 mVh : 1.2 - 2.0 m³/h : 1.0 kg/h : 1.5 - 6.0 kg/h

24. A process according to any preceding claim wherein step of the production of the reductant gas is conducted in a base region of the fluidised bed.

25. A process according to claim 24 wherein the step of the reduction of the metal oxide is conducted in an intermediate region of the fluidised bed.

26. A process according to claim 25 further including the steps: retaining the incompletely reduced metal oxide in the intermediate region of the fluidised bed; and transporting the reduced metal to an overflow region of the fluidised bed.

27. A process according to claim 26 wherein the reduced metal has a lower density than the metal oxide and incompletely reduced metal oxide.

28. A single stage fluidised bed reactor for direct reduction of a metal oxide to a metal product comprising: a support for a fluidised bed; a metal oxide feed to the fluidised bed; one or more reactant feeds for feeding reactants for production of a reductant gas within the fluidised bed, such that the reductant gas contacts the metal oxide to produce the metal product within the fluidised bed.

29. A reactor according to claim 28, wherein the reductant gas is produced by contact of the reactant feeds of a fuel with the metal oxide and a sub-stoicbiometric amount of oxygen in a base portion of the fluidised bed.

30. A reactor according claim 28 or 29 including a metal product discharge from an upper portion of the fluidised bed.

31. A reactor according to claim 30 wherein the fluidised bed in the fluidised bed reactor increases in cross-sectional area from the base portion to the upper portion.

32. A reactor according to claim 31 wherein the increase in cross-sectional area is in the range of 20% to 40%

33. A reactor according to claim 31 or 32 where the increase in cross-sectional area is approximately 30%.

34. A reactor according to any one of claims 31 to 33 wherein a taper angle of a wall of the fluidised bed reactor, between the base portion and the upper portion of the fluidised bed, is in the range of 1° to 15°.

35. A reactor according to claim 34 where the taper angle is in the range of 5° to 10°.

36. A reactor according to any one of claims 28 to 35 wherein the reactant feeds include a distributor for feed of the reactant gases so as to cause fluidisation of the fluidised bed.

37. A reactor according to claim 36 wherein the reactant feed comprises a recycled off-gas feed and a reactant gas feed.

38. A reactor according to claim 37 wherein the reactant gas feed is introduced to the fluidised bed above a level of the off-gas feed.

39. A reactor according to claim 38 wherein the reactant gas feed is directed downwardly toward the off-gas feed.

40. A reactor according to claim 39 wherein the distributor comprises: the reactant gas feed comprising a conical protuberance with an apex and a first gas aperture formed with a corresponding conical cap; and the off-gas feed comprising one or more tuyeres located in an annular arrangement about the base of the conical protuberance; where one or more tuyeres are adapted to project the recycled off-gas generally horizontally.

41. An apparatus for direct reduction of a metal oxide to a metal product, including: a fluidised bed reactor according to any one of claims 28 to 40; an off-gas recycling arrangement for recycling an off-gas discharged from the fluidised bed back into the reactor; and a CO₂ removal apparatus for at least partial CO₂ removal from the off-gas being recycled.

42. An apparatus according to claim 41 wherein the CO₂ removal apparatus is a monoethanolamine (MEA) scrubber.

43. An apparatus for direct reduction of a metal oxide to a metal product, including: a fluidised bed reactor according to any one of claims 28 to 42; an off-gas recycling arrangement for recycling an off-gas discharged from the fluidised bed back into the reactor; and a H₂O removal apparatus for at least partial H₂O removal from the off-gas being recycled.