RU2569093C2 - Removing carbon dioxide from waste streams through combined production of carbonate and/or bicarbonate minerals - Google Patents

Abstract

FIELD: chemistry.

SUBSTANCE: invention relates to removing carbon dioxide and other pollutants from a waste streams. A method of removing carbon dioxide contained in a gas stream includes the following steps: (a) obtaining sodium hydroxide in an aqueous mixture; (b) mixing the sodium hydroxide with carbon dioxide obtained from a source to obtain bicarbonate products or a combination of carbonate and bicarbonate products in a mixture; and (c) separating said products from the mixture, thereby removing carbon dioxide from the solid mineral product which contains 70 wt % sodium bicarbonate; where the hydroxide is obtained by: obtaining a sodium salt; mixing the salt with an acid and water, an acid and steam or an acid, water and steam to obtain protonated brine; and performing electrolysis of the protonated brine to obtain sodium hydroxide; where step (b) includes shifting the product equilibrium in favour of obtaining bicarbonate products.

EFFECT: high efficiency of removing carbon dioxide gas and high output of carbonate.

25 cl, 20 dwg, 3 tbl

Description

CROSS RELATIONS TO RELATED APPLICATIONS

This application claims the priority of provisional application for US patent 60/718906, entitled "Removing carbon dioxide from waste streams through the joint formation of synthetic carbonate minerals", issued September 20, 2005; US Provisional Application 60/642698, issued January 10, 2005; and provisional patent application US 60/612355, issued September 23, 2004. The full text of each of the above applications (including appendices) with the claims is given here, in particular, by appropriate reference to them.

BACKGROUND

1. The technical field to which the invention relates.

The present invention generally relates to the removal of carbon dioxide and, in addition, other pollutants from waste streams. More specifically, the present invention relates to the removal of carbon dioxide and other pollutants from waste streams by absorbing carbon dioxide and other pollutants from concentrated flue gas streams and then co-producing carbonate and / or bicarbonate materials that trap and neutralize any absorbed associated pollutants .

2. The prior art

Over the past four decades, industrial air emissions have been a growing concern at the national and international levels, both in the private and industrial sectors. In particular, attention was focused on gases that cause the greenhouse effect, which have the property of retaining solar heat in the atmosphere, leading to the "greenhouse effect". The greenhouse effect occurs when the heat received from the sun is absorbed in the atmosphere and hydrosphere of the earth, increasing the average temperature of the atmosphere, the temperature of the ocean, and other average and average temperatures of the planet, until the problem of climate change; this effect is usually associated with the effect on the heat balance of the Earth, although there is no consensus on the speeds, the degree to which materials burned by humans, and the degree, direction and magnitude of the effect. Despite the divergence in estimates, it is obvious to everyone that the removal of CO ₂ (and other chemicals) from the main sources of emissions will be beneficial if the costs are quite small.

The greenhouse gases are mainly carbon dioxide and are produced in urban power plants and large-scale industries at their location, however, they also form during any conventional carbon burning (for example, in cars, when clearing rainforests, just burning, and so on), although their most concentrated point emissions occur around the world in power plants, making the solution to the problem of reducing or removing emissions from these specific places attractive to The disposal technology. Since energy production is the main cause of greenhouse gas emissions, over the past thirty years, various methods have been investigated and intensively studied, such as reducing the carbon load, increasing efficiency, and sequestering carbon from the flue gases of a power plant.

Reducing the use of carbon involves the use of alternative carbon-free energy sources, such as nuclear, hydroelectric, photovoltaic, geothermal, and other sources of electricity, to reduce the proportion of energy produced solely from burning carbon. While each of these methods of energy production continues to contribute to the total volume of energy production, forecasts of the growth rate of global electricity consumption exceed the growth rates of energy production by these methods. Therefore, it is expected that greenhouse gas emissions will increase, despite increased use of non-carbon energy sources.

Efficiency gains typically came down to ways to improve carbon burning by pre-burning, decarbonizing, burning with oxygen, and so on, first reducing the amount of CO ₂ produced and then oxidizing all potential pollutants as completely as possible. In addition, the method leads to an increased amount of carbon dioxide emission energy released to increase efficiency. While advances in this area have led to an increase in combustion efficiency, there are few achievements in the field of heat recovery.

Attempts to sequester carbon (in the original form of gaseous CO ₂ ) have led to the development of many different methods that can be generally classified as geological, terrestrial, or oceanic systems. These methods are mainly associated with the transfer of the resulting carbon dioxide to real-world locations and the injection of carbon dioxide into geological, underground, or oceanic storage facilities. Each of these disposal methods is expensive for preparing CO ₂ for transportation, transporting, and pumping to a “carbon storage". Essentially, these methods are not economically feasible, and in many cases consume more energy than the original carbon produces.

Removal can also be carried out by various industrial methods, which include scrubbing, membranes, cheap O ₂ , and hydrates. However, the disadvantages of each of these technologies are large investments that increase to unprofitable levels and the extremely high impact of CO ₂ removal on the cost of electricity.

There is no reason to believe that all the disadvantages are listed above, rather they are one of many that can reduce the effectiveness of previously known methods for removing carbon dioxide from waste streams; however, the disadvantages mentioned here sufficiently demonstrate that all approaches known in the art are not satisfactory, and that there is a great need for the methods disclosed and claimed in this description.

SUMMARY OF THE INVENTION

The present invention provides methods and apparatus for removing carbon dioxide from waste streams that solve the problems described above. In general, the invention relates to methods for removing carbon dioxide from a gas stream by using a hydroxide compound to form carbonate and / or bicarbonate. Carbonate and / or bicarbonate can then be used in any various ways or simply discarded.

Various embodiments of the present invention provide advantages over currently known techniques for removing carbon dioxide from waste streams. A few potential advantages that are achieved in accordance with the invention are as follows.

Unlike processes that physically physically remove carbon to a distant place, the combined production of sodium carbonates and / or bicarbonates on an industrial scale by synthesis from sodium chloride and carbon dioxide converts unwanted carbon dioxide into chemical products directly at the energy production site, potentially eliminating the cost of transportation to the place of extraction.

Unlike other methods for decarbonizing flue gas streams that are not amenable to modernization, embodiments of the present invention can be used to modernize existing power plants, while significantly reducing the investment required to implement the decarbonization method. In addition, the decarbonization method is scalable, and can be implemented on a scale: pilot installation - semi-industrial installation - industrial installation, due to a gradual increase in the number of reactors.

Unlike other methods known in the art, the decarbonization method according to certain embodiments allows the extraction of carbon dioxide in the form of economically useful chemical products while producing useful by-products such as chlorine gas, sodium carbonate, and hydrogen gas. Since by-products of the decarbonization process are economically sought after, their costs offset the cost of extraction, and in properly designed systems, the carbon sequestration method is potentially beneficial in itself.

Thanks to the appropriate method for actively scrubbing the exhaust flue gases, other undesirable acidic pollutants, such as SOx, NOx, HgOx, and others, are actively extracted in the process. In addition, the scrubbing method can entrain and / or entrain other components of the gas stream and / or contaminants into sodium carbonate (e.g., coal ash, etc.), thus removing them from the gas stream.

Specific embodiments of the present invention include a method for removing carbon dioxide from a gas stream, comprising: producing hydroxide in an aqueous mixture; mixing the hydroxide with a gas stream to produce carbonate products (defined as products containing a carbonate group, CO ₃ ), bicarbonate products (defined as products containing a bicarbonate group, HCO ₃ ), or a mixture of carbonate and bicarbonate products in a mixture; and separating said carbonate and / or bicarbonate products from the mixture, thereby removing carbon dioxide from the gas stream. The hydroxide can be any form of hydroxide, including, but not limited to, sodium hydroxide, potassium hydroxide, calcium hydroxide, magnesium hydroxide, and aluminum hydroxide. Specialists in this field understand that it is possible to carry out a similar chemical interaction and decarbonization with any number of hydroxides or mixtures of hydroxides. In a number of preferred embodiments, the hydroxide is sodium hydroxide.

In specific embodiments, implementation, the method further includes a process of regulating the method in order to obtain mainly only carbonate products, or mainly only bicarbonate products. In a number of embodiments, the method further includes a process control process to produce a mixture of carbonate and bicarbonate products, a mixture of which may consist of approximately X% carbonate and Y% bicarbonate, wherein XY combinations may be any of the following: 1-99, 2-98 , 3-97, 4-96, 5-95, 6-94, 7-93, 8-92, 9-91, 10-90, 15-85, 20-80, 25-75, 30-70, 35 -65, 40-60, 45-55, 50-50, 55-45, 60-40, 65-35, 70-30, 75-25, 80-20, 85-15, 90-10, 91-9 , 92-8, 93-7, 94-6, 95-5, 96-4, 97-3, 98-2 or 99-1.

In certain embodiments, mixing occurs in two separate chambers, one chamber being used to produce carbonate products and the other chamber being used to produce bicarbonate products. In other embodiments, the mixing occurs in a bubble column or series of bubble columns. In still other embodiments, the recovery of carbonate and / or bicarbonate products from the mixture comprises a hot precipitation recovery method. In a number of embodiments, heat for the separation method is obtained as a result of heat exchange with incoming flue gases. Upon isolation, the separated carbonate can be in the form of an aqueous suspension or in the form of a solution of hydroxide, carbonate, and water at various concentrations, and in this form it can then be dried by any of a variety of methods. In some embodiments, the carbonate may not be dried. For example, a suspension of sodium carbonate can be used to treat hard water. There is no doubt that specialists in this field are aware of the numerous applications of carbonate obtained by the method of the invention, for example, mixtures of suspensions of sodium bicarbonate and sodium carbonate can be suspended in a tanker for use in various forms of production of detergent, in the manufacture of glass as a flux, and t .d., as well as for the above use in the water treatment process.

In specific embodiments, the method further comprises transporting carbonate products to a remote burial site; combining carbonate products with acid in a neutralization reaction to produce pure carbon dioxide; and pumping carbon dioxide into a carbon storage. In other embodiments, other gas stream components are neutralized and / or entrained / captured in the carbonate production process, including SOx, NOx, and mercury-containing material

In some embodiments, implementation, the preparation of hydroxide includes: obtaining a salt; mixing salt with water, steam, or both, to form a solution; and electrolysis of the solution to produce hydroxide. In specific embodiments, the solution is electrolyzed at a voltage of greater than or equal to 5 volts, while in other embodiments, the solution is electrolyzed at a voltage of less than 5 volts. In a number of embodiments, the solution is subjected to electrolysis at voltages from 1 V to 5 V, including 1.5 V, 2.0 V, 2.5 V, 3.0 V, 3.5 V, 4.0 V, or 4 , 5 in, or in any interval between any of these points.

In certain embodiments, the acid is added to the solution before it is electrolyzed. The acid may be any acid that can provide protonation of the solution, including, but not limited to, hydrochloric acid. For specialists in this field it is obvious that it is possible to carry out a similar chemical transformation and electrolysis with any number of acids or mixtures of acids. In a number of preferred embodiments, the acid is hydrochloric acid. In other embodiments, the amount of acid that is added to the solution is based on determining the optimal degree of protonation at which the lowest energy is expended in the preparation of the reagents and the highest energy is recovered from the products.

In still other embodiments, the electrolysis step is carried out in an electrolyzer having a catholyte space and anolyte space, and carbonate and / or bicarbonate products are returned to the catholytic space of the electrolyzer. In other embodiments, implementation, the energy required for this method is supplemented by the waste heat of the gas stream.

Other embodiments of the invention include a method for removing carbon dioxide from a gas stream, comprising: obtaining sodium hydroxide in an aqueous mixture; mixing sodium hydroxide with a gas stream to produce sodium carbonate, sodium bicarbonate, or a mixture of sodium carbonate and sodium bicarbonate; and recovering said sodium carbonate and / or sodium bicarbonate from the mixture, thereby removing carbon dioxide from the gas stream

In a number of embodiments, the method further includes a process for regulating the method in order to obtain substantially only sodium carbonate, or substantially only sodium bicarbonate. In other embodiments, the method further includes a process control process to produce a mixture of sodium carbonate and sodium bicarbonate, which may consist of X% sodium carbonate and Y% sodium bicarbonate, wherein XY combinations may be any of the following: 1-99, 2- 98, 3-97, 4-96, 5-95, 6-94, 7-93, 8-92, 9-91, 10-90, 15-85, 20-80, 25-75, 30-70, 35-65, 40-60, 45-55, 50-50, 55-45, 60-40, 65-35, 70-30, 75-25, 80-20, 85-15, 90-10, 91- 9, 92-8, 93-7, 94-6, 95-5, 96-4, 97-3, 98-2 or 99-1.

In specific embodiments, the mixing takes place in two separate chambers, one chamber being used to produce sodium carbonate and the other chamber being used to produce sodium bicarbonate. In other embodiments, mixing occurs in a bubble column or in a series of bubble columns. In still other embodiments, the recovery of sodium carbonate and / or sodium bicarbonate from the mixture comprises a hot precipitation recovery method. In a number of embodiments, heat for the separation method is obtained by heat exchange with incoming flue gases.

In specific embodiments, the method further comprises: transporting sodium carbonate to a remote burial site; combining carbonate products with acid in a neutralization reaction to produce pure carbon dioxide; and pumping carbon dioxide into a carbon storage.

