WO2005069965A2

WO2005069965A2 - Method and composition for treating sour gas and liquid streams

Info

Publication number: WO2005069965A2
Application number: PCT/US2005/002038
Authority: WO
Inventors: David Dutchover, Iii
Original assignee: Paradigm Processing Group Llc
Priority date: 2004-01-23
Filing date: 2005-01-21
Publication date: 2005-08-04
Also published as: WO2005069965A3

Abstract

Design and use of an aqueous solvent solution adapted for use in the removal of contaminate such as, acid gases and sulfur, from a feed stream, the design using computer modeling of a mass transfer unit based on mass transfer rate calculations, including a product preferably formed by adding sodium silicate, water, and at least one alkanolamine selected from the group consisting of monodiethanolamine (MDEA), diethanolamine (DEA), monoethanolamine (MEA), diisopropanolamine (DIPA), triethanolamine (TEA) and similar compounds. The solution can be formed by combining from 1 to about 99% by weight, and more preferably from 5 to 50%, sodium silicate with water, and combining the resulting sodium silicate solution with the desired alkanlolamine(s) at a ratio of from about 1:20 to about 1:1. A novel improved mass transfer unit is disclosed.

Description

Title: Method and Composition for Treating Sour Gas and Liquid Streams

CROSS REFERENCE TO RELATED APPLICATIONS This application references, incorporates by reference herein and claims priority based on U.S. provisional application 60/538,746 filed 1/23/04 of same title. FIELD OF THE INVENTION The present invention relates to the design of and methods of use of chemical compositions for the removal of in particular acid gases (including but not limited to H₂, CO₂, S,) as well as sulfur and sulfur containing compounds and COS and other contaminates, from a gas or liquid stream, and to the removal of certain heavy metals from a liquid stream. Certain embodiments relate to the removal of compounds from a fluid stream by contacting the fluid stream with a composition containing sodium silicate, an alkanolamine and water, the composition tailored to process parameters and requirements, the tailoring a function of mass transfer rate calculations in an absorber/adsorber. The invention includes the design of a novel improved mass transfer unit and preferred system equipment. BACKGROUND OF THE INVENTION The use of alkanolamines for removing acid gases such as CO and H₂S from gas streams is known in the art. Typically, the gas stream is contacted with an alkanolamine in an absorber or contact column. [Such contact unit is commonly, alternately referred to as an absorber and/or adsorber, a contact tower or column, a mass transfer unit (MTU) or an intimate contact unit (ICU).] The aJkanolamine may be in an aqueous solution which is typically passed through a desorber or other contaminant removal system and recycled back to the absorber or contact column. Although many improvements have been made in this process, and in the alkanolamines that are used, various problems remain that reduce the effectiveness and efficiency of conventional systems. One common problem is that of solvent foaming, in which the liquid alkanolamine solution and the gas that is being treated do not adequately separate, resulting in a froth or foam that fills the contact chamber and prevents operation. Corrosion of the process equipment is another common problem. Stability of the removal fluid can be an issue. In addition, some of the alkanolamines that are used in the art have an affinity for the hydrocarbon stream that is being treated, with the result that some hydrocarbons dissolve in the solvent stream, which is undesirable. Another problem is that many conventional amine solutions require elevated temperatures and/or pressures in order to ensure that the solubility of sulfur, sulfur compounds or other contaminants that are to be removed is sufficient in the stripping solvent. Still another problem in conventional systems results from the accumulation of heat stable salts. These factors typically require complex and expensive equipment to ensure effective operations. Originally, monoethanolamine (MEA), a primary aJkanolamine, was the standard compound used in the removal of acid gases. To prevent excess corrosion, the MEA concentration in the solvent had to be kept relatively low. This in turn required high solvent circulation rates, which resulted in high energy requirements. MEA also required periodic purification to remove degradation products. Purification usually involved the continuous thermal distillation of a small side-stream of the MEA. This side treatment maintained the MEA in an acceptable operating condition, but required an additional amount of energy to operate. In addition, the bottoms from the reclaimer represented a significant loss of MEA as well as a hazardous waste that is difficult to dispose of, both of which added to the operating expense. As energy and capacity requirements increased, another solvent, diethanomlamine (DEA), began to be used. Since DEA is a secondary alkanolamine, it is more stable, less reactive, and potentially less corrosive than MEA, and can be used in higher concentrations. The use of DEA increases the capacity of the solvent and decreases the overall energy requirement. However, it has been found that DEA does not sufficiently avoid the problems associated with MEA. It is also known that tertiary amines, such as methyldiethanolamine (MDEA), do not undergo the same type of chemical and thermal degradation that primary and secondary amines undergo. As a result, tertiary amines are much more stable in gas treating processes, and MDEA is becoming increasingly acceptable as a replacement for MEA and DEA. Not only is MDEA more stable than MEA and DEA, it is potentially less corrosive and can be used in even higher concentrations. Higher amine concentrations also allow increased capacity and lower energy requirements. Unfortunately, the reactivity of C0₂ with MDEA is much slower than with MEA and DEA. As a result, MDEA alone cannot be used in applications that require almost complete CO₂ removal. To get around this problem, DEA or MEA are sometimes added to MDEA to improve solvent reactivity. Such solvent mixtures have increased capacity and lower energy requirements than MEA or DEA by themselves, but are less stable than MDEA alone. Solvent degradation can make the benefits of using MDEA/MEA or MDEA/DEA mixtures uneconomical. It is desirable to provide a system for removing contaminants, and H₂S in particular, from gaseous and liquid hydrocarbon streams, and from natural gas streams in particular, that would be less temperature- or pressure-dependent than conventional systems, that would allow high throughput, avoid fouling and foaming, and that would be commercially cost- competitive with existing systems. An improved absorber/adsorber, or MTU, is also disclosed herein which amounts to a paradigm shift in regard to the traditional contact tower and tray approach. The novel mass transfer unit can offer orders of magnitude greater surface area in which the mass transfer can take place. The volume required for the novel mass transfer unit is orders of magnitude less than that of the conventional packed tower. The unit is scalable and promises essentially instantaneous chemical physical equilibrium and efficient gas to liquid ratios. Most importantly, the interfacial surface area is dramatically increased while the unit remains robust, reliable, safe, clean, inexpensive and easy to control. SUMMARY OF THE INVENTION In certain embodiments, the present invention comprises a method of use for an designed aqueous solvent solution formed by combining at least an alkaline silicate salt and water, the solvent solution composition designed by computer modeling using mass transfer rate calculations based on an anticipated feed stream and process equipment requirements. In many cases the designed solvent solution will include an alkanolamine. The designed aqueous solvent solution, in general, is less temperature or pressure dependent than conventional contaminate removal system solutions, allows higher throughput, avoids fouling and foaming to a greater extent, and is commercially cost-competitive with existing systems. In addition, the designed solutions are effective corrosion inhibitors, making them particularly effective for the removal of acid gases. In embodiments adapted for sulfur and sulfur compound removal, the present invention preferably features a designed a&anolamine composition that is formed by combining at least a sodium silicate and one or more alkanolamines dissolved in water. Surprisingly, the designed aqueous solutions have been found to provide both improved contaminant removal and reduced corrosion as compared to prior art commodity solvents, including those having the same concentration of alkanolamine. The designed solutions can incorporate a greater loading of contaminates than prior art commodity solvents. In addition, the present designed solutions are much less receptive to hydrocarbons than prior art solvents and increase the operating capacity of existing equipment by allowing greater throughput of the product streams. In certain embodiments, the present invention comprises the design of a tWophilic aqueous solution adapted for use in the removal of acid gases comprising H₂S, CO₂, COS, as well as sulfur and/or sulfur containing compounds or mixtures thereof, from gaseous or liquid streams. In these embodiments, the solution is designed by combining at least sodium silicate, water, and at least one alkanolamine selected from the group consisting of monodiethanolamine (MDEA), diethanolamine (DEA), monoethanolamine (MEA), dnsopropanolamine (DIP A), hiethanolamine (TEA) and similar compounds. The designed solution can be formed by combining from 1 to about 99% by weight, and more preferably from 5 to 50%, sodium silicate with water, and combining the resulting sodium silicate solution with the desired alkanlolamine(s) at a ratio of from about 1:20 to about 1:1. Known antifoamers, stabilizers, anticorrosion agents and surfactants can be modeled and added to the designed solvent solution, depending on the feed stream and the process conditions and requirements. In other embodiments, the present invention provides a process for the removal of acid gases comprising H₂S, CO₂, COS, or mixtures thereof, from a gaseous or liquid stream by contacting the stream with an aqueous amine solution in a contact zone, where the aqueous amine solution including at least an alkaline silicate, and preferably includes sodium silicate, and an alkanlolamine, the solution designed in accordance with a modeling process based at least in part on mass transfer rate calculations. The contact zone may include a trayed or packed column, or other MTU equipment particularly suitable for effecting the desired contact between the streams. The performance aqueous solution composition as well as the equipment and operating parameters are preferably designed for optimal performance through a modeling program using, at least in part, mass transfer rate calculations. Thus, the present invention comprises a combination of features and advantages which enable it to overcome various problems of prior art devices. The various characteristics described above, as well as other features, will be readily apparent to those skilled in the art upon reading the following detailed description of the preferred embodiments of the invention, and by referring to the accompanying drawings. The summary includes a mass transfer unit wherein acid gases or sulfur containing compounds can be removed from gaseous or liquid streams in a more efficient and cost effective manner. The mass transfer unit has application not only to stripping and absorption but to distillation and concentration processes. BRIEF DESCRIPTION OF THE DRAWINGS A better understanding of the present invention can be obtained when the following detailed description of the preferred embodiments are considered in conjunction with the following drawings, in which: For a more detailed description of the preferred embodiment of the present invention, reference will now be made to the accompanying drawings, wherein: Figure 1 is a schematic diagram of a system for removing contaminants from a sour feed stock; such as natural gas. Figure 2 is a schematic diagram of a system for direct treatment of a sour product; such as a NOX and SOX treatment system. Figure 3 is a schematic diagram of a system for removing sulfur fromtailgas. Figures 4A-4S illustrate the calculation involved in modeling mass transfer with chemical reactions, including distinctions between types of reactions. Figures 5A-5D illustrate the size dimension and structure of a mass transfer unit. Figure 5E illustrates a possible cavitation phenomena. Table I illustrates MTU process parameters as a function of diameter in length. The drawings are primarily illustrative. It would be understood that structure may have been simplified and details omitted in order to convey certain aspects of the invention. Scale may be sacrificed to clarity. DETAIL DESCRIPTION OF THE PREFERRED EMBODIMENTS SYSTEM Referring initially to Figure 1 , a generic system for carrying out one embodiment of the present invention includes a feed line 10, an absorber 20, a crystallizer/settler 30, a contaminate filter system 40, and a solution recycle line 50. The feed line 10 may pass through an inlet conditioning unit 12 before entering absorber 20. Contaminated solution from the bottom of absorber 20 exits via line 22 and enters crystallizer 30. If desired, an additive can be added to the solution in crystallizer 30. An exit stream 34 from the bottom of crystallizer 30 is pumped either into contaminate filter system 40 or through a cooling unit 34 and back into crytallizer 30. Filtrate leaves contaminate filter system 40 via line 42 and is preferably recycled back into crystallizer 30, while filter cake containing contaminants removed from the stream is intermittently removed from contaminate filter system 40 for disposal. The inlet gas in feed line 10, may be sour gas (e.g., natural gas, refinery fuel gas, HDS recycle gas, EOR gas streams, etc.) or sulfur tail gas. The inlet gas preferably enters the process through a gas conditioning step typically that is specific to the application. For direct treat applications, this gas conditioning step represents the addition of SO₂ to the process through either 1) injection of liquid SO₂; 2) burning of contaminate to produce SO₂; or 3) thermal or catalytic oxidation of H₂S from the inlet gas stream to produce a 2:1 H₂S:SO₂ ratio. For suJiur tail gas stream, (See Figure 3) this gas conditioning step is integrated into the bottom of the absorber as a quench system utilizing a recycle stream comprising a silicate containing design solution in accordance with the present invention. As the inlet gas, which in a sulfur removal process is relatively rich in sulfur compounds, passes upward though absorber 20, it closely contacts a solvent solution. The solvent preferably has an affinity for sulfur or sulfur compounds, so that the sulfur in the inlet stream tends to move into the solvent stream. Absorber 20 can be either a trayed or packed column, traditional absorber and can contain other peripheral equipment as necessary for optimum operation. Such equipment, known in the art, may include, optionally, an inlet gas separator, a treated gas coalescer, a solvent flash tank, a particulate filter, and a carbon bed purifier. Conventional absorber 20 preferably has a countercurrent configuration, such that the sulfur content of the gas stream is substantially reduced by the time the gas exits absorber 20, resulting in a relatively lean gas stream. Alternatively, flow through absorber 20 can be co-current or any other configuration. If the inlet gas is at pressure, then the solvent solution is preferably flashed down to near atmospheric pressure, producing a small flash gas stream that can either be recycled or used as fuel. For near atmospheric pressure applications such as sulfur tail gas treatment, no flash tank is required. Sulfur-rich solvent is removed from the bottom of absorber 20 and passed to crystallizer 30, where the sulfur, and/or other contaminants are flocculated, precipitated, agglomerated, or otherwise rendered separable from the liquid. This treatment returns the liquid solvent to a contaminant-lean state. In certain embodiments, settler/crystallizer 30 includes a cooling loop and the solution is retained for a calculated residence time that allows the formation of crystals. The solvent solution does not have to be cooled to any particular temperature for the crystals to form, but rather just enough so that the crystallizer temperature is sufficiently below the absorber temperature, so that the solid contaminate forms in the crystallizer rather than elsewhere in the system. For example, the absorber temperature could be 150°F, the crystallizer could be 120°F, and ambient air could be used as the coolant. Solution that is lean in contaminants is removed from crystallizer 30 and recycled to absorber 20 via line 50. If desired, heat is removed via heat exchanger 38. In an alternative embodiment (not shown), the filtrate in line 42 is added to the lean solution and recycled to absorber 20. In still another embodiment, line 36 is eliminated and all of the recycled solution passes through filter system 40. The slurry of crystalline product from the crystallizer is fed to a filter or centrifuge that produces a filter cake of elemental product for disposal or sale. Contaminate formed using the designed solutions of the present invention has been found to have a significantly larger crystal size and much better handling properties than contaminate formed from prior art solutions. Various known flocculating technologies can be used in conjunction with the present technology, to aid in removing the contaminate and regenerating the solvent solution. An example of a suitable technology is available from Eco Nova Corp. A low-boiling wash solvent may be used to wash the dry filter cake and remove residual solution from the contaminate. The solution/wash solvent stream from the filter may then be fed to a small, skid-mounted, solvent recovery still for separation. A final rinse with water preferably completes the wash cycle. The recovered solution, the wash solvent, and the water are all recycled and reused repeatedly. One of the primary differences between the present designed solution solvents and aqueous iron processes of the prior art is the contaminate formation mechanisms. Particles formed using the present design solutions range from 10 to 200+ microns in size, and have a very crystalline structure. The crystals contain few inclusions and are easy to wash, producing very clean contaminate product. The crystals settle on their own; no surfactants or wetting agents are used. In contrast, contaminate particles from aqueous-iron systems range in size from maybe 1-10 microns. The large clumps are simply many tiny particles held together by surfactants. The surfactants must be added to the system to allow the contaminate to sink and keep it from floating on the surface of the aqueous liquid. These loose agglomerates, in addition to containing a significant amount of surfactant, tend to be sticky and difficult to clean. Further, managing the additive levels can be difficult and ultimately lead to many of the plugging and foaming experiences that can occur in these processes. The preferred rate of inlet gas flow in the instant invention will depend on the size of the equipment and may be such that the equipment operates with a solution flow rate in the range of 400-600 gallons per minute. As is known in the art, an amine solution circulation rate depends on the amine concentration, the gas flow rate, gas composition, total pressure and treated gas specification. Nonetheless, it has been discovered that, all other parameters being equal, the use of compositions and design solutions in accordance with the present invention allows amine concentration to be increased, resulting in a significant increase in system operating capacity. An arnine solution circulation rate will typically be between 5 and 5000 gallons per minute (gpm) or more. By way of example only, a system using the present design solvents and apparatus can operates with a circulation rate of roughly 20 gpm per long ton of contaminate removed per day. In comparison, the available aqueous Iron Redox Processes circulate 1,000 to 3,000 gpm per long ton of contaminate removed per day. Since the present system allows a much lower fiowrate, the amount of hydrocarbons absorbed from the inlet gas is minimized, to a point as is by chemistry terms considered to be insignificant. SOL VENT COMPOSITIONS According to certain embodiments of the invention, the preferred solvent solution used in such systems comprises an aqueous solution formed by adding sodium silicate or potassium silicate and, optionally, one or more alkanolamines and water. More specifically, the solution preferably is formed by combining about 1 to 50% percent by weight of an alkaline silicate salt, such as potassium or sodium silicate, about 1 to 50% percent by weight alkanolamine, and the balance water. In certain preferred embodiments, the solution is formed by combining between about 10 and about 30 wt % sodium silicate, between about 10 and about 30 wt % alkanolamine, and the balance water. Depending upon design parameters, small amounts of commercial antifoamers, stabilzers, anticorrosive agents and surfactants may be added. The primary non aqueous component of the present solutions is preferably selected to have a high solubility for contaminate so that solution circulation rates can be low, resulting in smaller volumes and equipment sizes. All of the components of the present design solutions preferably have low volatility and high chemical stability. The alkaline silicate salt preferably comprises sodium silicate in a ratio of SiO₂ to Na O of between 1:3.22 and 1:1.80. One commercially avialable composition has a ratio of SiO₂ to Na₂O of 2.0 and is sold by Occidental Chemical Corp. under the trade name Oxychem Grade 50. wniie rne au aune silicate salt may be added to form the present solutions at about up to about 99% percent by weight, it is more preferably added at about 0.01 to about 30 wt%, more preferably 0.01 to about 12 wt % and still more preferably between about 0.01 and 5 wt%. Exemplary formulations include: Solvent #1 : A 10% to 90% of sodium silicate blended with a 10%-90% of MDEA. Solvent # 2: A 10% to 90% of sodium silicate is blended with a 10% to 90% of MDEA and 1% to 9% of DEEP. Solvent # 3: A 10% to 90% of sodium silicate is blended with a 10% to 90% of MDEA and 1% to 20% of DIPA. Solvent # 4: A 10% to 90% of sodium silicate is blended with a 10% to 90% of MDEA and 10% to 90% of DEA. Solvent # 5: A 10% to 90% of sodium silicate is blended with a 10% to 90% of DEA. Solvent # 6: A 10% to 90% of sodium silicate is blended with a 10% to 90 % of MDEA and a 1% to 20% of TEA. Solvent # 7: A 10% to 90% of sodium silicate is blended with a 10% to 90% of MDEA and 5% to 50% of MMEA. Solvent # 8: A 2% to 90% of sodium silicate is blended with 2% to 98% H₂O. Solvent # 9: A 2% to 90% of sodium silicate is blended with a 2% to 98% H₂O and a 1% to 20% of a compound to lower the melting point of the aqueous solution, such as methanol or any of the glycols. Preferably the designed solvent for a specific operation is specified by a computer modeling of the removal system, including the characteristics of the feed system, the equipment and the result requirements. Small amounts of commercially available additives may be included such as antifoamers, stabilizers, anticorrosive agents and surfactants. The following Table illustrates some amine variations of the present solvent compositions, the composition understood to also contain a base alkaline silicate, the compositions expected to incorporate the benefits of the present invention when designed and tailored further for a variety of circumstances, along with preferred modes of application for each.

