Process for producing trichlorosilane

The invention relates to an improved process to manufacture TCS in a polysilicon plant based upon combining a high temperature FBR process reacting metallurgical grade silicon, hydrogen, and silicon tetrachloride (STC) to make trichlorosilane (TCS) and a high temperature thermal converter to hydrogenate STC to TCS and hydrogen chloride.

Not applicable.

Not applicable.

FIELD OF THE INVENTION

This invention relates generally to processes to make halosilanes, most commonly trichlorosilane (TCS). It involves chemical synthesis of TCS from metallurgical grade silicon (MGSI), silicon tetrachloride (STC), hydrogen (H2), and hydrogen chloride (HCl) in a fluidized bed reactor (FBR) and also from only H2and STC in a gas phase thermal reactor. It presents advantageous methods to combine these two different processes to make TCS into a single process train thus minimizing both capital and operating costs.

BACKGROUND OF THE INVENTION

Chemical vapor deposition (CVD) reactors are used to produce polycrystalline silicon (polysilicon), the key raw material used in the manufacture of most semiconductor devices and silicon-based solar wafers and cells. The most widely used method for producing polysilicon is the Siemens reactor process generally according to the primary reactions below:
HSiCl3+H2→3HCl+Si
HCl+HSiCl3→SiCl4+H2
Net: 4HSiCl3→Si+2H2+3SiCl4

In commercially significant polysilicon processes, STC, H2, dichlorosilane (DCS) and HCl are all significant byproducts of the CVD operations along with lesser amounts of monochlorosilane (MCS), silane (SiH4), and trace contaminants containing metals, donor and acceptor atoms, and carbon bearing species. In addition to TCS, significant amounts of H2are fed to the CVD reactor and the single pass conversion of TCS is well under 50%. Variations of this process have been in commercial existence for about fifty years and are widely reported in the literature. In this process, high temperature polysilicon rods are placed in a reactor, and trichlorosilane (TCS) gas is passed over these rods. A portion of the silicon in the gas is deposited on the rods, and when the rods have grown large enough, they are removed from the reactor. The end product is in the form of polysilicon rods or chunks, which can be further processed into ingots, then sliced into wafers that are made into solar cells, for example.

In a related process, TCS is disproportionated to form silane (SiH4) and STC. The silane produced is used in many processes associated with semiconductors and other products, including making polysilicon in either a Siemens reactor or fluidized bed CVD process. The fluidized bed process makes silicon in irregular, but nominally spherical beads in diameters typically ranging up to about 2 mm diameter. The general chemistry of these reactions is as follows:
4HSiCl3→SiH4+3SiCl4
SiH4→Si+2H2
In the exhaust of a TCS CVD reactor, the chlorosilanes (Six—Cl4−x) are separated from the the non-condensable gases (H2and HCl) by relatively simple condensation steps and then followed by energy intensive and capital intensive processes consisting of compression, HCl absorption, HCl stripping, and H2pressure-swing adsorption (PSA) steps to remove HCl from the H2. The by-products of the STC converter are similar to those produced in the CVD reactors and have traditionally been subjected to the same or very similar off-gas recovery (OGR) designs.

In all commercial processes using TCS as a feedstock, the initial step of production of TCS can be further classified as direct chlorination process or the hydrochlorination process, both processes being widely referenced in the literature, though nomenclature is not necessarily consistent. It is the TCS thus produced that is reacted with H2in Siemens style chemical vapor deposition (CVD) reactors or disproportionated to make SiH4.

The process converting TCS into silane and the CVD-based Siemens process for manufacturing polysilicon both produce a large amount of the byproduct silicon tetrachloride (STC). For example, a maximum of about 20 kg of STC is made as a byproduct for every kg of polysilicon or silane produced. It is possible, however, to hydrogenate STC forming TCS by reacting STC with hydrogen in the gas phase in a reactor commercially referred to as an STC Converter at 800-1200° C. where 14-24% of the STC is typically hydrogenated in each pass through the reactor according to the following reaction:
SiCl4+H2→HCl+HSiCl3
The product TCS can then be recycled to a series of silane disproportionation reactors and separation steps to make silane, or to a CVD reactor for direct production of more polysilicon.