In some embodiments, implementation, the production of sodium hydroxide includes: obtaining sodium chloride; mixing sodium chloride with water, steam, or both, to obtain a brine; and electrolysis of the brine to produce sodium hydroxide and chlorine gas. In specific embodiments, the brine is electrolyzed at a voltage of greater than or equal to about 5 volts, while in other embodiments, the electrolysis is carried out at a voltage of less than 5 volts. In some embodiments, implementation, the solution is subjected to electrolysis at a voltage of from 1 to 5 V, including 1.5 V, 2.0 V, 2.5 V, 3.0 V, 3.5 V, 4.0 V, or 4 , 5 in, or in any interval between any of these points.

In some embodiments, acid is added to the brine before electrolysis. An acid can be any acid that can provide a protonation solution, including, but not limited to, hydrochloric acid. For specialists in this field it is obvious that it is possible to provide a similar chemical interaction and electrolysis with any number of acids or mixtures of acids. In some preferred embodiments, the acid is hydrochloric acid. In still other embodiments, the amount of acid added to the brine is based on the optimal degree of protonation at which the lowest energy is expended in the preparation of the reactants and the highest energy is recovered from the products.

In some embodiments, the electrolysis step occurs in an electrolysis having a catholyte space and anolyte space, and sodium carbonate and / or sodium bicarbonate is recycled to the catholyte space of the electrolyzer. In other embodiments, the energy required to implement the method is added due to the waste heat extracted from the gas stream. In still other embodiments, the method further comprises collecting chlorine gas, while in other embodiments, hydrogen gas is obtained. In some embodiments, hydrogen gas and chlorine gas are reacted to produce hydrochloric acid, which is added to the brine before electrolysis. In other embodiments, hydrogen gas is burned with atmospheric oxygen or oxygen from a chemical storage to produce water, while in other embodiments, methods include using hydrogen gas to generate energy. In some embodiments, the separation of sodium carbonate and / or sodium bicarbonate from the mixture comprises a hot precipitation separation method in which heat for the separation method is derived from energy produced by hydrogen gas. In other embodiments, hydrogen gas is burned with coal to improve emissions from burning coal, or is used to burn in a fuel cell to produce a constant electric current.

In some embodiments, the gas stream is a plant exhaust stream, while in other embodiments, the plant is a power plant that uses coal as a fuel source. In some embodiments, the exhaust stream includes CO ₂ and H ₂ O.

Specific embodiments of the present invention also include an installation, including, when using: an electrolysis chamber, comprising at least one cathode and at least one anode, a hydroxide production chamber; mixing equipment operably connected to the electrolysis chamber and a gas flow conduit; mixing equipment for mixing hydroxide from the electrolysis chamber with the gas stream to produce a mixture in which carbon, sulfur, and / or nitrogen compounds in the gas stream can react with hydroxide; and a separation chamber operatively associated with the mixing equipment and adapted to separate the mixture into separate gas phases and solid and / or liquid phases.

In some embodiments, the implementation, the electrolysis chamber includes a membrane cell, a diaphragm cell, and / or a mercury cathode cell. In specific embodiments, the mixing equipment is a batch reactor or a series of batch reactors, while in other embodiments, the mixing chamber is a chemical gas absorption apparatus or a series of such apparatuses. In other embodiments, the mixing chamber is a crystallization apparatus or a series of crystallization apparatuses, while in other embodiments, it is a bubble column or a series of bubble columns.

In specific embodiments, the installation further includes a drying chamber operatively associated with a separation chamber adapted to be used to remove liquid from the solid and / or liquid phase, while in other embodiments, the drying chamber, when used, is adapted to heat solid and / or liquid phase. In still other embodiments, the installation is further defined as being operatively associated with a power plant.

In some embodiments, the implementation, the electrolysis chamber in use is adapted to produce gaseous chlorine and sodium hydroxide from sodium chloride and water. In other embodiments, the mixing equipment, in use, is adapted to mix hydroxide from an electrolysis chamber with carbon dioxide from a gas stream to produce carbonate and / or bicarbonate products.

In still other embodiments, the present invention includes a method for determining the optimum operating voltage and magnitude of the electric current of a cell during its low voltage operation with respect to the region of increased values for a given V / I characteristic at a given degree of protonation. In other embodiments, the invention includes a method for determining a lower thermodynamic limit for an operating voltage for a given cell used in the method. In specific embodiments, a method for determining environmental performance (∂CO ₂ / ∂E) in general is provided for devices that remove CO ₂ from waste streams, while other embodiments include a method for determining environmental performance (∂CO ₂ / ∂ E) specifically for devices that are used in the present invention in any of its embodiments. Other embodiments include a method for producing extremely pure hydrogen gas at a low price of indifference, its cost equal to the content of recoverable energy.

The terms “includes” (in any form), “has” (in any form), “contains” (in any form), and “concludes” (in any form) are non-limiting conjugation verbs. As a result, a method or installation that "includes," "has," "contains," or "includes" one or more stages, or elements, has these one or more stages or elements, but is not limited to having only one or more stages or elements. Likewise, an element of a device or method that "includes," "has," "contains," or "encloses" one or more features, has this one or more features, but is not limited to having only one or more features. The term "use" should be interpreted in the same way. Therefore, for example, the stage of the method, which includes the "use" of specific information, means that at least the stated information is used, but the possibility of using other, not given information is not excluded. In addition, a structure that is formed in some way must be formed in at least this way, but can also be formed in a way or methods that are not indicated.

The indefinite article refers to one or more than one subject, unless expressly stated otherwise in the description of the invention. The term “other” means at least a second or more. The terms "mainly" and "about" mean at least a value close to a given value (and including it) or condition (preferably within 10%, more preferably within 1%, and most preferably within 0 ,one%).

As used herein, the terms “carbonates” or “carbonate products” generally refer to mineral components containing a carbonate group, CO ₃ . Therefore, the terms encompass both carbonate / bicarbonate mixtures and compounds containing exclusively a carbonate ion. The terms “bicarbonates” and “bicarbonate products” generally refer to mineral components containing the bicarbonate group HCO ₃ . Therefore, the terms encompass both carbonate / bicarbonate mixtures and compounds containing exclusively bicarbonate ion.

In the formation of bicarbonates and carbonates in some embodiments of the present invention, the term "ionic ratio" refers to the ratio of sodium ions in a product to the number of carbon atoms present in this product. Therefore, it can be indicated that the product stream formed from pure bicarbonate (NaHCO ₃ ) has an "ionic ratio" of 1.0 (Na / C), but it can be indicated that the product stream formed from pure carbonate (Na ₂ CO ₃ ), has an "ionic ratio" of 2.0 (Na / C). In order to expand the range, it can be indicated that an infinite number of continuous mixtures of carbonate and bicarbonate have ionic ratios ranging from 1.0 to 2.0.

In some preferred embodiments of the invention, hydrochloric acid is added to the feed brine of sodium chloride in an electrolytic cell used to produce chlorine and sodium hydroxide, causing the following reaction to occur:

H ₂ O + NaCl + aHCl + energy → NaOH + (1/2 + a / 2) H ₂ + (1/2 + a / 2) Cl ₂

In this equation, the symbol “a” stands for “degree of protonation”, and it is the ratio of protons (H ⁺ ions) to sodium ions (Na ⁺ ions).

As used herein, the term "sequestration" generally means methods or applications that partially or completely remove CO ₂ from emissions of point sources and allow you to store this CO ₂ in some form in order to prevent its return to the atmosphere. The use of this term does not exclude the consideration of any form of the described embodiments as “sequestration” methods.

As used herein, the term “environmental efficiency” is used synonymously with the term “thermodynamic efficiency” and refers to the amount of CO ₂ sequestered in a particular embodiment of the present invention for the energy expended (represented by the equation ∂CO ₂ / ∂E). Sequestration of CO ₂ expressed as a percentage of the total amount of CO ₂ installation; power consumption is also expressed as the total power consumption of the installation.

As used herein, the terms "low voltage electrolysis" and "NVE" mean electrolysis at voltages below 5 volts.

Descriptions of known processing methods, components, and equipment are omitted so as not to overshadow the present methods and devices with unnecessary details. Descriptions of the present methods and devices, including their descriptions in applications, are illustrative examples and do not limit the invention. For those skilled in the art, based on this description of the invention, some substitutions, modifications, additions and / or permutations may be apparent that fall within the scope of the patent claims, but are not explicitly listed in this description.

Other features and related advantages will become apparent as a result of reading the following detailed description of specific embodiments in conjunction with the accompanying drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

The following drawings are illustrative and are not limiting. The drawings are part of the present description, and they are included to further demonstrate some aspects of the present invention. The invention may become more apparent by reference to one or more of these drawings in combination with the description of illustrative embodiments presented herein:

In FIG. 1 is a schematic flow diagram illustrating specific embodiments of the present invention.

In FIG. 2A is a setup for illustrating the main features of one embodiment of the decarbonization portion of the present invention.

In FIG. 2B presents the experimental results of absorption / conversion.

In FIG. 2C shows experimental results for absorption / conversion.

In FIG. 2D is a graphical plot showing the gas / liquid contact distance (m, liquid depth) required to remove 90% CO ₂ .

In FIG. 2E shows a graphical plot of the ionic ratio in the product versus the percentage of CO ₂ absorbed in the test reactor.

In FIG. Figure 3 shows a graphical dependence of the temperature regime of a liquid in a reactor chamber depending on the indicated reaction time.

In FIG. 4 is a diagram illustrating a flooding study of column 5 '.

In FIG. Figure 5 shows typical operating curves of the current – voltage characteristics for various anolyte pH and temperatures.

In FIG. 6 is a block diagram of a system including a reactor in which hydrogen is not recovered in accordance with embodiments of the present invention.

In FIG. 7 is a block diagram of a system including a reactor in which, in accordance with embodiments of the present invention, hydrogen is recirculated to a fuel cell to produce direct current.

In FIG. Figure 8 shows the graphical dependence of the percentage of absorption of CO ₂ in a bubble column as a function of the depth of the liquid and the surface velocity of the gas at low surface velocities.

In FIG. 9A is a diagram showing maximum theoretical CO ₂ absorption and experimental results according to embodiments of the present invention.

In FIG. 9B provides indicative calculations and calculations for a flue gas model for a plant including specific embodiments of the present invention.

In FIG. 9C shows a decarbonizer charge and intermediate sodium hydroxide requirement for a plant comprising specific embodiments of the present invention.

In FIG. 9D shows the loading of the electrolysis section and the costs for a plant incorporating specific embodiments of the present invention.

In FIG. 9E presents waste heat calculations for a plant incorporating specific embodiments of the present invention.

In FIG. 9F provides calculations of energy balance and environmental performance for a plant incorporating specific embodiments of the present invention.

In FIG. 10 shows the environmental performance of various simulated power plants, including specific embodiments of the present invention.

In FIG. 11 is a graph showing the percentage of power saved per m ^{2 of} area for a plant with normalized NEV.

DESCRIPTION OF ILLUSTRATIVE EMBODIMENTS

The present invention relates to carbon sequestration methods in which carbon dioxide is removed from waste streams and converted to carbonate and / or bicarbonate products. Embodiments of the methods and installations of the invention include one or more of the following common components: (1) an aqueous decarbonization process in which gaseous CO _{2 is} absorbed by an aqueous caustic mixture and then reacted with a hydroxide to form carbonate and / or bicarbonate products; (2) a separation process in which carbonate and / or bicarbonate products are separated from the liquid mixture; (3) the process of electrolysis of brine to obtain sodium hydroxide, which is used as a liquid absorbent in the decarbonization process; and (4) the production and use of by-products of decarbonization and electrolysis processes, including gaseous chlorine, sodium carbonate and bicarbonate, and hydrogen gas. Each of these common components is explained in more detail below.

I. Benefit Overview

Like any other method or installation that allows you to achieve your goal, many embodiments of the present invention consume energy to absorb CO ₂ and other chemical products from flue gas streams and to achieve the other objectives of the embodiments of the present invention described here. However, one advantage of specific embodiments of the present invention is that they provide environmental performance that is superior to what is already known in the art, as explained in detail in Examples 5 and 6. As is evident from the data of Examples 5 and 6, additional waste heat recovery or energy production for energy supply of the method without the manifestation of the greenhouse effect, the use of distributed low-voltage electrolysis, and increasing the return of electrical energy when extracting hydrogen energy, ozhet further improve the environmental effectiveness of the process, up to the value or above, wherein the method is powered entirely by utilizing waste heat (and extracting energy from hydrogen) and provides virtually 100% absorption of emitted power plant CO _2.

In addition, one advantage of using the chemistry of the final interaction in specific embodiments of the implementation of the present invention, such as the absorption of weakly acidic CO ₂ , is that the method actually completely absorbs strong acids, SOx and NOx, and, to a lesser extent, mercury. Studies using the loading of SOx / Ar and NOx / Ar in single-stage decarbonizers showed 99% + removal of these components of the flue gas ("99% +" means that the presence of each pollutant in the case of processing the exhaust gas at a speed of 14 l / min was not detected in the final air flow by gas chromatography, that is, they were effectively removed). In specific embodiments of the present invention, concomitant scrubbing of NOx, SOx, and mercury compounds may be of greater economic value; that is, as a result of using embodiments of the present invention, coals that contain large amounts of these compounds can be burned in a power plant, in some embodiments, with less polluting effect than high-grade coals burned without taking advantage _{of the} specific CO ₂ absorption method embodiments of the present invention.