OPERATION CONDITIONS The operating pressure and temperature that are required at a contactor/absorber in order to allow removal of the acid gas components from a hydrocarbon stream depend in part on the efficacy of the solution as a removal agent and in part on the solubility of the various compounds in the aqueous and hydrocarbon streams. In general, operating pressures found within an oil refinery are less than 500 psig, which does not allow the necessary reaction between the conventional hydrocarbon treating solvent and the acid gas components to take place. Thus, in conventional H₂S removal systems, it is necessary to operate the contact unit at elevated pressure and temperature. In contrast, the highly reactive nature of the present design of aqueous solvent solutions does not require elevation of either pressure or temperature in the contact unit. Similarly, current technologies allow for a rnaximum acid gas loading of 0.45-0.50 mole of acid gas comprised of hydrogen sulfide, CO₂, COS and mercaptans to mole of amine. By contrast, the present designed solution compositions have successfully achieved a rich loading of solvent to a 0.71 mole acid gas per mole amine. This increased solution capacity results in greater system capacity. In some embodiments, pressure inside absorber 20 can be between 1.0 and 1200 psig, depending on the type of gas being processed. Temperatures inside absorber 20 are preferably between 80°F and 120°F. Amine Unit Operation With Solvents Before any consideration can be given to an amine unit operation, preferably the correct solvent has to be chosen and designed to meet the desired process requirements in order to provide the most efficient operation. Once me proper solvent has been chosen, and designed a good understanding of operating guidelines or theory of operation with the selected designed solvent can be realized. Choosing and Designing the Correct Solvent. Most processing requirements differ with each individual application. Consequently, commodity solvents are limited in their ability to operate efficiently, as they cannot be tailored to fit a specific application. Often operational inefficiencies and problems result when attempts are made to use a commodity solvent in an application not suited to its properties. The present design solvents are specifically designed to fit each individual application. Consequently, the desired processing requirements can be met in the most efficient manner. Selection and design of the proper solvent is preferably achieved via sophisticated computer modeling of the specific process conditions and requirements. Computer Simulation for Sour Gas/Liquid Product Treatment. An absorber/adsorber model is used to simulate gas sweetening and liquid treating, and the solution is compounded and designed based on the composition of the feed stream(s). The weight percent of the components in the solution, including silicate, water, and optionally alkanolamine, and/or other common additives is determined for a specific inlet composition. This simulation is based on a sophisticated mass transfer rate approach and uses actual trays, which do not need empirical staging efficiencies. In the gas processing industry, a theoretical Equilibrium Stage Approach has generally been used in the design and simulation and gas sweetening absorbers. Some type of tray efficiency typically modifies these models, which is determined from process data. In general, efficiencies depend on several variables including the contact time (gas and liquid rates, weir height), physical properties (viscosity, density, difjusivity, etc.) of the two phases, interfacial area (bubbling area, type and diameter of tray), and driving forces. Some of these factors vary from tray to tray even in an absorber. Thus, the common practice to use a single value for the overall efficiency is generally inadequate to accurately describe the process. Although one can possibly fit a single efficiency for each component to existing data, it is difficult to develop sound design for rating procedures based on these efficiency factors as needed by the Equilibrium Stage Approach. This problem worsens in the case of simultaneous absorption (as in the case of gas sweetening) due to strong interaction between the equilibrium and behaviors of the two gases. In contrast, a Rate Approach Model directly calculates the rate of absorption and/or adsorption, including on real trays. This approach, utilizing the types of calculation and considerations ϊliustrateα m ϊigures 4A-4S, involves calculating the driving forces, mass and heat transfer coefficients, interfacial area, and interaction efficiencies between the mass transfer and the reaction phenomena. Thus, a Rate Approach must have several theoretical models, such as vapor for/liquid Equilibrium Model, kinetic reaction model, hydrodynamic model, and a model to describe the effect of reaction on mass transfer. In addition, one needs to integrate the "point" behavior within an absorber to predict the performance of the entire column. In chemical absorption systems is also important to note that mass in energy balance interacts wrongly. Temperature affects the reaction rates in equilibrium behavior of the system hence the mass transfer rates. On the other hand, the mass transfer rates deteπnine the heat effects. All of the variables must be considered simultaneously in order to accurately predict an absorber performance. The preferred simulation of the instant invention predicts this complex behavior on each tray. This includes the liquid loading and temperatures, gas compositions and temperatures, strain flow rates, gas and liquid side mass transfer coefficients for each component, interfacial areas, heat transfer coefficients, physical properties, thermal dynamic constants, etc. As the gas treating industry evolves into an era of formulated solvents the need for a rigorous simulator becomes increasingly important. Equilibrium Models can no longer provide the estimates needed for plant design or rating. The present preferred model amine simulation performs these rigorous, simultaneous calculations. This simulation technique has been successfully tested against many actual plant data sets covering a wide range of conditions and several differently formulated solvents. The present preferred simulation, based at least in part on a mass and rate transfer approach, has been developed for evaluating commercial sour gas treating absorbers as well as liquid treaters commonly found in refineries. This model predicts the overall performance and the acid gas loading and temperature profiles for an absorber on a tray-by-tray basis. The simulated results have been verified by rrøtching actual plant loadings and temperature profiles. This simulator quantifies the effect of design and operating variables and provides a unique predictive tool for rating, design, and plant optimization purposes. General Operating Guideline_^ Due to the unique differences in properties in practice between the present design solvents versus the commodity solvents, differences in operating practices also exist. By following a few guidelines, successful performance and enhanced benefits from the present design solvents can be realized. Preferred operating parameters are listed as follows: • Solvent Concentration • Lean Solvent Temperature • Circulation Rate • Regenerator Conditions • Reboiler Conditions Solvent Concentration: The present design solvents operate most efficiently at a 50-60-wt% concentration (+ or - 5 wt%). MEA concentration is preferably limited to 15-20-wt%, and, DEA to 30-wt% concentration due to corrosion concerns. Therefore, substantially more capacity exists with the present 50-wt% amine concentration vs. MEA and DEA, even when the lower molecular weight of MEA is considered. Consequently, conversion to one of the present design solvents can create additional processing capacity and/or reduced energy requirements due to a lower circulation rate. Furthermore, at the 50-60 wt% concentrations, equilibrium rich acid gas loadings are lower with the performance solvents versus commodity amines. Consequently, the potential for corrosion in the rich solvent areas of the plant is reduced. In addition, the lower heat of reaction associated with the present solvents will produce lower rich solution temperatures, which further reduce the potential for corrosion. Operating above the 60 wt% concentration can lead to viscosity related problems such as inefficient acid gas absorption, increased hydrocarbon entrainment, poor heat transfer, and higher pumping requirements. Operating below the concentration can result in increased circulation rate and higher reboiler duties. Also, in selective applications, CO₂ rejection is decreased with the higher circulation rate and extra gas/liquid contract time. Lean Solvent Temperature. The preferred lean solvent temperature (fed to the absorber) depends on the type of gas treating application, and the particular solvent. In all applications, the lean design solvent is preferably 10-15°F hotter than the feed gas, especially when the feed gas contains greater than 2% ethane or heavier hydrocarbons. This is important to avoid condensing hydrocarbons in the amine solution, which can cause foaming. In all gas/liquid treating applications, lean temperature should be maintained below 130°F as above can result in increased circulation rate of off-spec product. In selective applications, the lean temperature should be kept as cool as possible to rfiaximize CO₂ rejection, but not below 80°F as viscosity related problems will occur. Due to vapor-liquid equilibrium considerations, H₂S removal is increased with cooler lean amine temperatures. Also, with the cooler lean temperature, CO₂ absorption kinetics is decreased. Consequently, circulation rate is minimized while CO₂ rejection is rriaximized. In CO₂ removal applications with performance design solvents, the lean temperature should be preferably maintained between 100-130°F depending on feed gas temperature and hydrocarbon content. Below 100°F, CO₂ absorption kinetics is tήndered. The hotter the lean solvent, the higher the circulation rate requirements due to vapor-liquid equilibrium considerations. When treating liquid hydrocarbon streams, it is recornmended to increase the lean solvent temperature to 130-145°F, provided the hydrocarbon would not vaporize at these conditions. This increased temperature will decrease the viscosity of the solvent and result in better amine/hydrocarbon separation in the contactor. Ajnine entrainment and solvent losses will then be kept to a minimum. Circulation Rate. With the present design solvents, a process evaluation can be performed, which specifies the design solvent circulation rate based on the design operating conditions. If actual operating conditions differ from the design conditions, the following formula can be used to approximate the required circulation rate: Required circulation rate (GPM) = (act. %H₂S) (act. %CO₂) (act. MMSCFD) (des. wt% Solvent) (Design GPM rate) x (des. %H₂S) (des. %CO₂) (des. MMSCFD) (act. wt% Solvent) Alternatively, Required circulation rate (GPM) = (32 or 41) (acid gas mole %) (vol. gas (mset)) (Design GPM rate) x (wt% solution) It is important to keep the circulation rate as close to design as possible. Usually the tendency is to over-circulate the solvent. When this occurs, needless energy is wasted in the reboiler due to increased sensible heat requirements and in electricity to drive the pumps.