Most commercial STC converter processes use molar feed ratios of H2:STC between 2.0:1 and 3:1 according the process ofFIG. 1. In this process, the STC converter reaction chemistry is insensitive to pressure, but the industrially practiced range has been predominantly 4-7 BarG. Most commercially available STC converters are retrofitted Siemens CVD reactors having multiple graphite rods for heaters and a flat baseplate where feed-through electrical connections are made to the graphite rods. When conducted in a Siemens style STC converter, it is costly to increase the pressure due to a large flat baseplate forming the bottom of the reactor and the reaction step itself does not benefit from increased pressure. In the entire STC converter process to make TCS, the OGR system which separates reactor effluent products and reactants and recycles unreacted feeds must be considered, and this part of the process benefits significantly from operation at higher pressure.

The OGR process is a very energy intensive process mainly due to the difficulties of separating H2and HCl. Since, hydrogen and HCl cannot be separated by a simple condensation process and would require excessive energy if separated via cryogenic distillation, alternate methods such as membrane separation processes have been evaluated. Even a membrane separation process is not deemed feasible for separation of H2and HCl. No durable membrane technology exists at a competitive capital cost. Consequently, polysilicon producers typically use STC to absorb HCl from the H2/HCl mixture. TCS, which has higher vapor pressure than STC, has also been used to absorb HCl. Whether using TCS or STC to absorb HCl from H2, significant amounts of the chlorosilane exist in the H2when it is subsequently fed to a PSA (pressure swing adsorption) or TSA (temperature swing absorption) operation to remove residual HCL and trace impurities. The presence of these chlorosilanes reduces the capacity of PSA or TSA adsorbent beds to adsorb the impurities of concern. Use of either TCS or STC for HCl absorption is an energy intensive process.

Another process to hydrogenate STC to TCS is what is commercially known as a hydrochlorination FBR process. In this process, MGSI, H2, and STC are fed into a FBR where the following net reaction occurs, typically at a temperature of 500 to 600° C. and typically 15-30 BarG.
Si+2H2+3SiCl4→4HSiCl4
In this process, approximately 15-25% of the STC fed is reacted in a single pass and typical molar feed ratios are 1.5:1 to 2.5:1 H2:STC. The process benefits significantly from higher pressures and temperatures due to improved equilibrium conversion. As a result, the process is very expensive due to high pressure and temperature ratings required on relatively large process equipments. Vessels must be made of fairly thick walls with high nickel alloys. Both the metals used and the fabrication techniques are quite expensive.

If STC could not be hydrogenated and recycled, there would be a huge loss of the primary raw materials silicon and chlorine and a cost for disposal of the byproduct STC. Thus, efficient polysilicon plants are built as a substantially closed loop processes as illustrated inFIGS. 1 and 2.FIG. 1Ashows a typical polysilicon plant utilizing hydrochlorination technology. In this drawing, a hydrochlorination plant2provides TCS to CVD operations1and the CVD operations return H2and STC to the TCS plant. Impurities and byproducts are purged from the TCS plant and make-up raw materials MGSI, H2, and STC/HCl are fed to the plant. The hydrochlorination plant produces TCS from MGSI and also hydrogenates STC from the CVD operations.

FIG. 1Bshows a typical polysilicon plant utilizing a direct chlorination plant4to make TCS from MGSI plus HCl and an STC hydrogenation plant3to hydrogenate STC plus H2to TCS. CVD operations1are the same as in the hydrochlorination based plant.

FIG. 1Cshows a silane plant5coupled to a TCS plant6, and CVD operations7. InFIG. 1C, TCS plant6can represent the non-CVD operations ofFIGS. 1A and 1B. In this FIG., TCS is supplied from the TCS plant to a silane plant which produces Silane and returns STC according the reaction:
4HSiCl3→SiH4+3SiCl4
It should be understood that inFIGS. 1A, 1B, and 1C, many operations occur within each block and additional minor feed streams may exist as feeds enter and byproduct and impurity streams may leave the processes.