In addition, the scalability of specific embodiments of the present invention can be extremely large; that is, since, in specific embodiments, the method controls the power consumption, the power consumption can actually be commensurate with the molecule of a particular absorbent obtained at any given time. In addition, it is possible to easily determine the exact amount of absorbed CO ₂ in practice: weigh the resulting carbonate / bicarbonate products, determine their ionic ratio by analysis, calculate to determine the moles of absorbed CO ₂ , and the amount of absorbed CO _{2 is} easily determined and confirmed as a result (a factor that may have the added benefit of intensifying the removal of CO ₂ and other chemicals from the flue gas).

Another additional advantage of specific embodiments of the present invention that distinguishes them from other CO ₂ removal methods is that, under certain market conditions, the cost of the products will significantly exceed the cost of the spent reagents, or the cost of the consumed electricity or depreciation costs for the installation. In other words, specific embodiments are industrial methods for cost-effectively producing chlorine, hydrogen, and carbonate products, while removing a significant amount of CO ₂ and associated contaminants. The implementation of all other competing methods is simply associated with additional operating costs.

II. Schematic diagram

In FIG. 1 depicts a simplified schematic flow diagram illustrating examples of main embodiments of the installations and methods of the present description. This scheme is offered for illustrative purposes only, and therefore it only depicts specific embodiments of the present invention, and in no way limits the scope of the claims. As shown in FIG. 1 , flue gas (TG) enters the process at point 901 , usually after it transfers the waste heat to a system that generates direct current from the waste heat. The TG introduced in this example as a gas mixture with a temperature of 300 ° C is first fed through a pipe 902 to the TG / bicarbonate 903 heat exchanger, in which the temperature of the TG in this example decreases to 120-140 ° C. Similarly, the VT passes through the anolyte / TG 904 heat exchanger and the catholyte / TG 905 heat exchanger, with a decrease in its temperature to 95 ° C, and then through the water / FG 929 heat exchanger, with a decrease in its temperature to an additional 30 ° C. The TG leaving the water / TG 929 heat exchanger is then introduced into the valve system, the temperature control system of the flue gas mixture 931 , in which the TG with a temperature of 30 ° C can be mixed with the flue gas with a temperature of 120-140 ° C, also supplied to the system adjusting the temperature of the flue gas mixture 931 by means of the hot flue gas treatment pipe 906 . TG mixtures with a temperature of 30-140 ° C can then be differentially introduced into the lower part of the carbonization absorption column 907 , which can be a packed or non-packed bubble column in which gas is injected or sparged, as a result of which gas forms bubbles that rise upward through fluid collecting in the upper outlet 908 . In this embodiment, the partially or fully decarbonated gas is then introduced and passed through the bicarbonate conversion column 909 , bubbled through the liquid in this column, removed further by a fan, and supplied to the outlet 910 .

The flow used in the TG / bicarbonate 903 heat exchanger is sodium bicarbonate / carbonate and sulfate, nitrate, mercury, and particles and aerosols absorbed from the flue gas in the absorption / conversion columns ( 907 and 909 ). By contacting it with an incoming TG with a temperature of 300 ° C, this liquid stream is heated to temperatures sufficient to generate water vapor with significant pressure, producing steam when the stream is introduced into the reservoir 911 , which is then condensed in a condenser 912 , to obtain distilled water, which is recycled to the regenerated H ₂ O 913 tank, which is then used, after any necessary treatment, to obtain brine in a brine mixer 914 . The flow used in the anolyte / TG 904 heat exchanger is a brine obtained by adding salts of group-1 and group-2 (in this example, NaCl) to water supplied either from a water supply from a water main 915, or partially or completely supplied from a tank regenerated H ₂ O 913 . This solution is protonated (acidified) by adding HCl in the form of gaseous HCl, absorbed by water, or from HCl from a warehouse, which is adjusted using a pH controller with a feedback of 916 . This stream circulates through the zone of anolyte 917 of electrolysis cell 933 . Similarly, the stream used in the catholyte / TG 905 heat exchanger is NaOH (aqueous), which circulates through the catholyte zone 918 of electrolyzer 933 . When the pH of the catholyte exceeds the minimum pH value (characterizing the concentration) at the measuring point 919 , concentrated NaOH is fed to the hydroxide storage tank 920 .

The stream used in the water / TG 929 heat exchanger comes from a sufficiently large water tank at a sufficiently low temperature to effect heat transfer. In some embodiments, this heat exchange system can be used as a “preheat” of bicarbonate / carbonate solutions before further heat exchange in a flue gas / bicarbonate 905 heat exchanger. In addition, in some embodiments, the injection into the reservoir of water from the water main 915 , the reservoir for storage of H ₂ O HX 937 , and the reservoir of regenerated H ₂ O 913 may be partially or completely combined.

The protonated brine circulating through the zone of anolyte 917 of electrolyzer 933 is electrolyzed to form chlorine gas, which is collected and transferred along chlorine gas line 921 , in this example, to sodium hypochlorite reactor 924 . Sodium ions and hydrogen ions (protons) are transferred through the membrane of the electrolyzer 933 to the catholyte zone 918 . There, sodium ions replace hydrogen ions in water, which leads to the formation of gaseous hydrogen, which is removed through the line for introducing pure gaseous hydrogen 922 into the H ₂ / O ₂ 923 fuel cell, where it is burned in atmospheric O ₂ to produce a constant electric current PT, which served in this example to the electrolyzer 933 , and pure water, which is recycled through a closed cycle of extraction of pure water 935 in the tank regenerated H ₂ O 913 . Chlorine gas supplied to the sodium hypochlorite reactor 924 is contacted (sparged) through sodium hydroxide supplied to the reactor from the hydroxide storage tank 920 . As a result, sodium hypochlorite is formed and it is stored for sale or further use as a feedstock. A certain amount of gaseous chlorine and HCl formed (using an excess of stoichiometric a, will create a continuous HCl recycling, excluding loss compensation) is burned here to produce HCl in fuel cell 925 , which is then recycled through the return line of acid gas HCl 926 to the brine mixer 914 .

The hydroxide produced or stored, or obtained from the starting chemicals and loaded into the hydroxide storage tank 920 , is used as an absorbent stream introduced into the carbonization / absorption column 907 . It is then passed through a bicarbonization / conversion column 909 and then sent (as a mixture of bicarbonate / carbonate in water) to a TG / bicarbonate 903 heat exchanger. After water is removed by evaporation, the concentrated suspension of bicarbonate / carbonate product is sent to the product bicarbonate tank 927 , where it can be dried for further processing or purification, or it can be sent for disposal or for sale.

Each of the above and shown in FIG. 1, the general components of the installations and methods of the present description are described in more detail below.

III. Water decarbonization (absorption) of CO ₂  from waste streams and its conversion to carbonate and bicarbonate

As noted above, in specific embodiments, the plants and methods of the present disclosure use an aqueous decarbonization process in which gaseous CO _{2 is} absorbed by an aqueous mixture of sodium hydroxide and then reacts with the hydroxide to form carbonate and bicarbonate products. In many embodiments of the present invention, unlike other carbon capture / sequestration schemes, sodium hydroxide is used as the main absorbent. Sodium hydroxide in various concentrations is known as an readily available CO ₂ absorbent. When carbon dioxide is contacted with aqueous sodium hydroxide, a continuous series of products can form, from pure sodium bicarbonate (NaHCO ₃ ) to pure sodium carbonate (Na ₂ CO ₃ ), and various conditions can be created to shift the equilibrium in any direction, even completely (or almost completely), and to precipitate bicarbonate, carbonate, or a mixed precipitate containing both compounds at a sufficient concentration (either as a result of chemical treatment or removal of water by various means).

When carbon dioxide is contacted with aqueous sodium hydroxide, the flow within the reaction chamber corresponds to the behavior shown in FIG. 3, as the reaction proceeds for the indicated time. Two stages of a sharp increase in temperature correspond to and determine two different reaction modes:

(1) A primary absorption step in which CO _{2 is} readily absorbed. The absorption capacity of the stream decreases as OH concentration decreases, and absorption ceases, and in some cases desorption begins when the OH content is completely exhausted. During this stage, the reaction is exothermic and almost exclusively carbonate is formed.

(2) A secondary conversion step in which CO _{2 is not} readily absorbed. The passage of the flue gas through the mixture does not lead to any significant absorption of CO _{2 by the} stream, but the stream cools significantly as a result of the heat loss of evaporation due to the evaporation of water, any loss of CO ₂ in the vapor state, and any endothermic reactions occur. During this stage, sodium carbonate already formed in the solution is converted to sodium bicarbonate, in accordance with the total stoichiometry of the reaction:

Na ₂ CO ₃ (water) + H ₂ O (W) + CO ₂ (water) → 2NaHCO ₃ (water)

This sequence, first carbonization and then bicarbonization, is confirmed by repeated operation of the apparatus shown in FIG. 2A (as explained in detail in Example 3), for and beyond absorption limits for liquids with different absorbent concentrations.

These two stages differ in the characteristics given in table 1 below.

TABLE 1 Stage Thermodynamics Product Absorption of CO ₂ Presence [OH] Carbonization Exothermic Na ₂ CO ₃ Intense Excess Bicarbonization Endothermic NaHCO ₃ Reduced to zero or negative Minimum

In embodiments of the present invention, the same chamber can be used to carry out two in situ processes, but the different nature of the reactions suggests that isolating the reactions into two chambers and separately optimizing them is an appropriate way for preferred embodiments. Regardless of the "internal arrangement" of transmission devices (that is, the degree of periodicity or continuity, a small number of chambers, capacities, stages, etc.), two fundamental processes occur at this molar concentration sufficient to ensure good absorption.

Therefore, since in some embodiments of the present methods and systems control processes for the production of pure and almost pure sodium bicarbonate can be implemented, in some embodiments of the present invention, therefore, one carbon atom is absorbed by one sodium atom produced by electrolysis, instead of ½ carbon atom (increase in achieved environmental efficiency by a nominal factor of 2 compared with carbonate production). Thus, the amount of energy for electrolysis and processing spent on obtaining a mole of hydroxide has twice as much "normal" absorption capacity when used to produce bicarbonate, compared with the absorption / energy efficiency in the case of carbonate production.

In various embodiments of the present invention, the entire range of bicarbonate / carbonate concentration ratios can be obtained. In preferred embodiments, concentrations, temperatures, pressures, space velocities, and so on, flows can be controlled to optimize the proportion of absorbed "available" CO ₂ to optimize bicarbonate formation.

In several embodiments of the invention, as a method of controlling the absorption rate of CO ₂ and other gases, pH control of the liquid absorbent (OH ion concentration) can be used. In some preferred embodiments, an increased salt / carbonate concentration may be used to further direct the reaction towards bicarbonate formation. Establishment of market prices for products and economic factors can determine the installation of a product enriched in carbonate for one period of time, and the receipt of a product enriched in bicarbonate for another period of time, and so on, with average Na / C, when it is formed the average effective use of ionic compounds designed for absorption / conversion.

As a result of separating two processes into two separate chambers and moving between chambers at the moment of OH depletion, formation and fall of temperature, and weakening of absorption, it can change the design and optimization method of decarbonization. It is obvious to a person skilled in the art that batch, pseudo-continuous, continuous, and so on, variants of this simple two-stage decarbonization process can be designed.

In addition, in order to realize absorption with the least energy consumption, in many preferred embodiments of the invention can be used bubbler reactor columns (packed and non-packed, with horizontal and without horizontal fluid flow, with horizontal or without horizontal fluid flow), which due to their design create a large liquid / gas contact area to facilitate mass transfer. The total advantages of this design are the freedom to use steps with a small height (3 m or less), at which, however, 90% + absorption is achieved with little resistance or overpressure when pumping flows. Therefore, they are designed with a wide horizontal surface to achieve industrial productivity (wide shallow dynamic areas or equivalent reactors), preferably with horizontal movement of flows for the possibility of their use in continuous mode. In many embodiments, gas-liquid contactors of many other configurations may be used if these devices allow the desired gas-liquid contact to be achieved.

In FIG. Figure 4 shows a flooding study of a 5 'column in which the resistance is approximately 0.01 psig, plus 1.52 psig of overpressure to overcome the 5' fluid depth. It is assumed that less than 1% of the energy of the power plant is spent on these losses and other costs of compression, and it is believed that they themselves are minimal and are not calculated in the examples. In FIG. 4 , it is confirmed that extremely low flow resistances result in extremely economical compression in the case of an effect in which the environmental performance of the device is not compromised by the excess energy spent on compressing or moving the gases.

The environmental performance of the embodiments of the present methods and installations is increased by doing the least amount of work to absorb CO ₂ , and one factor that reduces this efficiency is the degree of compression, pumping of liquid and air that is required to carry out that method. For this, two high-efficiency absorbers have been developed (capable of removing 99% CO ₂ from the incoming flue gas stream, which is 60% CO ₂ in N ₂ ) for working with "short steps" in which high CO ₂ absorption rates are achieved.

In preferred embodiments of the invention, a large liquid-gas mass transfer surface area is used (bubble column, packed or without packing, or equivalent fixed or moving fluid reactors) to achieve a high absorption rate at a low height of the liquid absorbent, thereby lowering the resistance required to bring the streams into contact, and therefore this “short stage design” requires the use of wide, short “dynamic areas” or their equivalents in ruboprovode, recesses, reactors, etc., for the effective absorption of large amounts of CO _2.