Furthermore, solvent over-circulation can result in a poor solvent regeneration, inability to make specification, corrosion, reduced selectivity, and solvent losses. If the solvent is under-circulated, high rich acid gas loading and corrosion can result, along with off-spec product gas. To keep the circulation rate optimized, a target rich loading should be established and analyzed routinely. Adjustments can be made according to the results of the rich solvent analysis. Circulation rate is an important variable when operating with the present design solvent compositions. By changing circulation rate, selectivity can be altered to meet the desired process conditions with the selective, performance solvents. An often chosen alternative to varying circulation rate to achieve the desired selectivity is to equip the absorber with multiple lean solvent feed points. This will reduce or increase the amount of absorber staging and decrease or increase the gas/liquid contact time, which in turn will increase or decrease selectivity. Operating options with the primary and secondary arnines (MEA and DEA) are limited, as these arnines are nonselective, and both CO₂ and H₂S are absorbed regardless of gas/liquid contact time. Regeneration Conditions. Substantial energy savings and reboiler capacity will result when the optimal present design solvent is used versus commodity amines. For example, typical MEA regeneration requirements are 1.5 lbs. steam gallon circulation. DEA usually requires 1.0 lbs. steam/gallon circulation. Regeneration requ ements for a performance amine can be as low as 0.7 lbs. steam/gallon of circulation. The main reason for this lower energy requirement is the lower heat of reaction associated with the present design solvents. Therefore, the acid gas is readily stripped from solution with less heat input with an MDEA-based solvent. This is evidenced in what is a normal lean loading for MEA, DEA, and a design solvent in accordance with the present invention: Solvent Lean Loading, m/m MEA .05 - .1 DEA .03 Present Solvent .006

In addition to the energy savings that result from this ease of stripping, the corrosive effects of the lower residual acid gas on the hot lean areas of the plant are substantially reduced. The regenerator operations and reboiler can be easily controlled and optimized if the regenerator overhead temperature and pressure are available since this is an indication of the amount of stripping steam being generated in the reboiler. The temperature and pressure are measured at the top of the regeneration, upstream of the reflux condenser. The reflux ratio-lb. mole H₂O/lb mole acid gas is defined as "the ratio in moles of water returned to the regenerator per mole of acid gas leaving the reflux accumulator." Optimized reflux ratios of 0.85:1 to 1.25:1 are recommended for regeneration of most of the present design solvents. MEA and DEA usually require reflux ratios of 2:1 to 4:1. Once the desired conditions are met, the lean acid gas loading should be deterrnined and heat input adjusted to achieve the required lean loading. At this time, a steam circulation value can be determined for the unit operators to follow. Reboiler Conditions. Operations of the reboiler should not require any special attention except for varying steam input to achieve the correct overhead temperature as mentioned above. The reboiler temperature is very insensitive to the heat input and should not be used to control the regenerator operations. Typically, a 50-wt% solvent in accordance with the present invention at 12 psig, will boil at about 252°F. It is important that the solution be at its boiling point to ensure the solution is regenerated. To protect against thermal decomposition, it is preferable not to exceed a maximum tube skin temperature of 320°F or a bulk solution temperature of 280°F. Finally, heat flux in the reboiler is preferably below 7000 Btu/hr/sq. ft. Change-out Procedures. Once the correct design solvent has been chosen, it is necessary to prepare the unit for startup with the solvent. The correct procedure depends on the type of startup being attempted. Generally, three different types of startups are encountered, with specific change- out procedures for each. These are: • Grassroots Application • Existing Arnine Unit with Unit Shutdown • Rurining Conversion Without Unit Shutdown Grassroots Application. Preparing a grassroots unit for startup is a simple procedure. 1) Caustic wash - A 2-3 wt% caustic solution should be charged to the arnine unit, heated, circulated for three hours, and drained from the unit. This should remove any particulates, welding chemicals, and grease from the unit. It is important to check with the packing manufacturer to ensure that the cutting oils have been removed from the packing, as a serious foaming problem could result.