FIG. 2shows a typical hydrochlorination synthesis plant. MGSI242is conveyed into a low pressure hopper201through line248. A series of valves is manipulated in lines248and202in order to transfer the abrasive solid through line202into high pressure hopper203. Valves are then manipulated to transfer the abrasive solid through line204into the hot and high pressure FBR205. MGSI reacts with hot feed gases coming in through line217and FBR product leaves through line206, where it may go through interchanger207and line208, or go directly to quench vessel209.

The primary function of the quench209is to stop fine particulates of MGSI and metals salts such as FeCl3and AlCl3from moving further through the process train. (Fe and Al are impurities typically present in MGSI that from volatile salts in the FBR). This is accomplished by countercurrent flow of a dilute slurry (lines230,231and pump233) of the fine particulates in STC and TCS through a packed bed against the rising vapors. A portion of the dilute slurry is fed through line232to stripper234where it flows counter current against a H2stream228fed into the bottom of the stripper and leaving the top of the stripper as line229saturated with volatile components. This concentrates the silicon particulates that may have been carried out of the FBR205and also the volatile metal salts into a residue stream. The residue stream235comprising STC, TCS and solids slurry is finally discharged into a low pressure slurry handling and recovery unit236where some portion of the STC and TCS present are recovered via stream237and join the condensate stream247forming stream223. Solids and residual chlorosilanes plus impurities in stream238are treated for final disposal in unit239. A typical practice is to hydrolyze and neutralize stream238with neutralizing media240, leaving waste stream241for disposal as dictated by local practices and regulations.

Vapor leaves the top of quench vessel209through line210where some of the chlorosilanes in the vapor stream are condensed in condenser249. Some of the condensate219is refluxed via line220and some is drained via line246and ultimately fed to crude distillation column244where STC and TCS are separated. Vapors211leaving condenser249continue to a series of condensers represented by250. Typically, the terminal temperature in the series of condensers250is −25 to −50 C. The H2and relatively small concentrations of chlorosilanes leave condenser250in stream212where they are compressed in compressor243, flow through line213to heater247and then join a stream of vaporized STC227coming from STC vaporizer226. Stream225represents STC feed from other parts of the plant (CVD and/or silane operations) and stream252is a blowdown from the vaporizer.

TCS245leaving crude column244goes to further purification and from there to CVD operations1. STC224leaves the bottom of the crude column and is fed to STC vaporizer226where it is vaporized and the mixed with H2coming from heater247. This combined stream214is the fed through interchanger207(if an interchanger is used) through line215to trim heater216and finally into the bottom of the FBR as stream217.

From the description above, it is apparent that the process involves many steps. The combination of high temperatures, high pressures, abrasive solids and corrosive environment make the process equipment quite specialized and expensive. Valves in MGSI lockhoppers and associated lines used to charge the FBR wear out. The FBR and quench vessel are typically made of expensive high nickel alloys that require specialized fabrication techniques which are also expensive. Handling abrasive slurries in bottoms streams at high pressures requires specialized pumps and careful design. Even then, these components have wear limited lifespans.

Since temperatures over 500° C. are required to achieve desirable kinetics and equilibrium in the FBR, heating feed streams uses significant thermal energy. Electric heaters typically used for this service are expensive to purchase and/or have limited life. Heat integration from the FBR exhaust stream before the quench is significantly complicated and limited due to fouling from deposition of volatile chloride salts of iron and aluminum byproducts created in the FBR from their respective metallic impurities in the MGSI.

For the reasons above, the hydrochlorination process is a relatively expensive process to build and to operate. The size of equipment is largely dictated by the H2flow through the FBR. Difficulties associated with availability of materials like metal plate and fabrication difficulties with the alloys used limit the size of FBR's and to a lesser extent quench vessels. Even with FBR's built as large as possible based on fabrication capabilities and process understanding, world scale plants today have as many as 8 or more parallel hydrochlorination trains. This is far more than are required for reliability. Multiple units are necessary primarily because the capability to build larger trains does not exist and/or the incremental cost of larger trains is prohibitive due to fabrication difficulties. It is reasonable to expect technology will evolve and larger units of substantially the same design can be built in the future, but these size increases using parallel technology are expected to be somewhat incremental and without significant impact on the economics. A solution increasing capacity and/or capital cost per unit capacity is quite valuable to the industry.