In the large-capacity industry and all related reference literature, the decarbonization reactions of the present description are usually considered to be limited by mass transfer. In practice, using packed or non-packed columns for absorption over a wide gas-liquid contact surface in the methods of a rising bubble through a liquid, the reaction seems to have small restrictions on mass transfer, or, in other words, using the present method of designing a bubble column for contact liquid-gas, apparently, allows you to skillfully overcome the limiting mass transfer: it was demonstrated that sparging from the zero level of the nozzle through the bubbler with the gas-oil contact distance bone only 30 cm, provides an instantaneous rate at 98% + absorption of CO ₂ (see. Figs. 2B and 2C, discussed in Example 3), and within the industry for significant periods of time, 15-25 minutes, retains the ability to flow 80% + absorption. This is hardly strongly limited by mass transfer, and a practical experiment even with the simple disappearance of the load demonstrates rapid mass transfer during this chemisorption.

Three examples of the design of absorbers with a high degree of absorption of CO ₂ / NaOH are described in detail in examples 1-3. The conclusions made in Examples 1-3 indicate that high absorption rates with short NaOH steps have been proven and demonstrated, which are suitable for the industrial removal of high levels of incoming CO ₂ at low resistance in reactors with sizes manufactured by manufacturers.

So, specific embodiments of the present invention relating to part of the methods and installations decarbonization include one or more of the following features:

(1) the use of short steps to achieve high rates of CO ₂ absorption at the stage of carbonization of the reaction;

(2) the isolation and processing of a carbonated stream in a bicarbonization process by continuous contact with a CO ₂ carrier gas (or other CO ₂ carrier gas with CO ₂ concentrations higher than the partial pressure of the CO ₂ formed from the absorbent stream);

(3) possession of the sequence of a method that can be used by routine control of state and concentration variables to obtain pure bicarbonate, pure carbonate, and all their various mixtures; and

(4) embodiments of the invention can be as effective as a 1: 1 sodium / carbon ratio in absorption; this optimizes the absorbed CO ₂ per kWh (environmental performance option, (∂CO ₂ / ∂E)) consumed in the production of the reagent.

IV. Product Highlighting

As noted above, in specific embodiments, apparatuses and methods of the present disclosure use an isolation method by which carbonate and bicarbonate products are separated from a liquid solution. Isolation of liquid solution products requires an appropriate method. The formation in the liquid phase of sodium bicarbonate (NaHCO ₃ or sodium bicarbonate) and sodium carbonate (Na ₂ CO ₃ or soda ash) in equilibrium with sodium hydroxide (NaOH or caustic soda) occurs in a wide range of temperatures and pressures and provides different end points of equilibrium at different partial pressures of CO ₂ . By controlling the alkali concentration, temperature, pressure, reactor size, liquid depth, and degree of saturation with carbon dioxide, precipitation of carbonate and bicarbonate can be caused. Alternatively, in some preferred embodiments, the carbonate / bicarbonate products may be separated from the water by heat exchange with the incoming flue gases. In addition, due to differences in solubility products for sodium carbonate and sodium bicarbonate, specific non-intuitive processing points can be achieved; for example, one of the features of the equilibrium of sodium carbonates in specific solutions of caustic soda is that the addition of heat leads to the precipitation of a solid; also, under specific conditions, it was shown that high (93% +) purity carbonates spontaneously precipitate from an aqueous solution.

Alternatively, in specific embodiments, the heat for the separation method can be obtained from hydrogen produced by electrolysis or by creatively utilizing the heat of the flue gases contained in the incoming flue gas stream. The crystallization method inherently allows the crystallizing mineral to be purified by a known crystallization purification method.

Exiting liquid streams, depending on the design of the reactor, may include water, NaOH, NaHCO ₃ , Na ₂ CO ₃ , and other dissolved gases at different equilibria. Dissolved trace emission components such as H ₂ SO ₄ , HNO ₃ , and Hg can also be detected. In one embodiment, to isolate / remove effluent liquid streams, for example, to remove / isolate water from carbonates (in this sense, “carbonates” means mixtures of carbonate and bicarbonate, potentially also containing hydroxide), any isolation method applied to any such the mixture would probably include the addition of thermal energy to evaporate the water from the mixture, the water may be boiled to evaporate the water using the reboiler 106 shown in FIG. 6 . Alternatively, by maintaining the partial alkalinity of the solution (e.g., NaOH at a concentration of about 1 mol) and then heating the solution in the separation chamber, relatively pure Na ₂ CO ₃ can be precipitated in the storage vessel, and the remaining NaOH will be recycled back to the reactor 200 . In other embodiments, pure carbonate, pure bicarbonate, and mixtures thereof in equilibrium concentrations and / or in suspension or in concentrated form can then be periodically loaded onto a truck / tanker truck. In other embodiments, implementation, liquid streams can be moved to the evaporation pools / fields, where a liquid such as water can be removed by evaporation.

According to FIG. 6 , using the reducible reactor design, the reactor design shown can extract the energy available in the hydrogen produced by electrolysis or in the form of combustible fuel, boiler gas, or in H ₂ / O ₂ fuel cells. The reactor 200 can be used to obtain steady state operation when NaOH and NaHCO ₃ can be obtained in proportions of approximately 50:50. Hydrogen gas obtained by initial electrolysis can be used to provide heat, NaHCO ₃ can be deposited in the deposition chamber 108 with the remaining NaOH returned to the reactor 200 . The suspension from the separation chamber 108 may be supplied to the water treatment chamber 110 , which may be connected to the separation chamber 108 . Alternatively, if necessary, the suspension may be stored and then fed into the water treatment chamber 110 .

OnFIG. 7 another reactor design is shown according to an embodiment of the present invention. In particular, onFIG. 7 Re-capture of some excess energy used to produce a hydrogen by-product is shown. The use of a highly efficient fuel cell in tandem in place may allow the generation of a constant electric current, which can be used to add and partially provide electric current to the electrolysis. From the mixing chamber300 the mixture, including, but not limited to, the percentage of NaOH, NaHCO₃, and NOx, SOx, and Hg are supplied to the selection chamber308. Highlight Camera308 can separate the mixture into solid and / or liquid phases by applying heat to the mixture. A drying chamber (not shown) of a separation chamber308 can remove liquids from the solid and / or liquid phase during the heat supply process. The resulting diluted form of NaOH is sent to the boiler306,at wherein diluted NaOH is concentrated by evaporation and the concentrate is sent to the mixing chamber300 on the line of the return cycle. Water from the boiler306 can be fed into the electrolysis chamber302in particular in a brine mixer302A. Formed Na₂CO₃/ NaHCO₃ (suspension) from the isolation chamber308 can be filed for industrial use. In one embodiment, the carbonate suspension may be directly or indirectly (e.g., retaining NaHCO₃ for later use in a process such as hard water treatment) filed in a water treatment plant310. Alternatively, NaHCO₃ It can be additionally cleaned, dried, shipped, and sent for use in other industrial applications.

The release of gaseous products raises concerns about how safe the release of either NaOH or the aforementioned components is, that is, the release of alkaline rain from a power plant should be excluded as well as the release of acid rain. However, sodium hydroxide is commonly used as an element of scrubbing in a power plant and is approved for use by the US Environmental Protection Agency (EPA). Transportation of sodium hydroxide to power plants, as well as methods for preventing alkaline release, are well known in the art. For example, a simple and inexpensive reflux condenser can prevent any significant release of NaOH in gas exhausts.

In a carbonate precipitation method according to particular embodiments of the present invention, carbonate equilibrium constantly binds carbon dioxide and absorbs gas upon contact with an almost instantaneous conversion to carbonate ion. The chain of chemical reactions can be limited by mass transfer, therefore, as soon as carbon dioxide is absorbed by the base, subsequent ionic reactions occur at a high speed.

The equilibrium of sodium carbonate has the peculiarity that when the temperature is raised, Na ₂ CO ₃ precipitates on its own, which makes it possible to withdraw it in the form of a suspension, with some partial removal with a suspension of NaOH. In one embodiment, treating this slurry with some of the wet chlorine produced in the chlorine cycle can be used to reduce NaOH to trace NaCl in NaHCO ₃ to levels that either correspond to or less than the sodium carbonate obtained from underground mining " thrones "or deposits. Essentially, the sodium carbonate / sodium hydroxide balance provides carbon with a complete transport chain from gas to liquid and to solid. In other embodiments, the carbonate cycle can advantageously be used as a medium for accumulating a suspension of ash, sodium hydroxide, and various other carbonates and impurities, and transporting the suspension for use as a road base.

V. Electrolysis of brine to obtain a liquid absorbent at low energies

As indicated above, in certain embodiments, the plants and methods of the present description use brine electrolysis to produce sodium hydroxide, which is used as a liquid absorbent in the decarbonization process. Brine electrolysis is an electrochemical process mainly used in the production of concentrated sodium hydroxide (caustic soda) and chlorine gas, and it is usually described in the relevant literature by the following equation:

2NaCl + 2H ₂ O + e- → 2NaOH + H ₂ (g) + Cl ₂ (g)

Electrolysis of brine can be carried out in three main types of standard electrolyzers: diaphragm, with a mercury cathode, and membrane. Each of these types of electrolytic cells produces the same end products from the same starting reagents. They differ from each other mainly in the manner in which the starting reagents and the final products are separated from each other.

In one embodiment, due to a number of factors, a membrane cell may be used. First, environmental concerns about the mercury hazard have reduced the demand for a mercury cathode cell. Secondly, diaphragm electrolyzers can produce relatively dilute sodium hydroxide, which contains significant concentrations of salt and ion chloride, and requires significant re-treatment / isolation to remove significant salt content from sodium hydroxide. Thirdly, improvements in the technology of fluorinated polymers have increased the resource and electrical efficiency of the technology using a membrane electrolyzer, in which it is customary to guarantee a life of more than 5 years in the markets for industrial goods. In addition, in preferred embodiments, the energy efficiency per ton of sodium hydroxide is greater than the efficiency for both a diaphragm and a mercury cathode electrolyzer.

Membrane electrolysis plants are characterized, but not limited to, by the following generalized input and output indicators:

Anode: input 26% NaCl + 2275 kWh / ton Cl ₂ → Cl ₂ (g) + 24% NaOH

Cathode: input H ₂ O + e- → 30-33% NaOH + H ₂ (g)

It should be noted that the energy consumption (for example, 2275 kWh per ton of chlorine) may depend on the individual designs of the cells. As such, costs may vary.

Many preferred embodiments may utilize membrane electrolysers in this function. Membrane electrolyzers have several advantages over other brine electrolysis processes. Firstly, membrane electrolyzers do not contain and do not produce any environmentally sensitive emissions (for example, mercury) and are electrically efficient compared to diaphragm and mercury electrolyzers. They also use concentrated / dilute / reinforced NaCl closed loop, so they can be well suited for use as a “salt closed loop” continuous process plant. Further, NaOH obtained in membrane electrolyzers without additional evaporation / concentration can naturally correspond to the concentration level for use in the decarbonization process (for example, 30-33% NaOH by weight). In addition, the hydrogen produced in membrane electrolyzers is "pure", approximately corresponding to "electronic purity", and relatively pure in NaCl or other impurities. Essentially, hydrogen can be compressed and pumped into the tank as gaseous hydrogen gas of electronic purity, used to produce energy on site, for example, to burn in a mixture with low-grade coal, or to increase the efficiency of the combustion technology. Alternatively, hydrogen can be used as boiler fuel for the separation processes that can be carried out after decarbonization. Membrane electrolyzer technology can also be easily scaled from laboratory to industrial scale by adding small additional plants. In addition, gaseous chlorine produced in the membrane process is less "wet" than chlorine produced by other standard electrolytic processes. Essentially, a one-stage compression cycle may be sufficient to produce chlorine with the purity required in the water treatment.

The foregoing relates to the state of the art described in the literature and applied in practice pertaining to the industrial production of chlorine gas and sodium hydroxide. However, the objectives of the specific embodiments of the present invention differ in some aspects, which leads to different chemical methods used to complete the various endings of specific embodiments of the present invention.

A. Use of low voltage electrolysis methods

In some embodiments, implementation of the present invention, brine electrolysis includes methods of low voltage electrolysis (NVE), thereby improving the thermodynamic efficiency of the method. Specific embodiments of the present invention do not produce sodium hydroxide solely as a final product, but use NaOH, primarily as an intermediate absorbent, to absorb CO ₂ from the flue gas before being released to the environment. Since the industrial production of chlorine and alkali usually involves transporting sodium hydroxide from the manufacturer to the place of use, transporting large quantities of water in sodium hydroxide is uneconomical, therefore caustic is usually concentrated to about 55 wt.% For transportation as a liquid by energy-intensive water removal as a result of evaporation, and in some cases, it is concentrated to an anhydrous solid, usually in granular form. This concentration is usually achieved by operating electrolytic cells at voltages higher than 5 V (called "with significant overvoltage"), at which the above-mentioned concentration of 30-35% NaOH is reached, followed by an evaporation cycle to achieve 55% (or to obtain a completely dried product in anhydrous, granular state, and so on).