2) Water wash — deionized or makeup-quality water should be charged to the arnine unit, circulated for two hours, and drained from the unit. If a caustic wash has preceded the water wash, it is important that all of the caustic is removed from the unit as treating problems can result from excess caustic in solution. This can be achieved by monitoring the wash water pH. More than one water wash might be necessary. 3) Charging the solvent - If the solvent is purchased as a concentrate, estirnate the system volume, add one-half of that volume as makeup water and the other half as concentrated solvent to the unit. Circulate the solution for one to two hours, measure solvent concentration, and note the liquid levels in the unit. Add the appropriate amount of solvent concentrate and water to bring the concentration to 50 wt% and liquid levels to the proper levels. Existing Amine Unit with Unit Shutdown. Switching from a commodity solvent to the present solvent compositions is only slightly more involved than preparing a grassroots unit. 1) Circulate and strip - bypass the gas around the absorber. Circulate and thoroughly strip the acid gas from the solvent that is to be changed out. 2) Drain - drain the used solvent and dispose of it via a reputable disposal company, reclaiming company, or to another unit as makeup. 3) Inspection - inspect the absorber and stripper vessels, exchangers, reboiler, and the surge and flash tanks for signs of corrosion, scaling, fouling, and plugging. 4) Filter change - mechanical and carbon filters should be replaced. 5) Acid wash - if significant scaling and fouling are revealed during inspection, a ild 2-3 wt% acid solution should be charged to the unit, circulated for two hours, and drained from the system. An alternative is a hydroblasting service. 6) Caustic wash - recommended only after an acid was for acid neutralization. 7) Water wash — a hot water wash is always recommended before the present solvent compositions are charged to the unit. Usually, this is the only wash required when converting from any given solvent to one of the present solutions. Follow the above procedure for the water wash and charging the solvent. Conversion without Unit Shutdown. Due to the complexity of large integrated units such as refineries, often it is not feasible to completely shut the unit down to change to a solution in accordance with the present invention. However, if the following procedure is performed, a solvent conversion can be achieved, with the benefits of the present solvent compositions immediately realized. 1) Reduce arnine concentration and reduce liquid levels to as low as possible. 2) Block off the surge tank, filters, flash tank and any other vessels that can be shut down. 3) Drain solution from the above vessels. 4) Replace mechanical and carbon filters. 5) Add concentrated solvent and water to the recommended operating strength. Startup and Optimization Procedures. Once the amine system has been prepared and the; solvent charged to the unit, startup and optimization can commence. Amine unit startup with the present design solvent compositions is the same as any amine. However, during unit optimization differences do exist. 1) Begin solution circulation and bring the system up to the proper temperatures as outlined. Maximizing the reboiler heat input during startup ensures a thoroughly regenerated solution. 2) Gradually bring the gas on-stream and up to the full rates. With excess circulation and stripping rates, all acid gas specificatioais should be achieved. 3) Gradually reduce the circulation rate until H₂S or CO₂ levels in the treated gas begin to increase and approach specification. At this point increase the circulation rate by approximately 2-10 % as a safety factor to the operation. 4) After the circulation rate has been optimized, gradually reduce the reboiler duty until the optimized overhead temperature is obtained. It may be necessary to increase the circulation rate slightly once the reboiler duty has been reduced to keep the treated gas within specification. Ultimately, true goal is to achieve a balance between the circulation rate and reboiler duty. After making changes in circulation or reboiler duty, the system should be allowed significant time to line out. It is important to monitor the operating data and mass balance around the system during startup and optimization as the data is useful for troubleshooting and optimization should process conditions change. Solution Monitoring And Control Solution monitoring and proper maintenance are important when using the present design solvent compositions to keep operations trouble-free, miiώnize costs, and to consistently achieve the benefits the solvent can provide. The results of a routine solvent analysis of the present design solvent compositions are to keep the solvent in the proper condition; the following quality factors should be monitored and controlled: • Solvent Concentration • Acid Gas Loadings • Contaminants • Foaming Tendency Solvent Concentration: This analysis should be performed daily. As mentioned above, solvent circulation rate is directly affected by amine concentration for a given set of conditions. This analysis is quick and requires only simple equipment. Acid Gas Loading: The lean and rich acid gas loadings, usually expressed as moles of acid gas per mole of solvent (m/m), are important because, of the insight they provide into absorber and regenerator performance and overall solvent utilization. A typical lean solvent loading with the present solvent compositions is .005 - .02 m/m. A significantly higher lean loading is indicative of improper regenerator operation or a mechanical problem in the regenerator. Continued operation under these conditions could cause acid gas corrosion in the regenerator or reboiler in addition of off-spec operations. Lean loadings below these levels are an indication of over-stripping and excess energy consumption. Rich solvent loadings are dependent on many factors and can provide insight into unit operations. Computer modeling of the systems operating parameters with the performance design solvent will determine an optimized rich loading. Analysis of the rich solution will then provide insight into unit optimization by comparing the actual rich loading to the optimized loading. If the rich loading is to low, either circulation rate can be reduced followed by a reduction in reboiler duty or a problem exists in the absorber that is hindering acid gas absorption. If the loading is too high, the circulation rate should be increased to avoid corrosion caused by flashing of acid gas from the rich solution in the heat exchanger. Contaminants: Solvent contarriinants can enter the system via the feed, makeup water, or be formed by thermal or chemical degradation of the arnine. Contaminants common to the amine system are: Heat-stable Amine Salts Makeup Water Impurities Thermal and Chemical Degradation Compounds Out of all the amines used in gas treating, including the performance amines, only MEA, DEA, and Sulfinol can be thermally reclaimed at atmospheric pressure because of their low boiling points. However, use of a reclaimer can result in high solvent losses and high-energy requirements. With the present compositions, a technology does exist to remove certain types of solvent contaminants via on-line separation. Also, the preferred tertiary amine-based present compositions are much more resistant to chemical degradation. All types of solution contaminants can result in operational problems such as corrosion or high costs. The contaminants should be monitored and their accumulation rrjrinimized. Heat-Stable Amine Salts: Heat-stable amine salts are created by the reaction between the arnine and highly acidic contaminants that enter with the feed. These salts are called heat-stable because they cannot be thermally regenerated under normal stripper temperatures. They will form regardless of amine type and can cause problems such as corrosion and foaming. Also, they reduce the acid gas capacity of the solvent by rendering a portion of the amine inactive. The types of acid anions that form heat-stable amine salts can be analyzed by ion chromatography. Total heat- stable amine salt content can be determined by a simple titration method. Depending on the type and quantity of acids present and the particular system, heat- stable arnine salt content should not exceed 10% of the active amine concentration. The present compositions have been developed for on-line removal of these salts. As a result, the high solvent usage and corrosion associated with these salts is substantially reduced. However, for feed containing a large amount of acid contaminants, water washing is recommended for removal of these contaminants before they contaminate the amine solvent. Makeup Water: Makeup water can contain contaminants that will accumulate in the system and cause operating problems such as foaming, corrosion, plugging and fouling of equipment. A metals analysis can be performed on the a ine solution and makeup water. Typical contarninants found in an arnine solution because of poor water quality along with other metals common to amine systems are shown in Table 1. Makeup water obtained form a steam source is most desirable. Boiler feed water is generally not acceptable due to additives that cause foaming. All amines are subject to some degree of thermal degradation. However, if the previously mentioned reboiler temperatures are not exceeded, nririimal degradation will occur with the present solvent compositions. In addition to being resistant to thermal degradation, columns operated using the present compositions are much less susceptible to chemical degradation than MEA or DEA. Consequently, problems caused by degradation compounds such as corrosion, high solvent makeup requirements, and foaming are significantly reduced with the present compositions. Foaming Tendency: The performance amine solutions do not have an inherent foaming tendency. Foaming is usually caused by surface-active contaminants that enter the system with the feed. These contarninants reduce the surface tension of the amine solution, which usually results in an increased foaming tendency. Foaming can result in high solvent losses, failure to meet acid gas specifications, and corrosion. A simple test can be performed in the laboratory for monitoring solution foaming tendency. This test involves measuring foam height and stability. Also, the effectiveness of certain antifoamers can be evaluated with this test. These results can be correlated with plant conditions, followed by the appropriate corrective action. Troubleshooting and Corrective Actions: If an amine unit is given adequate attention by following the above procedures and suggestions, good performance should result. However, occasionally problems can arise. Frequently, the causes of these problems are so obvious they are overlooked. If the symptoms can be correctly diagnosed and the proper corrective action performed, the problem can be controlled and eliminated before it causes serious additional problems. The three common arnine unit problems are: • Foarriing-Arnine Losses • Off-spec Product • Corrosion Foaming: Foaming can occur in both the absorber and regenerator. It results in solution carryover out of the top of the tower, high arnine losses, contamination of downstream processes, off-spec product, and corrosion. Foaming is usually caused by: • Hydrocarbons • Particulates • Mechanical Troubles • Degradation Compound

Foaming is indicated by increased pressure drop across the absorber or regenerator and by an unexplained drop in liquid level in the surge tank. A recording differential pressure gauge on the absorber and regenerator along with performing a foam test are methods used to monitor foam troubles. To cure foaming, the source has to be found and eliminated. The best cure is proper care of the design solution. Good antifoamers can be added to the solution. However, this should only be used as a temporary measure. Too much antifoamers can actually promote foaming. Hydrocarbons will condense in the amine solution if the lean solvent temperature is not maintained 10°F above the feed temperature. Also, liquid hydrocarbons will enter the system and cause foaming not only in the absorber, but are the main cause of foaming in the regenerator. Consequently, the rich solution flash tank should be checked for proper operation. If the solution is dark, then improved particulate filtration is necessary. Foaming can be caused by mechanical limitations if throughput is exceeding design. Finally, degradation compounds can cause foaming. However, this is usually the least likely source when using the present solvent compositions because of their chemical and thermal stability. Overall, the present compositions exhibit a lower foaming tendency than commodity amines. Off-Spec Product: Failure to meet specification is common problem in gas treating. This problem is easily remedied if the feed conditions are not exceeding design. If a commodity solvent is being used and feed conditions are exceeding design, a solvent switch to a performance design solvent will usually provide the necessary capacity. Design considerations aside, usually this problem is caused by: • Low Solution Flowrate • Low Amine Concentration • Inadequate Solvent Regeneration • Lean Amine Temperature Too High • Foaming • Mechanical Damage