Depending on efficiency of the process, energy costs can be either the greatest or second greatest cost in producing TCS in a polysilicon plant. Purchased MGSI is typically the other highest cost though economics vary from site to site. Reducing energy requirements are key to reduce operating costs.

FIG. 3shows detail for a typical STC hydrogenation plant3ofFIG. 1B. In this process, STC327from CVD reactors or a silane plant (stream326) and recycle stream325is fed to STC vaporizer329and then through line301where it is mixed with H2from stream302and fed through line303into STC converter304. The molar ratio of feeds is typically 2.0:1 to 3:1 H2:STC. In the STC converter approximately 18% of the STC fed is reacted to form TCS and HCL. The product stream305flows through a condensation train306consisting of one or more condensers and then through line307to a compressor308where the pressure is elevated to enable recycle of the H2and also to make subsequent condensation more efficient. After leaving the compressor, the gas flows through line309to condensation train310representing one or more condensers. Stream311then flows into a countercurrent HCl Absorber312where cold chlorosilanes313are flowing countercurrent to the rising gas coming from311. HCl and more volatile chlorosilanes are absorbed and exit through the bottom of the absorber via stream314. Stream314then flows to HCl stripper column315where the HCl is distilled away from the chlorosilanes via stream319and is directed to the direct chlorination plant4ofFIG. 1B. Chlorosilanes free of HCl leave the bottom of the HCl stripper in stream316. Part of the flow follows line321to crude column323where STC and TCS are separated. TCS324is fed to a purification train and then to CVD operations1ofFIG. 1. The remainder of the flow from the HCl stripper bottoms stream is fed through line322to absorber feed cooler318and then through line313into the HCl absorber. STC from the crude column is fed through line325where it is combined with STC326from CVD operations1, and fed to the STC Vaporizer. A small blowdown328from the STC vaporizer329purges impurities.

It should be understood that variations of this flow path utilizing various heat integration schemes are possible, and that the essential steps of an STC converter OGR used in industry are nearly universal. Those steps being condensation, compression, HCl absorbtion and HCl stripping. Many plants include a PSA (pressure swing adsorption) and/or TSA (temperature swing adsorption) step where the H2of stream302flows through a solid media to further adsorb HCl and other components from the H2. This TSA or PSA process consumes additional energy in that it must be heated and cooled (for TSA) or components are periodically purged into a low pressure stream (the PSA) process. The energy requirements in the TSA process are heating and cooling of the adsorptive media along with providing cooling and heating associated with the heat of adsorption and desorption. In the PSA process, either large amounts of purge gases are lost, or secondary recovery systems must be built to recover valuable gases from the purge.

The OGR in an STC hydrogenation process producing HCl byproduct is a significant capital investment and energy intensive part of the process. Refrigeration equipment is expensive and consumes large amounts of electricity to cool stream313feeding the HCl Absorber and run a condenser in HCl stripper319. Thermal energy required to strip HCl from the chlorosilanes in the HCl Stripper is also quite large. Since most STC converters operate at 4-7 Bar pressure and the HCl Absorber typically runs at 12-15 bar, significant energy is consumed in the compressor. Due to fairly low suction pressure, the compressors required are relatively large and capital intensive as well. When efficient high capacity converters as described in U.S. Patent Publication No. 2012/0328503 A1 are used, the cost of the OGR significantly exceeds the cost of the STC converters and most of the cost in the OGR is in the portions dedicated to removing HCl from H2and subsequently recovering the HCl in a relatively pure form. An OGR having lower capital and operating costs represents a high potential to reduce capital expenditure and energy consumption.