In most embodiments of the present invention, NaOH is not transported as a final product, and it is produced at suitable concentrations (as an aqueous absorbent stream) at voltages well below the standard 5 V + used in the operation of chlorine and alkali plants. The impact of the use of NVE on the thermodynamic efficiency of the present invention cannot be overestimated, since the energy consumed is accurately determined by a simple equation for direct electric current:

P = V (voltage) × I (current) × (current output)

Since the current is set by the electrochemical process (one pair of electrons per molecule, and so on), the power supply is almost completely regulated by the voltage necessary for electrolysis (an additional important factor is the current efficiency, which is also a function of the applied voltage). Since embodiments of the present invention use NVE with voltages that are easily achieved in standard electrolyzers, by changing the operating conditions to the lowest level of 2.1 V, by modifying the process and dimensions of a standard membrane electrolyzer, less power is consumed in the embodiments of the present invention (kWh) ) for each amount of formed NaOH, than in the case of traditional high-voltage electrolysis in order to obtain chlorine and alkali. For these reasons, preferred embodiments of the present invention include electrolyzers that are designed to use all available methods for carrying out low voltage operation, including but not limited to: operation at a small distance between electrodes, at elevated pressure, at a large membrane surface area, at varying anolyte concentrations and catholyte and protonation-ion relationships. Additional aspects of NEE are explained in detail in Example 4.

B. The "Point of Indifference" of Using Hydrogen

A typical process for the production of chlorine and alkali produces hydrogen gas, which is either burned in place as boiler fuel for evaporation (see above), or in many cases, its evolution is completely suppressed by using an air electrode; for example, a device that delivers oxygen-containing air to the cathode, causing an immediate reduction reaction to occur:

H ₂ (g) + 1 / 2O ₂ (g) → H ₂ O (w / g),

which changes the electrochemical summation of energies and lowers the voltage required for the production of caustic and chlorine (in some cases below 2.8 in various industrial literature) due to the fact that hydrogen is not produced. An advantage for preferred embodiments of the present invention is that they do not require hydrogen as a boiler fuel for evaporation.

Some embodiments of the present invention may include an “air electrode” process that suppresses hydrogen production (and nullifies the ability to extract energy from that hydrogen), but many preferred embodiments utilize hydrogen production for either energy recovery or for use as a source chemical raw materials for other processes.

Therefore, plants using specific embodiments of the present invention have a “point of indifference” regarding the use of hydrogen thus obtained; that is, the plant produces hydrogen to return its energy (for example, 60% of its energy content in the case of the example), and this amount of energy has a specific economic value, and instead of using hydrogen to return energy, hydrogen can be exchanged or sold in an amount that greater than or equal to this economic value of only one energy. Thus, plants of this type represent a kind of "source of hydrogen." Hydrogen is produced and consumed as a source for generating electric current, and anyone who wants to use hydrogen for other purposes should only reimburse its value as a source of electricity or its economic equivalent to maintain the economy of the process. This embodiment has beneficial consequences for future hydrogen economics schemes: having an affordable source of hydrogen available at a point of indifference to price that is lower than the natural energy content in it, a useful and beneficial characteristic of this technology can be proved.

C. Use of optimization methods for protonated brine

Specific embodiments of the present invention specifically optimize hydrogen production instead of suppressing it by the following means of optimizing protonated brine. The common chlorine and alkali industry in some cases uses the addition of HCl in the form of hydrochloric acid to the brine (NaCl) of an electrolyzer called anolyte space. This "protonation" of the brine has the effect of lowering the voltage required to produce a normalized amount of sodium hydroxide, in that it reduces the electric energy required to produce the reagent / absorbent of specific embodiments of the present invention. During normal operation of chlorine and alkali plants, the reason for this voltage-lowering effect is attributed to the “neutralization of NaOH that returns through the membrane,” a process that has been studied depending on the selectivity of the membrane, and so on, and is well understood. Generally speaking, it is believed that this "neutralization" produces only salt water [NaOH (returning) + HCl (anolyte) → NaCl + H ₂ O] and it is not proven that it changes the stoichiometry of the products. However, one point that has not been well studied or applied in the technique is that the addition of HCl does fundamentally change the stoichiometry of the equation as follows:

NaCl + aHCl + H ₂ O → NaOH + (1/2 + a / 2) H ₂ + (1/2 + a / 2) Cl ₂

This additional amount of hydrogen and chlorine produced is of particular interest to the present invention, since embodiments of the present invention rely on energy recovery from hydrogen (which increases when more hydrogen is produced, and is more economical when more chlorine gas is produced per mole of NaOH absorbent produced ) The fact that it produces additional hydrogen to extract energy and chlorine for sale, while the total electrical voltage (hence electric power and cost) is reduced, leads specific embodiments of the present invention to optimize "a" to achieve the lowest energy consumption to produce reagents and the highest energy recovery from products. The optimum usually lies between 0.5 and 1.0 M HCl in a solution of NaCl (aq.). However, optimization is specific to each cell design, size, concentration, temperature, and pressure mode. However, for each cell, there is an optimal degree of protonation (“a”) at which operation with the least energy is achieved. In practice, extremely high “a” values (> 0.15M at 90 ° C, for example) will likely instantly cause blistering on most commercial membranes. Of course, although HCl is the preferred acid for protonating the brine in the invention, many other acids known to those skilled in the art can also be used for protonating the brine.

D. Self-formation of HCl, which is used for protonation

Since specific embodiments of the present invention both use HCl inlet and production of gaseous H ₂ and Cl ₂ , protonation of the brine can be accomplished by self-exchange; that is, the gaseous products of H ₂ and Cl ₂ can be burned (in efficient fuel cells or in conventional kilns) to produce gaseous HCl, which can then be recycled to the anolyte to protonate the brine. The energy recovered from burning H ₂ / C1 ₂ is higher than the energy recovered from burning H ₂ / O ₂ . This increases the thermodynamic efficiency of the invention.

In some preferred embodiments, the absorption / conversion / regeneration pathway is used as the “initial concentration / absorption process”, which is then used to absorb and concentrate CO ₂ for use in geological, oceanic, or terrestrial sequestration methods (eg, methods that inject CO ₂ into carbon storage) using the following methods or similarly:

(1) All hydrogen is burned to produce gaseous HCl.

(2) All of the gaseous HCl thus obtained is reacted with sodium bicarbonate.

(3) As a result of the corresponding neutralization reaction, almost 100% of pure CO ₂ is released and salt water is regenerated, which can be recycled for further absorption cycles.

(4) As a result of this process, the invention is used to absorb, convert, and then release CO ₂ , with the resulting effect of removing gas from the flue gas stream, and concentrating it for further sequestration methods after this process.

E. Adding carbonates and bicarbonates back to the catholyte medium

Unlike chlorine / alkali fuel cells used in the chlorine and alkali industry, in some embodiments of the present invention, a mixture of carbonate and bicarbonate is also recycled to the catholyte (sodium hydroxide) space of the electrolyzer. Such methods have a wide range of changes, but each operating mode for the complete process has an optimal value for recycle of bicarbonate / carbonate mixtures to catholyte, since this can, at certain concentrations and conditions, reduce the voltage on the electrolyzer, and, therefore, the electricity consumption by the electrolyzer.

F. Use of waste heat recovery to heat liquids

Since specific embodiments of the present invention are used in the presence of a power plant or a large CO ₂ emission in the form of flue gas or other hot gases from combustion, there are ample opportunities for heat transfer, in contrast to standard chlor-alkali processes. For example, the temperature of a typical flue gas (for example, after treatment by electrostatic dust collection) may well be 300 ° C. Cooling this off-gas to a temperature below 300 ° C while heating anolyte and catholyte liquids (which, for NVE, should usually be maintained at temperatures> 90 ° C), allows specific embodiments of the present invention to operate without loss of electricity associated with anolyte heaters and Catholic.

G. Use of waste heat recovery for power supply of technological equipment

Typically, since the flue gas that is available in power plants exits at temperatures between 100 ° C (typical for scrubbing), 300 ° C (after the dust collection process), and 900 ° C (entering the dust collection system), or other of such values, as a result of processing, a significant amount of waste heat can be extracted by cooling the incoming flue gas by heat exchange using an electricity extraction cycle, an example of which is the ammonia-water cycle (patented Kalina process, for example), steam cells, or any such cycle that performs a similar thermodynamic problem. Since specific embodiments of the present invention are based on direct current electricity for the production of the reagent / absorbent for the present invention, the process can be directly supplied with electricity, in part or in full, by utilizing waste heat that occurs without the normal losses in the transformer associated with converting this DC power to AC power for other uses. In addition, by using waste heat to operate the engines, significant efficiencies can be achieved without the use of an electricity generation step at all. In some conditions, these amounts of energy can be found by utilizing heat to fully supply the embodiments of the present invention.

VI. The formation and use of by-products from decarbonization and electrolysis processes

As noted above, specific embodiments of the plants and methods of the present disclosure provide a number of useful by-products from decarbonization and electrolysis processes, including chlorine gas, sodium bicarbonate, and hydrogen gas. In some embodiments, hydrogen gas produced by the embodiments of the present invention is introduced into a hydrogen energy reuse cycle. In one embodiment, the present invention may include state-of-the-art techniques for using fluorinated polymer membranes to reduce the migration of chlorine ion for the decarbonization process. Consequently, the process can function without a large expenditure of energy and costs for the separation of the chlorine ion; decarbonization and isolation cycles do not contain chlorides. Essentially, an embodiment of the decarbonization reactor of the present invention can use sodium chloride, water, and spent carbon dioxide combined with electricity to form gaseous chlorine, sodium bicarbonate, and utilize energy by burning hydrogen as follows:

2NaCl + H ₂ O + 2CO ₂ + e- → 2NaHCO ₃ + Cl ₂ + [1 / 2H ₂ + atm O ₂ → e-]

A. Hydrogen Energy Recycling Cycle

Two methods have been developed that can utilize the hydrogen energy extracted by the embodiments of the present invention. The first is the co-combustion of hydrogen with coal to improve emissions of coal furnaces, and the second method involves obtaining a constant electric current in a hydrogen / oxygen fuel cell. About 22% to 32% of the energy spent on electrolysis can be recovered through the combustion of H ₂ / O ₂ . Alternatively, H ₂ and Cl ₂ can be burned in Cl ₂ and HCl directly, either by using a fuel cell, or by extracting direct current. Heat sources from the extraction of waste heat, or from the electrolysis reaction, which produces sodium hydroxide at a temperature of approximately 135 ° C, or from the decarbonization process itself, in which various heat of dissolution, heat of evaporation, and heat released in the exothermic reaction are absorbed, can be disposed of using well-known methods (that is, reboilers, and so on) at the locations of the power plant for heating gaseous products of combustion or other uses .

In one embodiment, the industrial production of direct current in a fuel cell may be advantageous because of its ease of use and safe operation at below atmospheric pressures. The immediate consumption of the hydrogen produced can also directly reduce the cost of electricity for brine electrolysis. In addition, since the hydrogen recovery cycle can be carried out during the off-peak electrical production of H ₂ , Cl ₂ , and NaOH, and H ₂ can then be used to provide peak loads with electricity, the present description provides reagents at low cost with subsequent production of auxiliary high costs with a peak load of electricity and the simultaneous implementation of the decarbonization process. The cost-effectiveness of the H ₂ energy recovery cycle to increase peak plant energy production by increasing electric current production by burning H ₂ as either fuel or in a fuel cell can provide for self-consumption utilization.

Alternatively, this refined carbon-based hydrogen gas source of electronic purity can be successfully used by municipal utilities, industrial companies, and power plants in municipal bus fleets, trains, and other public or private hydrogen fuel uses.

B. Other uses of decarbonization process by-products

In one embodiment, chlorine gas may be the primary oxidizing agent used to kill bacteria in water treatment plants around the world. Chlorine, and some 100+ derivatives of the original chemical raw materials derived from it, are often referred to as generating up to 30% + US GDP. It can also be used in the production of the largest industrial chemical product, hydrochloric acid. Chlorine is also widely used in the plastics industry as the most common non-carbon reagent.

Sodium carbonate is a by-product of the method described in the present invention, which is widely used in the manufacture of soaps, detergents, and shampoos around the world, and also as an additive in production. In addition, power systems, as well as private homes, use sodium bicarbonate to soften hard water as a result of the following reaction:

CaCl ₂ (water) + NaHCO ₃ (water) → CaCO ₃ (precipitate) + NaCl (water) + HCl (water)

A similar method can be used using sodium carbonates and bicarbonates from this method to carry out the exchange of ions with numerous salts of group 1 and group 2 to precipitate various carbonates.

Another example of by-products obtained from the decarbonization process is hydrogen gas. Hydrogen gas of electronic purity is the highest form of energy carrier. On the other hand, the hydrogen fuel produced can be burned with more “dirty” coals to reduce emissions from this fuel, it can be burned as a boiler feed during the separation process, or it can be used in fleets of hydrogen fuel.

In further embodiments, it is possible to effect the transport of carbon dioxide to remote locations of the sequestration in the following manner, or its equivalent:

(1) At the site of the power plant, CO ₂ and other pollutants are absorbed using the method or any variant thereof, along with the production of hydrogen, chlorine, carbonates.

(2) At the site of the removed sequestration, hydrochloric acid and carbonates are again combined in a neutralization reaction to generate substantially pure CO ₂ . The CO ₂ obtained by these methods can then be pumped into the carbon storage.

Using these methods, a similar net effect of transporting CO ₂ between the site of the power plant and the site of the sequestration is achieved without physical transport of CO ₂ as a result of liquefaction and transport by pipeline, road transport, and so on.