Amine concentration and solution flowrate should be monitored and increased to add the necessary capacity for proper acid gas removal. The lean solvent loading should be checked and adjustments made to the regenerator are the lean loading is too high. Also, the lean/rich exchanger should be checked for a leak, as rich solution will flow into the lean solution and increase the lean loading. This can be accomplished by analyzing the lean solution in and out of the exchanger. If the lean temperature is too high, the equilibrium capacity of the solvent is reduced and the temperature should be lowered. Finally, the absorber should be checked for corrosion, scale or otherwise damaged internals. Corrosion: Typically, corrosion with performance design solvents is miriimal and expensive metallurgy is not required. However, if improperly operated, or if solution contamination exists, corrosion can occur. Typical causes of corrosion are: • High Amine Loadings • High Heat-stable Amine Salts • Particulates • Foaming Inadequate solvent regeneration in the stripper tower will result in high residual acid gas loadings. This can cause corrosion in the stripper bottom, reboiler and other hot lean areas such as the lean/rich exchanger. Also, the solution pumps can corrode due to acid gas flas-hing form solution. The lean loading should be monitored and adjustments made to the regenerator if the loading is too high. High rich loadings can cause corrosion due to acid gas flashing form solution in the absorber bottom, heat exchanger and rich solution piping to the regenerator. If the acid gas loading analysis indicates an excessive rich loading, solvent circulation should be increased. To mimmize acid gas flashing, the pressure-reducing valve on the rich stream to the stripper should be located as close to the stripper as possible. Corrosion from heat-stable amine salts is usually limited to the reboiler and especially the reboiler tubes. This is due to the temporary disassociation of the acid anion form the arnine. The present compositions and systems can be used to reduce the heat-stable salt content. The present silicate solutions are very effective in reducing corrosion in these contexts. Erosion/corrosion is caused by a high amount of particulates in solution. Erosion usually occurs in areas with high solution velocities such as pump impellers and piping bends. Again particulate filtration is important to avoid this problem. Foaming in the regenerator will cause corrosion in the regenerator and reboiler. This occurs because the rich solution proceeds down the tower without having good contact with the stripping ^vapor. Consequently, proper regeneration does not occur and the aforementioned acid gas corrosion will occur. Thus, to reduce arnine solvent losses in alkanolamine gas treating plants, a systematic approach is vital. First, through an accurate plant inventory deteπnination of existing plant loss rate can be made on hourly or daily bases, and vs. GAS treating rates. Second, after collecting plant design drawings, plant operation data, and complete solvent analysis, characterize the type and amount of losses result from vaporization, solubility, entraήiment, and degradation. Then calculate the mechanical losses as the difference between losses in total inventory and estimates of characterized loss areas. Inspect the plant equipment to identify any individual large mechanical losses and estimate the amount of each. Finally, rank each category of losses from highest to lowest and focus operation in equipment changes in that order. Advantages of the present systems and compositions.

The present silicate-containing alkanolamine design solutions will benefit the process outlet gas separator unit. The solution family rejects hydrocarbon absorption, which hydro carbon absorbtion promotes solution carryover into the outlet process pipeline. The present silicate-containing alkanolamine design solutions will benefit the contactor unit. Partial Pressures. The operating pressure in the contactor/absorber is the physical mechanism required to initiate a kinetic chemistry necessary to remove the acid gas components from a hydrocarbon stream. Most operating pressures found within the oil refinery are less than 500 psig, which does not allow the chemistry between the conventional hydrocarbon treating solvent and the acid gas components to have an immediate reaction. The profile at the contactor/absorber becomes very linear if plotted. The highly reactive chemical makeup of the present compositions, however, allow for a definitive profile, which can be plotted and evaluated around the contactor/absorber. The bulk of the reaction occurs exothermically and immediately on the lower trays of the process tower. Acid Gas Loadings. Perhaps the singular most important characteristic of a treating solvent is its reactivity. Current technologies allow for a maximum acid gas loading of .45-.50 mole per mole of acid gas comprised of hydrogen sulfide, CO₂, COS and mercaptans. The present technology has successfully achieved a rich loading of its solvent technology to a .71 mole per mole. This capability equates to more available pound moles able to remove more acid gas from hydrocarbon process. Higher mass transfer capacity increases deeper, more affective removal of acid gases associated with the inlet stream. Solubility Concerns. The higher weight percent at which the present system is allowed to operate does not incur any concerns regarding the solubility of any hydrocarbon stream into solution. The base chemical the alkaline silicate, in the present compositions does not have a chemical affinity for the hydrocarbon stream. Traditionally, an undesirable rag layer is observed forming in the site glasses of the process vessels throughout the system. This rag layer is undesirable because often the rag layer contains the high end octane points that will have to be replaced in the polished fuel. Conventional alkanolamines have historically had a problem regarding the natural affinity for the hydrocarbon stream. This problem has forced end users of such products to do the following: Reduce the weight percent of the hydrocarbon treating solvent. Increase the weight percent of the hydrocarbon treating solvent. End users generally realized an increase in fuel usage for regeneration purposes. The Present silicate-containing alkanolamine design solutions will benefit a flash tank unit. The family of performance solutions increases the capability of the various solvents to flash regenerate acid and hydrocarbon gases out of solution as a result of the initial rich control valve located between the process contactor and a process flash tank. The Present silicate-containing alkanolamine design solutions will benefit a cooler unit The process solvents provide a passivating protection to the internal surface for the tube bundle. In addition, the heat duty is reduced for design applications. The Present silicate-containing alkanolamine design solutions can be safely operated at higher weight percent than amine solutions found in today's market Solvent Percentage. The present system and compositions can easily increase the weight percent of solution of the desired components in any of its products to 60 weight percent without the experiencing any problems regarding viscosity or the ability to introduce it as a virgin product into the process system. The introduction of higher weight percents allows higher mole concentrations capable of removing kinetically higher concentrations of acid gas on a stable, continuous basis from the inlet hydrocarbon stream. The Present silicate-containing alkanolamine design solutions will benefit a gas stripper unit Reducing the stream input required to regenerate the rich solution cascading through the stripper tower. The Present silicate-containing alkanolamine design solutions will benefit a reboiler unit. Reboiler Duty. The high loadings of 0.71 mole per mole require less heat duty to achieve the lean loading of .003 mole per mole required to return back to the contactor/absorber to repeat the continuous kinetic chemical reaction required necessary to remove the acid gas stream from a hydrocarbon components' main stream to achieve the necessary outlet specification. The calculated heat of reaction required to regenerate the Selarnine Products is calculated at 410 BTU/lbs. for hydrogen sulfide and 419 BTU/CO₂. The Present silicate-containing alkanolamine design solutions will benefit a reflux condenser unit by reducing contaminates normally found utilizing typical amine solutions. The Present silicate-containing alkanolamine design solutions will benefit a reflux accumulator unit. Degradation Problems. The present solvent solutions are highly resistant to degradation induced by contaminants introduced by the sour process hydrocarbon stream or by thermal degradation. This advantage over conventional amines nrinimizes the possibility of corrosion taking place internally within the process system.

The Present silicate- containing alkanolamine design solutions reduce all known types of internal corrosion within the process. Corrosion Concerns. The present solvent solutions include compounds that are able to enter the matrix of the metallurgy of which the process plant is constructed. The chemistry involved to achieve this is primarily a function of temperature. The process plant utilizes for the removal of acid gas, components from a hydrocarbon stream that are generally operated on average, at a temperature of approximately 150°F. This passivity feature provides the protection required for CO₂ applications where H₂S is not present to put down a passivating film. The passivating effect is critical to mirώnize the possibility of the various types of corrosion inherently found in alkanolaniine plants (Le. crevice attack, erosion problems, galvanic corrosion, and hydrogen damage etc.)

The Present silicate-containing alkanolamine design solutions can be used for preventing corrosion in vessels and in pipelines Corrosion Proofing. When pressure testing vessels and piping, substituting a heated solution according to the present invention in lieu of plain water as the test fluid will provide corrosion resistance. The heat, pressure and contact involved in the pressure testing and corrosion treating process will drive the compounds from the solution into the matrix of the metal and provide corrosion resistance and can in many cases alter the metallurgical requirements. The Present s cate-containing alkanolamine solutions, when used in conjunction with specialized equipment for "Mass Transfer" and with optimum filtration/separation/clarification, can also facilitate the removal of the varying sulfur species and other elements from water, oil, crude, diesel and NGLs. Finally, the Present sificate-containing alkanolamine solutions can be used for removing the "heavy" metals (such as nickel and vanadium) from the various fuel and crude oils used for instance in power plants. This process could solve the problems in power plants associated with these metals. This process would elirriinate the need to add magnesium as used for corrosion resistance in power plants. DIRECT TREAT APPLICATIONS The present systems and compositions can be used in a direct treat application with both a contaminate burner and flash gas recovery. Referring now to Figure 2, one

Pilot Results Applicable to Direct Treatment Cases Since the inlet gas stream at the pilot unit did not contain SO₂, the SO₂ had to be introduced to the process, as would be the case in commercial direct treatment applications. Because of the relatively small size of the pilot unit, it was most economical to purchase liquid SO₂ and inject it into the pilot unit. For much of the testing, the SO₂ liquid was injected directly into the absorber. Injecting liquid SO₂ into the lean solution stream being recycled to the absorber was also successfully demonstrated. If liquid SO₂ or a contaminate burner were used in a commercial application, SO would be carried into the absorber via the lean solution. The SO₂ is highly soluble in the solution and, in the case of a contaminate burner, would be added through a small scrubber column on the contaminate burner exhaust. Gas-phase injection upstream of the absorber was also successfully tested. A potential concern for direct treatment cases is the possible loss of the SO₂ source via