All patents, patent applications, provisional patent applications and publications referred to or cited herein, are incorporated by reference in their entirety to the extent they are not inconsistent with the teachings of the specification.

SUMMARY OF THE INVENTION

The present invention provides a process minimizing both capital and operating costs to make trichlorosilane (TCS) in a polysilicon plant by simplifying the STC converter OGR, improving heat integration, and combining STC converter and hydrochlorination processes in ways not previously envisioned despite many years practice of each process. To avoid the capital and operating costs to remove HCl from H2in the OGR of the STC converter based plant, the H2stream is flowed through a fluidized bed reactor to react substantially all of the HCl and a portion of the H2. This can be done using either a direct chlorination FBR or a hydrochlorination FBR, but is especially advantageous in a hydrochlorination FBR. Both processes are described. The net benefit of the hybrid hydrochlorination/converter process (the hybrid process) is a nominal 60% increase in the capacity of a hydrochlorination process having the same H2flow (ie. 1.6 times the capacity of a hydrochlorination process of the same size) as a traditional HC process while using approximately 25%-50% less energy per unit TCS produced. This can be retrofitted into an existing hydrochlorination plant at a very favorable capital investment cost for capacity expansion versus building either a new hydrochlorination train or an STC converter hydrogenation unit.

When coupled with a direct chlorination process, the size of the FBR must be increased and a means to recycle H2from the direct chlorination FBR incorporated. Given the significant modifications required, this is probably not attractive as a modification to an existing plant, but is an option for new construction.

DETAILED DESCRIPTION OF THE INVENTION

Throughout the description, where apparatus, compositions, mixtures, and composites are described as having, including, or comprising specific components, or where processes and methods are described as having, including, or comprising specific steps, it is contemplated that, additionally, there are compositions, mixtures, and composites of the present invention that consist essentially of, or consist of, the recited components, and that there are processes and methods of the present invention that consist essentially of, or consist of, the recited processing steps.

It should be understood that the order of steps or order for performing certain actions is immaterial so long as the invention remains operable. Moreover, two or more steps or actions may be conducted simultaneously.

It is contemplated that methods, systems, and processes of the claimed invention encompass scale-ups, variations, and adaptations developed using information from the embodiments described herein. Methods and processes described herein may be conducted in semi-continuous, and/or continuous operation. Reactors may be single-stage or multi-stage, and may be singular or plural without explicitly stating so It is contemplated that methods of the invention may be implemented in completely new facilities or combined or supplemented with existing reactors, systems, or processes that are known in the art such as fluidized bed reactor (FBR) processes used to make TCS from hydrogen chloride (HCl) and metallurgical grade silicon, or FBR processes to make TCS from hydrogen, STC and metallurgical grade silicon. Known, suitable techniques for separation of reaction products, recirculation of reactants, isolation and purification of reaction products, etc may be adapted for application in various embodiments of the claimed invention.

The mention herein of any publication, for example, in the Background section, is not an admission that the publication serves as prior art with respect to any of the claims presented herein. The Background section is presented for purposes of clarity and is not meant as a description of prior art with respect to any claim.

As used herein, a “halosilane” is understood to be a compound of the general formula RnSiZ4−n, where the radicals R are identical or different and are each hydrogen or an organic radical, such as an alkyl group CnH2n+1, and n is 0, 1, 2 or 3. Z is a halogen from the group Fluorine, Chlorine, Bromine, Iodine. When Z is chlorine (Cl), the halosilane is a chlorosilane. For example, in one embodiment the radicals R are —H, —CH3, or a combination thereof. In certain embodiments, each R is —H. In certain embodiments, the integer n is 0, 1 or 2. In certain embodiments, n is 0. In preferred embodiments, the halosilane is tetrachlorosilane (silicon tetrachloride, STC). In other embodiments, the halosilane is a bromosilanes, chlorodisilane or methyltrichlorosilane. Anywhere a chlorosilane is referred to, other halosilanes are considered equivalent. i.e. bromosilanes, fluorosilanes, or iodsilanes.