EXAMPLES

The following examples are included to demonstrate preferred embodiments of the invention. These embodiments provide a practicable economic solution to reduce or even practically remove carbon dioxide and other contaminants from power plants by obtaining and using the resulting reagents for sale or in industrial processes and / or by returning energy.

For specialists in this field it should be obvious that the methods described in the examples, which follow the methods already presented, disclosed by the author of the invention for proper functioning in practical application, can be considered the preferred modes for its use. However, it should be apparent to those skilled in the art that many changes can be made in the light of the present description in the specific embodiments described, and similar or similar results can additionally be obtained without departing from the spirit and scope of the invention.

Example 1

Design of a bubble column reactor CO ₂ / NaOH in graphical ways

In a bubble column reactor designed for this example, there are four main threads, namely:

(1) The fluid flowing in the bubble column stream at a given volumetric flow rate (Vl = fluid volume over time); in the selected case, the volumetric flow rate at the inlet is equal to the volumetric flow rate at the outlet. Therefore, both are designated as Vl). In this example, Vl = 0.001 m ³ / s.

(2) Vg0 = volumetric flow rate of the incoming gas, which will be partially or fully absorbed by the liquid absorbent. In this example, Vg0 = 0.05 m ³ / s.

(3) Vg = volumetric flow rate of the exhaust gas. In this example, Vg = 0.02 m ³ / s.

The bubble column reactor was designed based on the above conditions. Sixty percent of the incoming CO ₂ in the flue gas was to be removed by sparging through a solution of concentrated sodium hydroxide. The reaction is limited by mass transfer. The aim of the example was to calculate the dimensions of the reactor (height and diameter) needed to remove 99.9% of carbon dioxide. P = 2 atm, T = 298 ° K. Using the graphic data shown in FIG. 8 , this example describes the design of both a high reactor (2.36 m) and a low reactor (0.41 m) . In FIG. Figure 8 is a graph showing the percentage of CO ₂ absorption in a bubble column as a function of liquid depth and interfacial gas velocity at low interfacial velocities (Schumpe et al., 1979).

Column solution with a height of 2.36 m

The conversion reaches 100% at a reduced speed (Ug0) of approximately 0.04 s / s. This speed is in the area of bubbling (in solutions similar to water, it is approximately 0.05 m / s). Knowing the required volumetric gas flow rate (Vg0), the column diameter was calculated:

Ac = Vg0 / Ug0 = 0.05 / 0.04 = 1.25 m ²

Dc = 2√Ac / π = 2√1.25 / π = 1.26 m ²

Therefore, the conversion of 99% of the incoming 60% CO ₂ requires a column 2.38 m high, 1.25 m ^{2 in area} , 1.26 m in diameter, and a total volume of 2.95 m ³ .

0.41m column solution

Conversion in a 0.41 m high column requires a reduced gas velocity of about 0.02 m / s. Like the above calculations:

Ac = Vg0 / Ug0 = 0.05 / 0.02 = 2.50 m ²

Dc = 2√Ac / π = 2√2.5 / π = 1.78 m ²

Therefore, the conversion of 99% of the incoming 60% CO ₂ requires a column with a height of 0.41 m, an area of 2.50 m ² , having a diameter of 1.78 m, and a total volume of 1.03 m ³ .

So, through this example, it has been demonstrated that shorter columns are more efficient per unit volume when removing carbon dioxide from the exhaust gas; for this example, 3 times. Therefore, for preferred embodiments of the present invention, the design aims at short stages and / or multi-stage reactors consisting of short stages.

Example 2

CO bubble column design ₂ / NaOH by determining the mass transfer coefficient

The purpose of this example was to determine the mass transfer coefficient, kla (moles / sec / volume), from a theoretical construction. From this example, it was determined that this correlation method could potentially lead to inconclusive results; that is, this example emphasized the difficulty of predicting real results from theory due to the uncertainty of measuring some basic parameters. Therefore, only experimental scaling can finally determine the result for a large decarbonization plant.

The following equations for gas capture ( εg ) and molar transfer ( kLa ) follow from the correlations established by Akita and Yoshida (1973), and they are applicable to carbon dioxide and water systems for columns with relatively large heights and diameters (i.e.> 0.1 m):

Gas capture

εg

----- = C * [g * Dc ² * ρL / σ] 1/8 * [g * Dc ² / νl1] 1/12 * [µg / √g * Dc]

(1-εg) ⁴

and

mass transfer coefficient

kLa (1 / s) = [Cco2 * Dco2-h20 / Dc2] * [νL / Dco2-h20] 0.02 * [g * Dc2 * ρL / σ] 0.62 * [g * Dc3 / νl2] * ε11

Where εg is the gas capture coefficient

Cco2 = concentration of CO ₂ in the exhaust gas

Dc = column diameter

νL = 0.0001 m ² / s

σ = 1 cP = 0.1 Pa · s

ρL = 998 kg / m ³

and since Dco2-h20P = 1,661 m ² · Pa / s,

then, Dco2-HO = 1,661 / 5,067 · 10 ⁵ = 3,278 · 10 ^-6 m ² / s

The driving force is the difference between the equilibrium concentration of carbon dioxide (Cco2 *) and the current concentration of carbon dioxide in the liquid phase (Cco2), which in this example is assumed to be zero; that is, the sodium hydroxide present instantly neutralizes aqueous carbon dioxide, the “acid." The molar transfer rate per unit volume of the reactor can, therefore, be written as:

Nco2 = kLa * [ Cco2 * -Cco2] = kLa (Cco2 *)

The molar transfer rate required to remove 99.99% of carbon dioxide in the exhaust gas. It is believed that CO ₂ is an ideal gas under column conditions.

Cubic volume / sec:

Vg0 = 0.05 m ³ / s :. Vco2 = 0.6 * Vg0 = 0.03 m ³ / s

Mol / sec:

Vco2 * P / RT = [0.03 M 3 / s * 5 atm] /

[0.082m3 atm / kmol K * 298] * 1000 gmol / 1 kmol = 6.14 mol / s

Nco2 (mol / s) = 0.999 (mol CO2 removed / mol CO2 received) * 6.14 mol / s = 6.13 mol / s

Other fluid properties required for the correlation model.

Having established the initial reduced velocity at the limit of the bubbling flow (Vg0 = 0.05 m / s), the column area and diameter were calculated:

Ac = Vg0 / Ug0 = 0.05 / 0.05 = 1.0 m ²

Dc = 2√Ac / π = 2√1.0 / π = 1.13 m ²

For gas-phase capture reactions, this example sets C = 0.2 and it is assumed that the reduced velocity (Ug) is the average between incoming and outgoing speeds; Ug = 0.035 m / s = arithmetic mean of (0.05 m / s and 0.02 m / s), and using the program to solve the equation, it was found that εg = 0.055.

Next, the molar transfer rate constant was calculated:

$$\begin{array}{l}kLa(\mathrm{one}/\mathrm{from}e\mathrm{to})=\frac{0.6\cdot \mathrm{3,278}\cdot 10-6{\mathrm{<figure-callout\; id="2"\; label="m"\; filenames="00000005.png,00000010.png"\; state="\{\{state\}\}">m</figure-callout>}}^2/\mathrm{from}e\mathrm{to}}{{[0.6m]}^2}\cdot \\ \frac{0.0001{\mathrm{<figure-callout\; id="2"\; label="m"\; filenames="00000005.png,00000010.png"\; state="\{\{state\}\}">m</figure-callout>}}^2/\mathrm{from}e\mathrm{to}}{\mathrm{3,278}\cdot 10-6{\mathrm{<figure-callout\; id="2"\; label="m"\; filenames="00000005.png,00000010.png"\; state="\{\{state\}\}">m</figure-callout>}}^2/\mathrm{from}e\mathrm{to}}\cdot \left[\frac{9.8m/\mathrm{from}e{\mathrm{to}}^2\cdot {(0.6m)}^2\cdot 998\mathrm{to}g/{\mathrm{<figure-callout\; id="3"\; label="m"\; filenames="00000005.png,00000010.png"\; state="\{\{state\}\}">m</figure-callout>}}^3}{\mathrm{0,0696}\mathrm{to}g/\mathrm{from}e{\mathrm{to}}^2}\right]\\ \cdot \left[\frac{(9.8m/\mathrm{from}e{\mathrm{to}}^2)\cdot {(0.6m)}^3}{{[0.0001{\mathrm{<figure-callout\; id="2"\; label="m"\; filenames="00000005.png,00000010.png"\; state="\{\{state\}\}">m</figure-callout>}}^2/\mathrm{from}e\mathrm{to}]}^2}\right]\cdot (\mathrm{0,055})\mathrm{1,1}=0.386/\mathrm{from}e\mathrm{to}\end{array}$$

kLa = 0.386 / s

Returning to the equation for the driving force, the reactor volume (V) was calculated:

V = Nco2 / [(kLa) (Cco2 *)] = 6.13 mol / s / (0.386 / sec * (103.6 mol / m ³ )) = 0.15 m ³

and

Hc = 0,15 m ^3/2 1 m = 0.15 m

Thus, the dimensions of the bubble column are Dc = 1.13 m and Hc = 0.15 m, leading to a significant difference from the actual results in real bubble columns.

A few assumptions can explain the differences between correlation models (which are considered the best models for this mode of operation) and real results:

(1) It was assumed that sodium hydroxide has the properties of water (density, surface tension, and so on); and

(2) The concentration of CO ₂ in solution cannot be well characterized as zero; it is a more probable quantity of calculations; for example, if the effective concentration of CO _{2 is} not equal to zero, then the driving force will be less, and a higher column is required.

It should also be noted that this theoretical correlation is violated under the condition, which is also its effectiveness: since terms such as ( νL = 0.0001 m ² / s) are often squared in the denominator, small changes in these numbers lead to gigantic effects. This type of theoretical construction is good for selecting the approximating function after the fact, but is not a good predictor of mass transfer for design purposes. In addition, there are striking hydrodynamic effects on the absorption / conversion rate of CO ₂ , so it should be designed at different and proportionally increasing scales.

Example 3

Design of a bubble column (depth) CO ₂ / NaOH from experimental data

The observation that the reliance on specific embodiments of the present invention on the “short stage efficiency theory” described herein (3 m or less of gas-liquid contact distance, or height of the liquid phase to achieve> 90% absorption) is supported by two different calculation methods that are found in compliance with the design used in chemical technology. However, in specific cases (as noted above), certain simplifying assumptions were made in these design calculations, therefore, an experimental verification was necessary and carried out with the results shown in FIG. 2B and 2C (explained in more detail below).

In each of these technological experiments, a certain absorption of CO _{2 was} achieved at a certain gas-liquid contact distance (that is, the height of the liquid in cases of a non-packed, open bubble column); for example, 20% absorption through 30 cm of liquid.

If the gas was then passed through a second column of a similar design and under the same conditions, then the same absorption would occur again; that is, 20% of the remaining 80% of the original CO ₂ would be absorbed again. This dependence ultimately subsides; however, given the high absorption characteristics of the liquid absorbent, and the tendency for chemisorption to occur with stable absorption of even diluted CO ₂ in the flue gas, this effect is ignored for this example and 90% CO _{2 is} removed in the design.

Obviously, additional flow passages through the absorption liquid would again reduce the remaining CO ₂ by 20%, and so on, until a sufficient number of flow passages was achieved to achieve the desired absorption level (in this case, 90%).

This leads to the design of a "series of main steps", in which the depth of the liquid (the depths of many steps with a height of 30 cm) is determined, which is necessary to achieve 90%. Assuming that each stage absorbs at the same% CO ₂ / distance as the previous one, the results shown in FIG. 2B and 2C, and graphically shown in FIG. 2D and 2E .

In FIG. 2A shows an apparatus for examining the main features of one embodiment of the decarbonization step of the present invention. The installation of FIG. 2A (or a unit with a similar action) can be operated according to the procedure below:

(1) NaOH (for example, 1M NaOH in water at 25 ° C.) is loaded into carbonization unit 801, with or without a nozzle, to a depth of 30 cm with the test load.

(2) The flue gas, simulated or real, in this case, a typical flue gas exhaust from coal heating (16% CO ₂ , 84% N ₂ , SOx / NOx <1% or at natural ppm levels) is introduced into the carbonization unit 801 , effectively irrigate in a column without a nozzle, effectively irrigate or distribute in packed columns, move through a liquid, and remove. The gas is at a temperature of 25 ° C, CO ₂ with a flow rate of 2 l / min is mixed proportionally with other gases, rises up the column with a circle diameter of 4 "; the pressure in the system can be 1 atm psig or lower.

(3) It can be noted by measuring the concentration of incoming CO ₂ (by gas chromatography of a sample taken, for example, or by measuring the concentration of CO ₂ in the stream) that CO _{2 is} absorbed by a liquid, that temperatures increase (an exothermic reaction), and analysis is taken by a liquid the samples will show the presence of carbonate / bicarbonate / hydroxide equilibria, indicating that not only the absorption of CO ₂ occurs, but also its conversion to the carbonate or bicarbonate form. Field experience indicates that these key “transition points” are in equilibrium with pH:

a. At pH <= 8.3, bicarbonate is predominantly formed.

b. At pH> = 10, carbonate is predominantly formed.

(4) The reaction of absorption / conversion to carbonate continues actively and exothermically until, at a given gas flow dynamics, at whatever speed CO ₂ is absorbed / converted, the exothermic phase of the reaction ends, the temperature first does not change and then drops, and the absorption capacity of a liquid, which decreases as the OH ion concentration decreases, quickly reaches zero at this point. When absorption levels begin to fall, usually the pH value approaches 8.3 or close; at pH> 8.3, absorption is relatively active.