37 embodiment of a direct treat apphcation includes an inlet feed line 10, an absorber 20, a crystallizer/settler 30, a contaminate filter system 40, and a solution recycle line 50. In addition, a preferred system includes a flash vessel 35, which receives contaminant-rich solution from the bottom of absorber 20 and reduces the pressure thereon so as to remove the low-boiling components of the rich stream prior to the stream entering crystallizer/settler 30, and a compressor 33 for compressing the flashed components prior to re-injection into absorber 20. In the embodiment shown in Figure 2, the system further comprises a contarninate burner system 45, to which a portion of the contaminate recovered in filter 40 may optionally be fed for oxidation, along with an oxygen containing stream such as air. The combusted contaminate (e.g. SO₂) from burner 45 is fed into a second absorber column 70. A portion of the lean solution from stream 42 is passed through column 70, preferably in a countercurrent flow configuration, and strips the SO₂ from the exhaust gas, which is then vented to a flare. As discussed above, the addition of SO₂ can be performed in three different ways, and the decision will be very site-specific, depending on location and quantity of contaminate produced. For example, the injection of liquid SO₂ is economic only for small contaminate tonnages and transporting it to the site could present potential concerns for plants in urban locations. Contarninate burning has been used in many industries and is a proven approach, but will have a higher capital cost than liquid SO₂ injection for small applications. It is also possible to thermally or catalytically oxidize one-third of the inlet H₂S to SO₂. If the stream to be treated is an acid gas stream or other stream of no value, then combusting one-third of the stream can be an economic way to produce SO₂. If the stream to be treated is valuable, then it may be possible to catalytically convert 1/3 of the inlet H₂S to SO₂ without destroying the valuable constituents of the stream. SUMMARY OF DIRECT TREAT NEEDS AND APPLICATIONS Gas containing very little contaminate (e.g., less than 0.2 LT/D) can be cost-effectively treated with nonregenerable scavenging chemicals. This can be performed by injecting a liquid scavenger directly into a pipe containing the sour gas (direct injection) or by passing the sour gas through a tower containing a liquid or solid scavenger. Gas containing more than 25 to 30 LT/D of contaminate is generally processed by first separating the acid gases with an arnine unit and then processing the arnine off gas in a sulfur plant to produce molten elemental contaminate. However, gases containing medium amounts of contaminate (e.g., between 0.2 and 25 LT/D) have generally posed treatment challenges to industry. Often, gas in this niche has been processed in plants using liquid redox technology. While liquid redox technologies 28 often appear to be the economic choice, these plants frequently have operability and reliability problems. Furthermore, no liquid redox technology has been demonstrated to reliably recover contarninate in a single step if the pressure is greater than around 250 psig. The present systems and design compositions may be used in a number of direct treatment applications in the 0.2 - 25 L T/D range, including the following: - Natural gas (especially at high pressures); - Refinery fuel gas; - Refinery hydrodesulphurization (HDS) recycle streams and vent streams (especially as refineries address potential needs for additional Contaminate removal capacity); - CO₂ for enhanced oil recovery; and Geothermal vent gas.

SULFUR TAIL GAS APPLICATIONS Referring now to Figure 3, a preferred system for using the present compositions in sulfur tail gas applications includes an inlet stream 110, an absorber 120, a solvent recycle line 144 and a contaminate removal loop 142. Contaminate removal loop 142 includes a crystallizer 130 and a filter/wash system 140. Solution from the bottom of absorber enters contaminate removal loop 142 via line 143 and enters crystallizer 130. Solution leaves crystallizer 130 via line 145 and is cooled and recycled back to crystallizer 130. A portion of stream 145 is preferably diverted to filter/wash system 140 via line 146. Solution leaves filter/wash system 140 via line 147 and flows back to crystallizer 130. Solution that is lean in contarninants is removed from crystallizer 130 and recycled to absorber 120 via line 150. If desired heat is added via heat exchanger 138. In addition, an overhead system 170 receives the sweet gas stream from the top of absorber 120. Overhead system 170 includes a condenser/separator 174 and an incinerator 180. Condenser/separator 174 separates the sweet gas stream into a water stream 175, a solvent stream 176, and an exhaust stream 177. Solvent stream 176 may be recycled back to absorber 120 via line 144. Exhaust stream 177 is sent to incinerator 180, where it is burned. Since the present solutions have high boiling points (typically above 400°F) and can dissolve/absorb elemental contaminate, they can be used in direct contact to cool sulfur tail gas. In this mode, a portion of the contaminate-rich solution from the bottom of absorber 20 is run through a cooler 23 and across a bed of packing 25 in the lower portion of 20 absorber. This cools the tail gas from its entering temperature of 270-300°F down to 140-180°F (5-10°F above the water dewpoint). This eliniinates the need for a separate water quench system that 29 other processes frequently require. When applied to sulfur tail gas, the contaminate recovered with the present solutions is expected to be sufficiently pure so that it can be mixed with the sulfur plant contarninate and sold. Contaminate from a pilot unit has been blended with pure sulfur contarninate and analyzed. The analyses proved that the blends were as pure as the original sulfur Contaminate, within the capabilities of the analyses. The present compositions do not remove or convert COS or CS₂ in the solvent tail gas. If these compounds are present in significant quantities, titanium dioxide catalyst can be used in the first sulfur catalyst bed in conjunction with warm bed temperatures to eliminate the COS and CS₂. SUMMARY OF SULFUR TAIL GAS NEEDS AND APPLICATIONS While the requirements may vary some across states, new contaminate plants over 20 LT/day in the United States are generally required to achieve efficiencies of 99.8+%. Smaller facilities and existing grandfathered plants may be allowed lower levels of control. ADVANTAGES FOR SULFUR TAIL GAS TREATING The most widely used approach for sulfur tail gas treating has been the use of H₂S recycle processes. In these processes, a reducing gas generator (burner), and hydrogenation reactor are used to convert all contarninate species to H₂S. The hot gas is then passed through a waste heat boiler and a quench tower to reduce the temperature of the stream. The H S is absorbed and reconcentrated with an amine system. Finally, the H₂S is recycled to the front end of the Suifur process. These types of processes have historically had the highest overall contaminate removal efficiency, but are expensive, with the tail gas treating process often exceeding the cost of the Sulfur plant. Many of the other processes have had the drawback of not being able to achieve the high efficiency of the H₂S recycle processes, particularly when the normal swings in tail gas H₂S:SO₂ ratio adversely affects their removal capabilities. The present compositions address the shortcomings of the previous processes by having significantly lower capital and operating costs than the H₂S recycle processes, while still achieving the same high efficiencies. A summary of advantages is as follows: Achieves 99.8+% efficiency: Contarninate vapor is very soluble in the present solutions and is scrubbed efficiently (to 10 ppm or less), Operates at lower temperatures (150-180°F) versus other processes (> 250°F) that are limited by equilibrium at higher temperatures, and 30 present solution's buffering capacity maintains removal during H₂S:SO₂ ratio swings. Low cost, simple plant: No reducing gas generator, No hydrogenation reactor, No waste heat boiler, and No quench water system No recycle of H₂S, CO₂, or SO₂: Does not reduce Sulfur reaction furnace temperature or capacity, and Simplifies control of the Sulfur unit. WATER BALANCE CONSIDERATIONS Sour product is typically saturated with water at the inlet to the contaminate recovery process. An additional amount of water is made as a result of the reaction to form elemental contaminate product. To remain in balance, the water entering with the sour product and the water made by reaction must exit the system. The considerations for direct treatment of sour gas streams are different than for tail gas applications and considerations of the overall outcome of the end stream outlet. WATER BALANCE WITH DIRECT TREATMENT OF SOUR GAS The present process can typically operate at temperatures from about 130°F to approximately 180°F or higher. Sour product temperatures vary, but are frequently less than 100°F for direct treat applications. Since the gas gains in temperature across the present system, water formed by reaction typically exits with the sweet gas. This is currently a preferred method of niamtaining the water in balance with the present process. Although non-aqueous, the steady state concentration of water dissolved in the present solution may reach as much as 1-2 wt%. Water can be stripped from the solution. This means that plants that use sour contaminate burners to produce the SO₂ have another important water outlet. Under typical contaminate burner/SO absorber operating conditions, much if not all of the water formed by reaction can be stripped out of the solution via the SO₂ absorber. WATER BALANCE WITH TAIL GAS TREATMENT In tail gas treatment, the tail gas typically enters at temperatures of 270°F or higher, and the water dewpoint of this gas ranges from 130 to 175°F. In these cases, the tail-gas treating unit is simply operated at temperatures that are 5-10°F above the water dewpoint. OTHER APPLICATIONS The present compositions can be used to advantage in a wide variety of industrial 31 applications involving contaminate pollution. In many cases, the present compositions will reduce extremely high levels of sulfides, thiosulfate and sulfites (e.g. above 1000 mg/L to less than 10 mg/L). The following examples are illustrative of the diversity of applications. Treatment of biological upsets in pond and lagoons: in both the food processing and pulp and paper industries, biological treatment systems can suffer anaerobic conditions that lead to the formation of H₂S and other sulfides. Batch treatment with the present compositions not only rids ponds and lagoons of odor and poisonous sulfides; it also leaves behind a residue that contains excess oxygen and helps prevent future anaerobic conditions. Treatment of sour waters from a refinery: one refinery compared the use of the present compositions to stripping to remove sulfides from a sour water stream. Using an H₂S:SO₂ ratio of 3:1 at a pH of 8.5, the refinery successfully reduced sulfides to required limits and avoided the capital cost of the stripper. Pretreatment of process wastewater: in a refinery environment, sulfide remaining in the waste stream after sour water stripping can be oxidized to thiosulfate by air and steam. The present compositions can then be utilized to convert the thiosulfate to sulfate. This will allow municipal collection systems to be accepted as treated wastewater. Hydrogen sulfide scrubbers: all municipalities experience problems with high sulfide levels after primary treatment of sewage. A caustic scrubber system should be installed to absorb the hydrogen sulfide. Introduction of the present compositions into the scrubbing solution will oxidize the sulfide to sulfate. Geothermal applications: Condensed steam from geothermal power plants contains high levels of sulfides. The present compositions can be introduced along with an iron catalyst into the scheme to address this problem. Industrial water reuse: commercial car wash operations under a zero discharge permit often experience hydrogen sulfide odors in their water even after the water has gone through settling, ion exchange and carbon adsorption treatment units. By treating with the present compositions after the carbon absorption unit, the odors are expect to be completely eliminated. Municipal odor control: wastewater treatment plants often have severe hydrogen sulfide odor problems associated with the wastewater. Dosing with the present solutions would allow the facility to control its dissolved and airborne sulfide levels. Sludge treatment: wastewater treatment plants often experience high levels of hydrogen sulfide in the sewage sludge. The hydrogen sulfide volatilizes at the thickener and 32 causes severe odor and health problems. By adding the present solutions prior to sludge thickening, such plants can successfully combat the hydrogen sulfide and related problems. Treatments of sludge in pond closings: many industries are closing and/or refurbishing treatment ponds to meet stringent RCRA guidelines for permeability. Some of these ponds contain sludge that is high in sulfides. Treating such ponds with the present solutions is expected to reduce sulfide levels from several thousand mg/L to less than 20 mg/L. Further Operating Considerations Use of the present design solvents can save energy, provide capital savings on new equipment or provided a capacity increase with existing equipment. Also, operational problems such as corrosion can be substantially reduced. However, using the present solvent compositions requires that the solvent be properly used and maintained, in order to avoid problems that can negate the benefits of a performance solvent. Provided a few basics are adhered to, the present solvent compositions will perform well with little difficulty and attendance, while reducing costs and providing the aforementioned benefits. Good performance starts with proper selection of equipment, followed by good operating practices and solution monitoring and control. Finally, understanding the occasional problem that could occur and the proper corrective action will lessen the severity of such problems and quickly allow normal operations to resume. The above considerations are often related. The following discussion explains how each applies to gas treating with the present compositions. Selection of Preferred Equipment. Selection of the proper equipment is the first step towards achieving a successful gas treating operation. Many of the problems common to arnine unit operation can be minimized and avoided via installation of equipment that will lessen the amount of contaminants entering the system, remove contaminants from the system, reduce arnine losses, and reduce the potential for corrosion. Specifically, five pieces of equipment that are preferably incorporated into an arnine unit constructed in accordance with the present invention (Figure 1) are: • Inlet Separator • Treated Gas Knockout/Coalescer • Flash Tank Separator • Particulate Filter • Carbon Purifier