As used herein, “portion” means a part or all of the whole. A “part” as used herein means the whole from which components have been partially removed therefore effecting the relative concentration of those components in each part.

As used herein, interchanger is understood to mean a heat exchanger exchanging thermal energy between two process streams. Any heat exchanger can be an interchanger if heat is transferred between two process streams.

As used herein, countercurrent is understood to mean flows in opposite directions such that a hot fluid might enter one end of a heat exchanger flowing one direction while a cold fluid enters the opposite end of the heat exchanger and flows directly opposite the direction of the hot fluid. It may also mean, for example, a gas flows in substantially one direction (typically upward) while a liquid flows in substantially the opposite direction (typically downward) such as in mass transfer equipment comprised of trays and mass transfer packings.

As provided herein, an STC Converter capable of operation at pressures greater than 12 bar and preferably greater than 18 bar, and with a pressure vessel rating preferably equal to about 30 bar or more is installed in what would normally be the suction line212of the H2compressor243in a hydrochlorination plant ofFIG. 2. The new hydrogenation reactor preferably features a highly efficient heat exchanger having minimal gas residence time on the product gas side and a circular heating zone contained at the radial center and axial end of the hydrogenation reactor. After leaving the heating zone, gas flows around the heating zone and back into the interchanger where it heats the feed gas. In addition to the advantages over Siemens style hydrogenation reactors, this design also has significant advantages over other designs referenced by patents previously cited in simplicity, ease of fabrication, maintenance, capacity, and cost.

FIG. 1is representative of major process steps within most existing closed loop polysilicon production facilities. It is a summary block flow diagram of entire polysilicon plants. It shows three configurations of known commercially significant plantwide configurations. Additional operations could be defined within a polysilicon facility and minor recycle streams may exist that are not shown. Details not shown are not highly relevant to this invention and are omitted for brevity and ease of understanding essential concepts.FIGS. 1A, 1B, and 1Care previously described in some detail within the background section of this document.

FIG. 2is a representative schematic diagram of the hydrochlorination TCS plant ofFIG. 1Asubstantially the same as that described in U.S. Pat. No. 8,298,490 B2. For hydrochlorination technology, it shows the major unit operations associated with synthesis of TCS and STC. Separation of a crude TCS product is shown, but additional purification steps as well as minor recycle streams are not necessarily shown as they are not essential to this invention.

FIG. 3is a representative schematic diagram of the STC hydrogenation portion of a direct chlorination plant ofFIG. 1Bas practiced in industry. It shows the key steps to hydrogenate STC to TCS and the associated OGR to recover TCS and HCl products from the hydrogenation reaction and to recycle unreacted H2and STC streams. It shows separation of a crude TCS product. Additional purification steps as well as minor recycle streams are not necessarily shown as they are not essential to this invention.

FIG. 4is a representative schematic diagram of the hybrid process of this invention where an STC converter462(or STC converters) and simple condensation train464are installed in series with the hydrochlorination FBR, its quench, and condensation train. Like numerals withFIG. 2can be considered substantially the same as those ofFIG. 4. The hybrid process ofFIG. 4uses the same H2compressor243asFIG. 2displacing substantially the same amount of H2gas. InFIG. 4, streams413,414,415and417are essentially the same as respective streams213,214,215, and217ofFIG. 2except that the four 4XX streams will have a higher concentration of HCl (˜6 vs ˜0.1 mole %) versus the four respective 2XX streams ofFIG. 2. These four streams are otherwise similar. Streams424and445are substantially the same function and composition as224and245respectively.

Streams461,463,465, and466(STC converter related streams; STC feed, combined feed, exhaust, and exhaust condensate) serve largely the same functions respectively as streams301,305,320, and303. Pressures in the 4XX streams are greater than typical practice for the 3XX streams and compositions are slightly different. Equipment426is an STC vaporizer feeding STC vapor to both the FBR205and the STC converter(s)462. Stream470is a distillation vent.