(5) The liquid is transferred to the bicarbonization column 803 and the exhaust gas is introduced back into the liquid. The absorption of CO ₂ ceases, and in some cases, as will be shown, becomes negative (the liquid gives up a certain amount of CO _{2 to the} gas stream passing through it). The temperature of the liquid continues to fall, partly due to some inconsequential evaporation into the moving gas stream, but also due to the bicarbonization reaction that occurs between the previously formed sodium carbonate and the remaining "free" CO ₂ , which is dissolved in the liquid.

(6) The equilibrium continues to shift towards bicarbonate, and optimization of the initial hydroxide concentration, liquid and gas temperatures, pressures, flow rates and speeds, steepness, and so on can be carried out, up to the point of obtaining pure bicarbonate (99% +).

In FIG. 2B and 2C show the results of several series of load-load tests (a specific concentration of NaOH was created in the decarbonization system, as shown in Fig. 2A ). Several key points are illustrated by the data in FIG. 2B and 2C :

(1) The conditions can be accordingly applied for reproducible production of either pure carbonate (experiments 4 and 14) or pure bicarbonate (experiments 28 and 32), and can be adjusted to achieve different results (or "ionic ratios" between the extreme values of 1, 0-2.0).

(2) It was found that the dimensions of the reactor, which are the result of this study, for all cases with significant absorption, providing distances for gas-liquid contact, usually less than 3 m, are sufficient to achieve 90% absorption of the incoming gases. Therefore, it has been shown that in order to achieve absorption rates corresponding to the limits of thermodynamic efficiency, short stages with low resistance should be designed. In other words, the physical CO ₂ removal method works at absorption levels that can obviously correspond to or exceed the thermodynamic efficiency of the system. Such high absorption rates (inlet gas, outlet gas) are not attributed to energy and, therefore, due to the production of CO ₂ . Therefore, maintaining the rates of CO ₂ absorption from the stream and the thermodynamic efficiency of the plant as two clearly distinguishable values is important to avoid confusion.

The results shown in FIG. 2B and 2C (absorption of CO _{2 by} liquid) and the ionic ratio in the product (1.0 = bicarbonate, 2.0 = carbonate) are shown in FIG. 2D and 2E . Several important conclusions can be made based on FIG. 2B and 2C :

(1) Instantaneous absorption rates of about 98% are noted for single-stage absorption of incoming CO ₂ .

a. Pure bicarbonate (NaHCO ₃ ) was obtained in solution under conditions in which 25% of the incoming CO _{2 was} absorbed in a single-stage bubble column at a gas-liquid contact with a depth of 0.30 m liquid depth / gas-liquid contact distance. By extrapolating to 90% absorption, we find that 3 meters of contact distance is sufficient to absorb 90% of the incoming CO ₂ .

b. Pure carbonate (Na ₂ CO ₃ ) was obtained in solution under conditions in which 70% of the incoming CO _{2 was} absorbed in a single-stage bubble column at a gas-liquid contact with a depth of 0.30 m liquid depth / gas-liquid contact distance. By extrapolating to 90% absorption, we obtain that <2 meters of contact distance is sufficient to absorb 90% of the incoming CO ₂ .

c. Different absorption depending on the ratio of carbonate ion in the products indicates that between these extreme points there is a continuous series of solutions.

(2) The liquid absorbent retains its absorption characteristics for industry acceptable periods of time (e.g., 15-240 minutes in these examples).

(3) The input variables of the reactor (concentration, temperature, pressure, gas flow rate, contact time, and so on) can be adjusted to obtain pure bicarbonate, pure carbonate, or any mixture thereof.

(4) Using these laboratory results to design a reactor with 90% CO ₂ absorption results in a value of less than 3 m for gas-liquid contact distance (e.g., approximately liquid depth, column height), and 1 m for acceptable industrial processes in many areas.

Example 4

NVE analysis for various chemical conditions

In FIG. 5 is a diagram showing the performance characteristics of low voltage electrolysis for various chemical conditions. It shows some typical results in which a membrane electrolyzer for producing chlorine and alkali is operated under non-standard conditions, namely:

(1) the pH of the anolyte (protonated brine) is adjusted using a closed pH cycle, adjusted by adding HCl (hydrochloric acid in water), at pH 1.0, 2.5, and 5.0;

(2) the anolyte temperature is maintained at a given value using a closed cycle of fluid flows heated by electric heaters; and

(3) the voltage is regulated for each liquid / protonation / temperature condition by recording the electric current reached by 0.01 m ^{2 by the} electrolyzer to produce chlorine and alkali.

In FIG. 5, the set of experimental data of this example is noteworthy, which are graphs of the actual experimental voltage versus current (expressed in current densities, kiloampere / m ² indicated on the graph) for 0.01 m ² electrolyzer, 13 mm distance between electrodes, operated under various combinations of temperature and degree of protonation of anolyte liquid brine (controlled in this series of experiments using a closed pH cycle controlled by adding HCl (l) to the brine cycle).

Noteworthy are the following points regarding these typical results in FIG. 5 :

(1) At high voltages (5 V), which are usually used in such electrolysis cells to produce chlorine and alkali, the maximum current is reached (and for this electrolyzer, therefore, the maximum current density).

(2) At a higher brine temperature at the same pH, a higher current density at a given voltage is established.

(3) At a lower pH of the brine, a higher current density at a given voltage is established compared to a higher pH of the brine.

(4) These general trends (higher temperature, higher acid concentration) can be optimized using standard experimental design techniques to design the sizes / components of each individual cell to obtain the optimum (kA / m ² · v) for this cell. A similar experiment on any electrolytic cell to produce chlorine and alkali at an increased working pressure leads to the conclusion that an increased working pressure also increases (kA / m ² · in).

(5) The slope of these lines (ΔV / ΔA) is initially large, with relatively large voltage drops occurring with relatively low current / current density drops; however, after the inflection point of the curve is reached (approximately (2.5 V, 10 A / 0.01 m ² ), subsequent voltage drops result in sharper drops in current, and therefore current density.

(6) This inflection point and its surroundings for operating conditions is an optimal compromise solution for the ratio of voltage and current to current density for efficiency from an economic point of view. Optimization with the standard method of experimental design allows you to achieve the optimal low voltage conditions for any real cell of an embodiment of the invention.

(7) In the context of this example only, conditions for anolyte, including pH 1.0, 90 ° C, have the best current-voltage characteristics and, therefore, the optimal working line presented among these various driven lines.

(8) The main disadvantage of electrolysis at low voltage is the concomitant decrease in current density; kA / m ² decreases with decreasing voltage. Since the system can produce the same number of Na ⁺ ions to absorb the same amount of carbon dioxide, m ² the surface area of the membranes should proportionally increase; for example, if the current density drops by 50%, then a membrane with a surface area twice as large is required to produce a sufficient amount of liquid absorbent. This greatly affects the installation cost, since the installation cost of producing chlorine and alkali is almost proportional to the surface area of the membrane. Low voltage electrolysis provides a number of advantages that, in specific embodiments of the invention, can optimize low voltage conditions that compensate significantly or completely for this drawback of the need for a large surface area of the membranes. Namely, the service life of membranes and electrolytic cell components that are operated under milder / less energy-intensive conditions that can extend the life of the cell and / or membrane can be tested in practice. Designing specifically for low voltage conditions may be accompanied by some ability to impose less stringent requirements on certain materials and performance that are not so important in embodiments in which low voltage is used. Some of these degrees of freedom in design can lead to low cost electrolyzers that partially or fully compensate for the increase in the cost of electrolytic cell membranes, which was originally due to low voltage / low current density operation. In these and many other ways, NVE systems, although they require a larger membrane surface area than standard electrolytic cells with a membrane to produce chlorine and alkali to produce the same amount of NaOH, can completely or partially reduce some of these additional costs and operating costs.

(9) The trade-off between the advantage (low voltage and therefore lower power consumption) and the disadvantage (higher surface area and lower current density) can be optimized using the procedure described in Example 7. For a pH of 1.0 / 90 ° C shown in FIG. 5 (which, for this small set of example conditions, is the best V / I characteristic for the operation of NVE), Voptlve can be calculated, and Ioptlve can be obtained from the above dependence. Therefore, for the given dimensions of the design of the electrolyzer, all conditions can be fulfilled in terms of temperature, pressure, brine concentration, degree of protonation, membrane samples, and so on, to obtain the best V / I curve or working line, and then the optimal point on this curve can be calculated by the method of Example 7. In this case, Voptlve is 2.88 V and the current density Ioptlve is 1.04 kA / m ² .

(10) In Example 7, the current at Vopt = 2.88 V is approximately 5 A in cases of slight protonation and / or low temperature. In the same example, this current (and therefore the current density) was more than twice as large, i.e. 10.4 A.

(11) Additional protonation of the brine, temperature, pressure, concentration, geometric arrangement of the electrolytic cell components, electric fields, and conditions can be similarly optimized to obtain the best value (kA / m ² · v), but protonation per se increases the amount of stoichiometric hydrogen, thereby increasing the return on energy of the system. It is important to note that optimization for the lowest CO ₂ absorption / conversion energy can also be achieved by optimizing the indicator (kA / m ² · in) of the system, which lowers the energy required to produce a liquid absorbent, but which simultaneously optimizes the hydrogen available for energy recovery (and then optimizes the efficiency with which the energy of this available hydrogen is extracted), the total energy for the process can be optimized to its lowest value for a given cell of a particular design.

(12) Based on the fact that embodiments of the present invention can effectively absorb CO _{2 with} extremely dilute hydroxides (0.2M or lower was demonstrated) compared to the concentrated hydroxide typically obtained in the production of chlorine and alkali (usually 33-35% by weight , then additionally concentrated by evaporation), the design of electrolyzers for the production of alkali chlorine for operation at low concentration (as well as operation at low voltage) can provide new degrees of freedom for opt imitation of construction under these non-standard conditions.

Embodiments of the present invention are unable to receive more hydrogen energy than the energy used to produce this hydrogen. Otherwise, the result will be a violation of the second law of thermodynamics. This imposes a limitation on the minimum voltage that can be applied to the cell. Assuming 100% efficiency in the return of hydrogen, and using an energy content of 39,000 kWh / ton H ₂ (reference value of the Energy Information Service, EIA, USA), a value of 1.55 V can be obtained for the minimum voltage. A person skilled in the art can, for any system with a given hydrogen / electricity return efficiency and a selected value for the energy content of the system, calculate the minimum attainable voltage for this system.

In practice, thermodynamic inefficiencies (including, but not limited to, I2R losses, inefficient use of current in the cell, waste heat loss, and so on), and the requirement of a small overvoltage for operation, increases the minimum voltage achievable for a given cell. The above figures vary slightly depending on the value of "a," the degree of protonation, although it changes the amount of hydrogen available for energy recovery.

In this case, the current density at low voltage determines the magnitude of the electrolysis areas (strongly affecting capital costs) required to produce the amount of caustic, and at minimum voltage, the required area is extremely large. Therefore, for operation, some voltage is required above the minimum voltage, the value chosen during the design, depending on the compromise, capital costs / environmental efficiency. The current efficiency (the fraction of the current spent on the production of the product) drops at low voltage, so optimizing the operation of electrolysis at low voltage is not the same as achieving the operating condition only at low voltage. The existing processes are designed for operation in the NVE mode (below 5 V), and at these voltages below 5 V, the energy consumption relative to traditional methods increases significantly.

Example 5

Thermodynamics of large-scale plant design

For this example, a model plant (including specific embodiments of the present invention) is described, characterized by a mode for a full-scale plant, and the amount of energy required to extract a given amount of CO _{2 is} quantified, which is limited within statistical limits by various means and methods, including the following:

(1) The thermodynamic efficiency (∂CO ₂ / ∂E) can be approximated as ΔCO ₂ / ΔE for sufficiently short intervals of changes in E (energies).

(2) Certain simplifying assumptions can be made regarding plant design, such as those shown in FIG. 9A , among them:

a. The main expenditure of energy occurs in the process of electrolysis; energy consumption for pumping, compression, regulation, and so on is considered minimal relative to the energy expended in obtaining reagents (electrolysis) and in the recovery of hydrogen energy. These values are taken equal to zero or <0.1% of the energy consumed in electrolysis operations.

b. The energy consumed by electrolysis can be represented by approximately the following equation:

Eout = V * I * EFFcurrent

Where: V = voltage on a working cell

I = current required to produce chemical products by electrochemical semi-reaction, including more than 1: 1 stoichiometry caused by protonation of the brine. The example uses the 0.05 HCl / NaCl ratio of protonating ions consumed in electrolysis.

EFFcurrent = current output, determined by the amount of current used in the actual production of chemical compounds, the remainder being lost with I2R losses, and so on. The value used in the example is 97%; Each cell has its own unique current output, which decreases and changes over the life of the cell.

c. The energy extracted from the combustion of hydrogen (regardless of how, combustion in the form of boiler gas, combustion in a fuel cell, and so on) is expressed as:

Ein = 39000 kWh / ton of compressed H ₂ * ton of H ₂ * EFFdc

Where: tons of H ₂ = tons of hydrogen obtained by the process, including hydrogen obtained by stoichiometry of more than 1: 1, caused by protonation of the brine.