Inlet Separator. 33 One of the wisest investments that can be made is in adequate removal of contarninants that can enter the system with the sour gas. A variety of contaminants such as solids, down- hole or pipeline treating chemicals, liquid slugs cased by volume surges or line pigging, compressor lubricants, and in refining applications, large amounts of sponge oil and acid contaminants can be eliminated via a proper inlet separator. These contarninants promote foaming and can hinder and shutdown operations if allowed to enter the system. The design of the inlet separator depends on the type of gas being treated and the level of expected contaminants. However, most are gas-liquid separators equipped with an impingement baffle and coalescing device. In refining applications, consideration should be given to a combination separator/water wash for removal of both hydrocarbons and acid impurities that are generated in upstream processing. Outlet Gas Knockout and Coalescer. The outlet gas knockout is located downstream of the absorber. It serves to miriimize controllable solvent losses due to entrainment. However, its main function is to protect downstream processes and minirnize solvent loss from uncontrollable carryover that is usually caused by mechanical malfunctions or foaming. Design considerations should include proper sizing for liquid slug handling capacity, along with a coalescing element to remove entrained mist. The liquid dump valve should be properly sized for withdrawal of large amounts of solvent that result from foaming or upsets. When treating a liquid hydrocarbon stream, it is advisable to install a coalescer downstream of the liquid/liquid contactor. This vessel allows additional time for separation of arnine and hydrocarbon and serves to recover aπiine during upset conditions. A coalescer is a horizontal vessel usually with a 3:1 L/D ratio. Residence time is 10-30 minutes. A coalescing element is located between the hydrocarbon inlet and outlet. Flash Tank Separator. A flash tank separator is located on the rich arnine stream, upstream of the lean/rich exchanger. It serves three purposes: degassing of volatile, dissolved hydrocarbons; separation of heavier liquid hydrocarbons; and vaporization of a portion of the acid gas in solution. This vessel is important because it prevents hydrocarbons from entering the regenerator where they can cause foaming or cause problems in the processing of the acid gas stream. Also, the volatile hydrocarbons can be collected and used as fuel. The design should consider residence time, velocity and quiescence. Liquid residence time varies between 10-30 minutes depending on the amount of hydrocarbon disengagement 34 required. Design should include a settling arrangement and drain for the removal of liquid hydrocarbons. For direct treatment of sour gas, the inert vapor components in the inlet gas will reach steady-state concentrations in the solvent solution. The steady-state concentration in solution depends on the vapor concentration of each component in the inlet gas. As an example, the inlet gas at the pilot unit host site (being an EOR application) has quite a few heavy hydrocarbons; as a result, the steady-state concentration in the pilot unit solution is believed to be about 11 wt% in the lean solution stream. These additional hydrocarbons act only as a diluent to the main solution components and do not affected operating performance at the pilot unit. In feet, it is expected that slugs of hydrocarbon contaminants could be purposely added to the system at up to approximately 14 volume percent with no effect on operations. In contrast, hydrocarbon contarninants often cause extreme operating problems in aqueous processes, even though there are much lower concentrations of hydrocarbons at steady state. For the reasons cited above, occasional slugs of liquid hydrocarbon do not harm operation of the present process. However, slugs of liquid containing very heavy components will stay in the system for a long time. Continuous slugs of heavy hydrocarbons would result in undesirable dilution of the present solutions. As with any process, appropriate knockouts/separators should be used to prevent liquids from entering the system, Particulate Filter. Particulate filtration is essential for mamtaining good solution quality. Solids in the amine solution can cause foaming, fouling and plugging, resulting in high arnine losses, erosion/corrosion and damage to the absorber and regenerator internals. Solids filters are usually of three types: cartridge, mechanical or pre-coat. Cartridge filters are the most popular because of ease of change-out and maintenance. Design and location of the particulate filter depends on filtration requirements. The filter can be full flow or slipstream and can be located on both the lean and rich sides. With the present solvents, filtration requirements are usually less because of low corrosion resulting in less iron sulfide particulate. Solution filtration requirements depend on the micron size and quantity of the particulate. Generally, 10-micron size filters are used with a 20-30 micron size suitable for particulate removal. Particulate in the circulating solution should be kept below 0.02% by weight. The filter should be equipped with a differential pressure gauge to monitor removal effectiveness, and be able to operate up to pressure differentials of 25-35 psig. Above this, the 35 elements could collapse and become ineffective. A good filter supplier can help recommend the proper filter. Carbon Purification. An active carbon bed will remove surface-active contaminants such as hydrocarbons, compressor oils, pipeline corrosion inhibitors, and well-treating chemicals that promote foaming. Therefore, a carbon bed is recommended in performance amine service due to its ability to indirectly reduce arnine consumption, and the costs associated with plant shutdowns along with good process control. The carbon bed should be designed to treat a 5-10% slipstream of the cool lean amine stream. On small systems, the carbon bed can be full flow. A 20-rninute contact time between amine solution and activated carbon is recommended. Superficial velocity should be two gallons per minute per square foot of cross-sectional area. A L/D ratio of 4/1 is recommended for the carbon tower. Pilot Unit Testing And Results Summary A pilot unit was constructed that was capable of treating a stream from enhanced oil recovery that is approximately 80% CO2 and 10% methane, with the remainder being heavier hydrocarbons. The gas contains approximately 2,000 ppmH₂S and is at roughly 300 psig. After conducting parametric testing, the plant was at steady-state (except for some limited additional parametric testing), and continued to meet specifications. Some design improvements were subsequently implemented. The pilot unit was evaluated on eight criteria; including foaming, plugging, chemical cost, removal, and byproduct make rates/removal, impact of contaniinants, Contaminate quality, and materials of construction. A surnrnary of the evaluation results are presented in Table 1. Overall the pilot unit demonstrated good performance in all areas, validating that the present inventions represents a significant breakthrough for these applications Table 1. Summary of Pilot Unit Evaluation Areas Evaluation Results Area Foaming Not experienced in pilot operations