Key advantages of the hybrid process are as follows;1) The capacity of a hydrochlorination process can be increased approximately 60% or more with investment only in a larger (or supplemental) STC vaporizer, an STC converter, and a fairly simple and low cost condensation train following the STC converter. The cost of this is estimated to be substantially less than 60% of the cost of an entire hydrochlorination FBR process built from scratch.2) The dominant reaction in the direct chlorination FBR
3HCl+Si→HSiCl3+H2is very exothermic. The net reaction in a hydrochlorination reaction,
3HSiCl3+2H2+Si→4HSiCl3is slightly endothermic. By constructing the hybrid process where both reactions occur simultaneously in the hydrochlorination FBR ofFIG. 4, temperatures in the FBR will be approximately 20-25 C higher than the feed temperature versus 20-25° C. lower than the feed temperature ofFIG. 2representing a significant cost reduction to heat the feed stream in the hybrid process.3) Trim heaters216typically have a limited life impacted significantly by operating temperature. The ability to operate them at lower temperature while keeping the FBR temperature the same or higher increases the lifespan of the heaters.4) Operating the hydrochlorination FBR at higher temperatures with the same or lower trim heater temperatures is possible due to the exothermic reaction. Increasing FBR temperature will increase the conversion of STC to TCS in each pass, increasing capacity by more than the approximately 60% previously stated.5) The hybrid process enables opportunities for heat integration not present in either process individually, further reducing the energy costs. Examples of heat integration steps are detailed onFIG. 5.6) In sites with scarce space available for new construction or expansion, the hybrid process provides an opportunity consuming considerably less ground area.

FIG. 5is a schematic diagram of a hybrid hydrochlorination/converter process showing key heat integration steps enabled by combining a process with STC converter462and hydrochlorination FBR205. Saturated vapor stream210leaving quench tower209is at a high enough temperature that it can be used to provide a major portion of the heat load required to vaporize the STC in stream519. This is possible because the pressure in stream210is higher than stream466and because the desired ratio of H2:STC in stream466is lower than in stream210. After stream210flows through coil516in STC vaporizer506it flows through line501to interchanger502, through line503to condensation train504consisting of one or more heat exchangers, and back through interchanger502in stream507. Pressure reducing valve514between lines509and511controls the feed pressure to the STC vaporizer506, thus controlling the H2:STC ratio in stream466. Condensate from504flows into streams223and220. A supplemental heating coil520can be used to provide operating flexibility between composition and pressure in stream466.

Gas leaves STC converter462and flows through line463to interchanger508where it heats stream213. Cooled stream523flows to condensation train510consisting of one or more heat exchangers and then through line525to compressor243. Small amounts of STC and TCS remain in stream213that is predominantly H2and HCl. Heated stream513flows into STC vaporizer512which has supplemental heating from heat transfer media523flowing through heating coil518and out through stream525. STC is fed to vaporizer512through line521. Saturated stream515flows through heater514and then to the FBR as previously described forFIG. 4. With thoughtful design practices known to those skilled in the art, much of the heat required to vaporize STC in vaporizer512will be provided by the sensible heat in stream513as it cools to the saturation temperature of stream515. Streams517and522are blowdowns to remove highboiling impurities from the process.

FIG. 6is a representative schematic diagram showing details of the direct chlorination TCS plant4ofFIG. 1B. The STC hydrogenation process3shown inFIG. 1Bis detailed inFIG. 3. InFIG. 6, MGSI626is fed to direct chlorination FBR601and HCl gas624is fed to the bottom of the FBR601. The reaction of HCl624with MGSI626to make TCS and STC is highly exothermic and heat is removed with heat transfer media628through cooling coils or equivalent in FBR628. Essentially all HCl fed in stream624is reacted in FBR601. Reaction products TCS, STC, and H2flow through line602to quench operation603. The configuration and function of quench603is largely the same as for vessels209,234and249of the hydrochlorination process ofFIG. 2, namely to remove silicon solids and volatile salts which are removed in stream610. Some condensed TCS and STC leave the quench through stream612and go to distillation column607where TCS and STC are separated. H2and uncondensed chlorosilanes leave the quench operation in stream604to enter condensation train605consisting of one or more heat exchangers. Condensate608flows to distillation in line627. TCS from distillation630goes through a purification train609where impurities616are removed and then to CVD operations1via stream632. In the traditional direct chlorination process, H2is vented to atmosphere through stream606and is lost. The direct chlorination FBR601, its quench603and condensation train605typically operate at 2-7 bar pressure, thus requiring colder condensation temperatures in the condensation train to achieve the same percentage recovery of chlorosilanes than is required in the hydrochlorination process TCS separated in distillation column607is fed to a purification train and finally to CVD operations1. STC from column607leaves through line614and go to the STC hydrogenation operation3. Sub-processes within CVD operations1are the same as has been previously described. STC618and H2byproducts620are recycled to STC hydrogenation3. TCS622from hydrogenation is fed to CVD operations1.