36 a rrialfunction of the SO₂ injection system or the contarninate burner. The effects of losing the SO₂ source were evaluated in the laboratory bench-scale unit by shutting off the SO₂ feed during some portions of the testing. It typically took 6 to 12 hours before the removal started to decline. Testing at the pilot unit although less extensive, also confirmed the buffering ability of the alkanolarnine solution. Because of the specific chemistry of the solution, a "buffer" capacity is achieved that allows the system to continue achieving removal even without the addition of SO₂ for significant periods of time. Another issue for the direct treatment case is the impact of gas pressure. For high- pressure gas streams, losses of the inlet gas through flashing can be a concern. During the pilot testing at 300 psig, only 1.2% of inlet hydrocarbons were lost in the flash tank. Operation at higher pressures would result in more absorption and greater potential losses. In these cases, a flash tank vessel would likely be used, with the flash gas recycled to the absorber (the use of a two-stage flash would be determined by the overall economics of the gas losses versus the cost of additional flash stages and recompression). For the pilot unit, approximately 1.9% of the inlet CO₂ was lost in the flash gas, but only 0.02% or less of the H₂S was flashed from the solvent solution. (H₂S levels in the flash gas were typically below 20 ppmv.) This suggests that the H₂S-to-contarrιinate conversion reactions were largely complete by the time the solvent solution reached the flash tank. In many direct treat applications, this will be an acceptable level of H₂S flashing; however, if necessary, the H₂S in the flash gas could further be reduced by adjusting the reaction time allowed in the bottom of the absorber or in a separate reaction vessel. The effect of CO on gas treating processes is also often important. For aqueous liquid redox processes, applications with high partial pressures of CO₂ have historically been difficult since CO₂ reduces the aqueous solution's pH, thus inhibiting H₂S removal. Because the present solutions can be designed to be nonreactive, CO₂ can be designed to be absorbed only to a very limited extent. As demonstrated on the pilot unit (85% CO₂ stream at 300 psig, which is roughly equivalent to the CO₂ partial pressure of a 24% CO₂ stream at 1000 psig), the present solutions are not affected by CO₂. Pilot Results Applicable to Sulfur Tail Gas Applications Many of the same results from the pilot unit that demonstrated robustness of the present solutions for direct treat applications also apply to sulfur tail gas applications. For example, the proven "buffering" capacity of the solution with regard to SO₂ means that the process should be able to continue to run for an extended period even if a sulfur unit upset (or normal short term fluctuation) causes the H₂S:SO₂ to deviate from the desired 2:1 ratio. Similarly, the demonstration of the process on a stream with high partial pressures of CO₂ provides the certainty of knowing that the CO₂ in the tail gas stream will not affect the process operation. There are other issues unique to sulfur tail gas that are not addressed by the direct treat case, but that have been addressed during the pilot testing. During one portion of the pilot testing, the SO₂ was vaporized and injected into the gas upstream of the pilot unit. The pilot unit continued to operate well and achieve the removal specification, demonstrating that there is no removal issues associated with absorbing both H₂S and SO₂ simultaneously from the gas phase into the solvent solution. Calculations relative to applying the process to sulfur tail gas indicate significant economic and performance advantages and benefits versus other approaches to 99.8+% contaminate control in sulfur plants. Figure 5 A illustrates an external perspective of a novel absorber or mass transfer unit. The absorber of Figure 5 provides a 6 inch flow diameter and is 10.5 feet long. Such equipment should be able to provide 450 gpm nominal liquid flow capacity. Figure 5 illustrates the relative size between a comparable novel mass transfer unit of the instant invention and a conventional packed tower absorber. Figure 5C is a cut away view of an embodiment of the present invention. In a test case using the embodiment of Figure 5C, water contaminated with MTBE entered through a body B inlet Bl and transversed the MTU longitudinally, in a spiral path, from left to right. Inlet Bl was structured to impart a spiral motion to a feed stream. The feed stream was propelled at a speed of approximately .3 mach The force and spiral configuration imparted to the feed stream of contaminated water longimdinally traversing the MTU chamber C resulted in a thin spiral ribbon of water formed against screens S within the MTU chamber. A removal stream, in the instant experiment fresh air, entered chamber C outside of the screens S through a pair of inlets RSI. Johnson screens were used in the experiment. The removal stream or air was sucked and/or pressured from the exterior of the chamber through the screens S, through the spiral ribbon of water, and into the radially central portion of the chamber. In the process the fresh air removal stream picked up and moved MTBE molecules, both being drawn to the interior of the chamber as a result of centrifugal force on the feed stream The fresh air with the MTBE exited the MTU chamber and through the central outlets CO at both ends of the chamber body. The feed stream exits the MTU longitudinal chamber at a peripheral outlet PO defined at one end of the chamber. Centrifugal force maintains the heavier water molecules against the confining screen S within the chamber. The lighter MTBE molecules together with the lighter air migrate to the center of the chamber in a known fashion The novel MTU results in an essentially instantaneous separation and removal of products from a continual supply of new material. Gas bubbles can be created with nearly molecular dimensions. A designed cavitation effect, illustrated in Figure 5E, can be utilized to clean structural surface. In the novel MTU it can be seen that the liquid phase is continuous wherein the gas phases is discontinuous. It can be seen that the novel MTU provides an extremely large interfacial transfer area, orders of magnitude larger than that of conventional equipment, hi such a manner the novel MTU can achieve a rapid approach to chemical equilibrium. A high capacity can be provided with significantly smaller volumes, size and weight. The smaller size and weight permits a lower cost, even with upon the use of expensive materials, as well as a greater capacity to handle high temperature and high pressure. Figure 5D illustrates a 3-inch diameter unit and a 6-inch diameter unit, the 6-inch diameter unit having a length of 7 Vz feet and the 3-inch diameter unit having a length of approximately 5 Vi feet. Table 1 illustrates anticipated MTU process parameters as function of diameter and length of the unit. In general the novel MTU is particularly suited for VOC removal, evaporation, absorption and gas liquid reactions. A few petrochemical applications include gas dehydration with TEG, gas sweetening with amines, sea water deoxygenation, refinery desalter effluent treatment, sour water treatment, bilge water/tank bottoms treatment, VOC stripping of produced water, alternative sulfer recovery processes, ground water remediation and chemical reaction processes. While preferred embodiments of this invention have been shown and described, modifications thereof can be made by one skilled in the art without departing from the spirit or teaching of this invention The embodiments described herein are exemplary only and are not limiting. Many variations and modifications of the system and apparatus are possible and are within the scope of the invention Accordingly, the scope of protection is not limited to the embodiments described herein, but is only limited by the claims which follow, the scope of which shall include all equivalents of the subject matter of the claims.

Claims

CLAIMS WHAT IS CLAIMED IS:

1. A method for use in the removal of acid gases comprising H₂, CO₂, S, COS, or mixtures thereof, from gaseous or liquid feed streams containing said acid gases, said method comprising: computer modeling of process conditions and requirements to select an aqueous solvent solution adapted for use in an adsorber/absorber unit, the aqueous solvent solution comprising at least 5% to 50% by weight of an alkaline silicate salt; the computer modeling taking into account a rate of adsorption and/or absorption in a mass transfer layer.

2. The method of claim 1, wherein the solution includes about 10% to about 50% by weight of an alkanolamine.

3. The method of claim 1, wherein a requirement is deep CO₂ removal and wherein the solution contains about 10% to 90% by weight MDEA and about 5% to 50% by weight by MEA.

4. The method of claim 2, wherein the solution includes a compound selected from the group consisting of heterocyclic amine condensates, alkyl amine condensate, oil-soluble alkyl arnine condensate, polymeric airiinoalcohol, and combinations thereof.

5. The method of claim 1 wherein the solution further includes at least one of a corrosion inhibitor, a stabilizer, an anti-foamer and a surfactant.

6. The method of claim 2 wherein the solution includes MDEA.

7. A method for removing sulfur and/or a sulfur containing compound from a sulfur- containing feed stream, comprising: computer modeling of process conditions and requirements to select an aqueous solvent solution adapted for use in an adsorber/absorber unit, the aqueous solvent solution being at least a tMophilic solution formed by combiriing at least an alkaline silicate salt, an alkanolamine and water, the computer modeling taking into account a rate of adsorption and/or absorption in a mass transfer layer upon contacting the sulfur/compound-containing stream with the tMophilic solution.

8. The method according to claim 7 wherein the sulfur/sulfur compound-containing stream comprises H₂S and gaseous hydrocarbons.

9. The method according to claims 1 or 7 wherein the solvent in the solution is in a 45% to 65% by weight concentration.

10. The method according to claims 2 or 7 wherein the solution is formed by combining at least 10% by weight MDEA.

11. A system for use in the removal of acid gases and sulfurs comprising H_2s CO₂, S, COS, sulfur compounds or mixtures thereof from gaseous or liquid streams containing said acid gases and/or sulfurs, said system comprising: contacting the acid gas and/or sulfur containing stream with a thiophilic solution formed by combining at least an alkaline silicate salt, an alkanolamine and water in an ICU, the ICU modeled by a mass transfer rate based computer program; and including designing at least a portion of the compound of the solution using the computer program.

12. A system for use in the removal of acid gases and/or sulfurs comprising H₂, CO , S, COS, sulfur compounds or mixtures thereof from gaseous or liquid streams containing said acid gases and/or siilfur, said system comprising: contacting the acid gas/sulfur containing stream with a tMophilic solution formed by combining an alkaline silicate salt, an alkanolamine and water in an arnine unit, the amine unit incorporating at least two pieces of equipment from the group consisting of an inlet separator, a treated gas knock out/coalescer, a flash tank separator, a particular filter, and a carbon purifier.

13. A system for the removal of acid gases and/or sulfur comprising H₂, CO₂, S, COS, H₂S, sulfur compounds or mixtures thereof, from gaseous or liquid streams containing said acid gases and/or sulfur, said system comprising: designing a performance solvent solution including at least an alkaline silicate salt and an alkanolamine and water for use in an adsorber/absorber unit; wherein the designing includes optimizing system equipment operating parameters and solvent composition using a computer model based upon mass transfer rates in the adsorber/absorber unit, the model taking into account input stream parameters and design requirements.

14. The system of claim 13 wherein the model takes into account output stream parameters.

15. The system of claim 13 wherein an operating parameter optimized includes solvent concentration.

16. The system of claim 13 wherein an operating parameter optimized includes lean solvent temperature.

17. The system of claim 13 wherein an operating parameter optimized includes circulation rate.

18. The system of claim 13 wherein an operating parameter optimized includes regenerator conditions.

19. The system of claim 13 wherein an operating parameter optimized includes reboiler conditions.

20. The method according to claims 1 or 7 wherein the alkaline silicate salt includes a sodium silicate.

21. The method according to claims 1 or 7 wherein a lean temperature of the solvent solution is 10-15 °F hotter than the feed stream.

22. The method according to claim 7 wherein the alkanolamine is selected from the group consisting of monodiethanolarnine (MDEA), diethanolamine (DEA), monoethanolamine (MEA), dusopropanolamine (DIPA), triethanolarnine (TEA) and combinations thereof.

23. The method according to claim 22 wherein the aJkanolamine contain MMEA.

24. The method of claims 1 or 7 wherein the lean solvent temperature is maintained below 130°F

25. The method of claims lor 7 wherein the lean solvent temperature is above 80 °F.

26. The method of claims 1 or 7 wherein the solution further includes at least one of a corrosion inhibitor, stabilizer, anti-foamer and a surfactant

27. A method for treating sour waters from a refinery comprising applying a solvent solution comprising an alkaline silicate salt and an alkanolamine to said sour waters.

28. A method for preheating process wastewater containing sulfur comprising applying a solvent solution comprising an alkaline silicate salt and an alkanolamine solution to said process wastewater.

29. A method for enhancing operation of a hydrogen sulfide scrubber comprising applying a solvent solution comprising an alkaline silicate salt and an alkanolamine solution to the scrubbing solution in said scrubber.

30. A method for improving operation of a geothermal power plant comprising applying a solvent solution comprising an alkaline silicate salt and an alkanolamine solution to condensed steam in said power plant along with an iron catalyst.

31. A method for improving the quality of industrial water comprising applying a solvent solution comprising an alkaline silicate salt and an alkanolamine solution to said industrial water.

32. A method for improving sludge treatment comprising applying an alkaline silicate salt and an alkanolamine solution to said sludge prior to sludge thickening.

33. A mass transfer unit (MTU) for use in the removal of contaminates from a contaminated feed sfrearn, mcluding a removal stream, comprising : an intimate contact unit (ICU) chamber extending between a contaminated stream inlet and a contaminated stream outlet; at least one inlet into the chamber for a mass transfer removal stream; at least one outlet from the chamber for the removal stream and contaminates; the inlet for the contanύhated feed stream structured to spiral a contarninated feed stream ribbon of fluid through the ICU chamber at a speed of approximately .3 mach; and the ICU chamber structured to pass the removal stream inwardly through the spiraling contaminated feed stream along substantially the whole length of the ICU chamber.

34. An MTU for separating chemical compounds in a fluid feed stream using a removal stream, comprising: a body defining a chamber having at least a significant longitudinal direction and axis; structure for directing one of a feed stream and a removal stream in a spiral path longitudinaly through the chamber; boundary structure adopted to substantially contain the one spfraling fluid stream in the longitudinal direction through the chamber while permitting the other of a feed stream and a removal stream to pass through the boundary structure in the chamber toward the longitundinal axis in an inward direction; and a removal stream inlet and outlet, structured with the body to circulate the removal stream through and out of the chamber, including through the boundary structure in a radially inward direction.

35. The MTU of claim 34 wherein the one stream spirals through the chamber at a speed of approximately .3 mach.