FIG. 7is a schematic diagram of the process using direct chlorination and STC converters without separating HCl from the H2produced in STC converters. Unlike the hybrid process involving STC converters and hydrochlorination where STC converters can be added into an existing hydrochlorination loop ofFIG. 2with very minimal modifications, this process is very different from a typical direct chlorination TCS plant and the design is believed practical only if built as a new facility. In this process, H2and HCl in stream718and MGSi626are fed to direct chlorination FBR701operating at approximately 330° C. and 7 barG. The H2will go through the FBR largely unreacted, but substantially all of the HCl present will react to form STC and TCS in essentially the same proportions as in FBR601. The FBR must be larger diameter to deal with the increased volume of gas flow due to the H2in stream718. Heat removal from FBR701by heat transfer media will be slightly less than in FBR601because the H2flowing through the reactor will take out some of the heat. Gas stream702leaves FBR701to a quench operation703, then to line705and condensation train706whose design and function is similar to Quench603and condensation train605. Vapors leave condensation train706to be joined by H2produced in CVD operations and feed compressor711via line710. Impurities are removed in line610. Temperatures must be lower in condensation train706than condensation train605due to the much higher concentration of non-condensible H2. Condensate streams704and707combine in708and with condensate719from the condensation train717of the STC converter715. Combined stream720flows into distillation column721which separates TCS and STC. STC leaves column721through line725, is joined by STC coming from CVD reactor ops1in stream726. The combined flow727then goes to STC vaporizer728. The STC is vaporized into stream713and joins the H2in stream712coming from compressor711. The combined flow is stream714flowing into converter715.716is the exhaust stream from converter715containing TCS and HCl along with unreacted H2and STC fed to STC converter715in stream714. Condensation train717condenses substantially all of the chlorosilanes from the H2in stream716into stream719which then flows to distillation column721. Vapor stream718leaves condenser717and flows back to the FBR701. H2stream712can be preheated in an interchanger with stream716(not shown) and fed to STC vaporizer728to reduce heat load on the vaporizer728and condensation train717. TCS722leaves distillation column721to go through purification train723prior to being fed to CVD operations via stream724.

Constructive Examples

Standard engineering techniques using chemical engineering simulation software widely used in industry (Chemcad, Version 6, provided by Chemstations, 3100 Wilcrest Drive, Suite 300, Houston, Tex., USA and Aspenplus, Version 8, provided by AspenTech, 20 Crosby Drive, Bedford, Mass., USA) were used to model and compare the energy consumption per unit TCS produced in a traditional TCS plant associated with polysilicon production. The processes ofFIG. 5,FIG. 2, andFIG. 3are compared (in these comparisons, the energy for separation of STC from TCS by distillation in equipments244and323are not included). Physical properties, thermodynamic models, and equations of state known by those skilled in the art to be sufficiently accurate for industrial applications were used. The industrial range stated below is based on the inventor's industry experience of designs and operating parameters associated with current industrial state-of-the-art processes ofFIGS. 2 and 3and are consistent with prior art descriptions of this document.

It is understood that the foregoing examples are merely illustrative of the present invention. Certain modifications of the articles and/or methods employed may be made and still achieve the objectives of the invention. Such modifications are contemplated as within the scope of the claimed invention.