EFFdc = Efficiency of the hydrogen extraction process when the initial energy of hydrogen gas is converted into direct electric current. 60% of the generated hydrogen energy is extracted into direct current. This is a ready-to-use direct current hydrogen / atmospheric oxygen fuel cell; A similar value is obtained for the process of hydrogen compression (the cost of which is 15% of its energy for implementation) and, then, assuming that 85% of the hydrogen energy is returned in hydrogen-consuming processes, such as hydrogenation processes, gasoline reforming processes, and so on.

d. The return energy from the waste heat recovery was extracted from the heat of the incoming flue gas stream. Incoming heated gases are cooled when they enter the process. In some embodiments of the invention, this cooling can be carried out by absorbing the waste heat and converting this heat into direct current, which can be used to supplement / fully provide energy / increase the energy of the process that constitutes the invention. In this example, no additional waste heat recovery is provided.

The plant model for this example includes modeling flue gas exiting a power plant under normal operating conditions, as shown in FIG. 9B . This includes significant assumptions regarding the composition of fuels, the efficiency of the combustion processes themselves, the relative proportion of elements in the combustion processes, and so on. Assumptions for this example are shown in FIG. 9B and they correspond to the flue gas outlet of a typical brown coal-fired power plant with a thermal capacity of 10,000 BTU / kWh.

For a given flue gas outlet, hydroxide is needed, the need for which can be calculated. A few assumptions should be made here. The ion ratio (“ion ratio” is the Na / C ratio in the absorption / conversion reaction) is the same as the ratio of these elements in the resulting solid products. In the case of pure bicarbonate, the value will be 1.0, in the case of pure carbonate, the value will be 2.0, and for mixtures of bicarbonate and carbonate, the value will be between 1.0 and 2.0. The calculation of alkali requirements for this example is shown in FIG. 9C . For the example shown in FIG. 9C , the ionic ratio is 1.0.

For this hydroxide need, there are corresponding electrolysis needs based on the amount of water, salt, square meters of membrane surface area (scalar for electrolyzers of this kind), and current density (in itself is a function of the electrolyzer design, the chemistry of the process, and operating conditions ; here the indicator 3 kA / m ² ) is used. These calculations of electrolysis requirements for this example are shown in FIG. 9D .

For a given amount of electrolysis current under protonation conditions, a given amount of hydrogen gas is generated, which in itself represents the amount of energy available for extraction, or hydrogen is used chemically in other technologies. In this example, hydrogen is converted to direct current using a fuel cell operating at EFFdc = 60%.

For a given amount of flue gas being processed, there is a certain amount of waste heat that can be extracted from it with a certain conversion efficiency into direct current, and this recoverable electricity can be used to supplement the direct current consumed in electrolysis. The waste heat for this example is shown in FIG. 9E , with an efficiency selected in this case equal to 25%, a value that is exceeded by various methods for producing direct current from waste heat currently existing in this field.

The individual components of energy input and output shown here, the resulting effect of these energy transfers can be summarized, as in FIG. 9F . Here, energies are presented in kWh and as a percentage of the power of the main unit, and the calculation of environmental efficiency for this example is shown.

In some embodiments of the present invention, an additional H ₂ / Cl ₂ fuel cell may be used to burn gaseous hydrogen and chlorine to recover the HCl used to protonate the brine. In particular, a “super-stoichiometric” amount of HCl can be recycled, and theoretically, the need for a stock of HCl added to the system can be eliminated. In real practice, a certain amount of HCl recharge should be periodically added to the system. Burning H ₂ in Cl ₂ gives more energy than burning H ₂ / O ₂ . Thus, a small amount of additional energy is obtained. However, fuel cells are not inherently completely efficient in nature, so the “gain” in energy from using chlorine as an oxidizing agent is more than offset by losses inherent in the extraction method. Therefore, in this example, it is believed that these effects are mutually annihilated and are reduced to zero. Therefore, the two effects to some extent cancel each other, but, however, lead to general losses. However, the effect of the presence of additional protons during electrolysis should sharply catalyze the production of NaOH at low voltages and lead to a high value (kA / m ² · v) at these low voltages. Therefore, for any given installation, optimization can be made of processing a given amount of H ₂ / Cl ₂ into HCl and protonating the incoming brine with this amount of HCl. At a certain optimum value (which is usually between a = 0.05 and a = 1.0 M, or near pH 1 at 90 ° C), the hydrogen / chlorine fuel cell losses (which are superior to the small gain provided by chlorine compared with oxidation oxygen,) and the energy gain for the hydroxide (the best value of kA / m ² · in) will be simultaneously optimal for the entire system. It should be noted that in this example, only the burning of H ₂ / O _{2 is} calculated; burning of H ₂ / Cl ₂ has a small thermodynamic gain due to the higher rate of chlorine oxidation, but the counteracting effect of the fuel cell inefficiency contributes to a slightly negative, but considered minimal effect.

A. Calculation of environmental performance

The calculation of environmental efficiency (∂CO ₂ / ∂E) and ΔCO ₂ / ΔE for this example was carried out as follows:

(1) It has been suggested that there are three installations:

a. The main power plant (shown as an example in the model of the exhaust gas in Fig. 9B )

b. Installation absorption / conversion of CO ₂ (which requires supplementary power to process the flue gas from the main power plant and returns a portion of this energy is by the combustion of hydrogen or by the calculated energy of hydrogen if hydrogen is the end-product and is not combusted).

c. The third, additional power plant, which provides the energy necessary for the installation of absorption / conversion of CO ₂ . In this example, it was assumed that the characteristics of this power plant are the same as those of the main plant.

(2) The following aspects related to CO ₂ and energy consumed in the processing of 100% of the main power plant were then calculated:

a. CO ₂ from the main unit (flue gas model)

b. The energy produced by the main unit

c. Net consumption of energy consumed by the process of absorption / conversion of CO ₂

d. The net energy input consumed by the additional power plant is assumed to be the same as the net energy input consumed by the CO ₂ absorption / conversion process.

e. It is assumed that the CO ₂ produced by the additional installation is proportional to the energy produced by the additional installation and with the same ΔCO ₂ / ΔE as for the main installation.

(3) The following results for the above calculations were obtained:

a. The main installation is a plant with a thermal capacity of 10,000, producing 1 GW continuously for one year as a basis, produces 8.76 gigawatt-hours each year and produces 7446068 tons of CO ₂ per year, on average 1176 kWh / ton CO ₂ .

b. Installation absorption / CO ₂ conversion - for this example calculation (a = 0,10, in 2.1 mode, pure bicarbonate produced, 15% of hydrogen energy consumed for compression costs pumping / compression and advantages of the waste heat recovery is not included), 3,648 gigawatt-hours is required for the absorption / conversion of a 100% main unit.

c. Additional power plant - the plant in this example produces the energy consumed by the CO ₂ absorption / conversion plant, 3,648 gigawatt-hours, and itself produces (based on the above value 1176 kWh / ton CO ₂ ) a total of 3101209 tons of CO ₂ , which, supposed to be released into the atmosphere.

d. The total energy generated is therefore 12.48 gigawatt-hours. The total energy supplied is therefore 8.76 gigawatt-hours. The total CO ₂ generated is therefore 10.55 megatons. The total CO ₂ emission is therefore 3,101 megatons. 29.1% of the total energy is consumed in the process of absorption / conversion of CO ₂ . 71.9% of the total CO _{2 is} consumed.

Several key points are illustrated by the above calculations:

(1) Arithmetically demonstrated that the following formulas can be used:

% energy consumed = 1-% CO ₂ consumed

% CO ₂ consumed = 1-% energy consumed

This forms a working line, called a single working line, shown in FIG. 9A.

(2) For this example, (∂CO ₂ / ∂E) and ΔCO2 / ΔE are algebraically identical, namely:

ΔCO ₂ / ΔE = (∂CO ₂ / ∂E) = 0.291 / 0.719 = 0.41

Further, extrapolated cases can be simulated in which the CO ₂ emitted by the additional power plant is itself subjected to the process of absorption / conversion of CO ₂ in another plant # 2 of correspondingly lower power, and this absorption / conversion plant # 2 is respectively supplied with energy with additional energy installation # 2, and so on, which leads to results such as in table 2 for the first five series of iterations.

TABLE 2 Energy
setting Produces
CO ₂
(megatons) Produces
the energy
(kWh) Absorbi-
quantity Necessary
brought
energy Basis 7446068 8760000000 7446068 3648448267 Iteration 2 3101209 1519540497 Iteration 3 1291621 632872705 Iteration 4 537947 263584854 Energy
setting Additional
CO ₂ %
total
E %
total
CO ₂ Environmental
efficiency
(∂CO ₂ / ∂E) Basis 3101209 29% 71% 2.40 Iteration 2 1291621 37% 89% 2.40 Iteration 3 537947 40% 96% 2.40 Iteration 4 224049 41% 98% 2.40

A few points regarding table 2 are important for the model:

(1) It should be noted that the efficiency of the process, either in the case of a basis, or in any cases of successive iterations, constantly gives the same value (∂CO ₂ / ∂E) for the system; this indicator is considered constant for systems that approximate the limitations of this model, and it is called, for these purposes, the environmental efficiency of the process.

(2) It is clear that the value of (∂CO ₂ / ∂E) is constant for all solutions; therefore, a solution can be obtained when the number of iterations is assumed to be infinite; that is, when the installation is operated so that it consumes 100% of the CO ₂ produced by the installation, simply by using the following equation:

1 / (∂CO ₂ / ∂E) =% plant capacity required to absorb / capture 100% of the produced CO ₂

In the case of the example, the calculation leads to a value of 41.6%.

(3) On the other hand, it is obvious that when the usable power spent on absorption / conversion is zero (neglecting waste heat recovery), for this process condition, the absorbed and converted CO _{2 is} also zero. Therefore, all working lines for plants of this type theoretically intersect at a point (0% energy, 0% CO ₂ ).

(4) For data of any two points in a linear system, straight-line solutions for work lines can be built that determine the performance of the CO ₂ absorption / conversion process using the following methods:

a. for each working condition, a basic case is solved, and the obtained solution point of a single case (% energy,% CO ₂ ) is plotted as% CO ₂ (y axis) versus% of usable power (x axis);

b. for this case, calculate (∂CO ₂ / ∂E), and solve the case for x-coordinate at y = 100%; and

c. believe that all lines pass through the origin. In real systems, there may be some energy consumption (regulation, environment, and so on) even with zero absorption, so this case is idealized. In practice, these lines may be slightly curved and do not extend from the origin.

(5) Thus, a family of work lines can be created for this type of CO ₂ absorption / conversion process.

(6) On the same type of graph, graphically competing technologies can also be plotted and compared, for example:

a. A competing technology absorption using MEA (methylethylamine) consumes 30% of plant capacity for absorption of 58% CO ₂ emitted before absorption was applied.

b. In addition, an estimated 15% of the plant’s capacity is spent on liquefying this CO ₂ as a result of applying ultrahigh pressure and a refrigeration cycle (45% power / 58% CO ₂ ).

c. Further, this would give a value for (∂CO ₂ / ∂E) equal to 1.24; however, there is unaccounted for energy required for transporting / injection / CO ₂ content at the carbon sequestration site.

d. It is graphically shown that this competing technology is less efficient than an example of a CO ₂ absorption / conversion unit using a method that is one embodiment of the invention; that is, this model shows that competing technology would require 70% + plant energy to remove 100% of the CO ₂ it produces. Attention should be paid to these points in relation to competing technology, shown graphically in FIG. 9A (as indicated on the diagram):

i. According to 2005 EIA estimates, CO ₂ absorption using MEA technology requires 30% of the plant energy to absorb 58% CO ₂ from the produced flue gas. (The position in the diagram in Fig. 9A (30%, 58%) for absorption-processing only is noteworthy.)

ii. By similar estimates, the compression / liquefaction of this CO ₂ consumes another 15% of the plant's capacity, shifting the operating point of such a plant to a value (45%, 58%).

iii. There is no reliable estimate of either the energy required to transport liquid CO ₂ through a pipeline or other means of transportation, nor the amount of energy needed to pump or pump this CO ₂ into a carbon storage of various nature, nor the amount of energy that may be required for eternal storage of CO ₂ in specified storages. However, despite the fact that these additional energies are not measurable, it seems reasonable to assume that they are not equal to zero. Therefore, the environmental efficiency of this method is logically worse than the (45% / 58%) compromise in energy-consumption to provide a certain advantage while reducing CO ₂ . By extrapolating this to the case of 100% CO ₂ removal, the MEA / liquefaction / disposal method required more than 70% of the plant energy. It should be noted that typical competing absorption technologies cannot achieve 100% absorption; that is, a 58% CO ₂ absorption value was for a plant that treated 100% of the flue gas outlet.

B. Calculation of environmental performance limits [(∂CO ₂ / ∂E) max]

In practice, for a given system that effectively converts all NaOH produced into NaHCO ₃ as a result of CO ₂ absorption, the main energy component is kWh / mol NaOH. Despite the fact that the energy per mole of NaOH is proportional to both voltage and current, the current is set by stoichiometry of chemistry. Thus, the energy consumed per mole of CO ₂ is primarily optimized by achieving the lowest voltage condition at which hydroxide is efficiently produced.

The minimum voltage at which the electrolysis system, according to the embodiments of the present invention, operates (generated by varying concentrations, geometric dimensions, flow rates, and so on) can be determined by observing the current density (kA / m ² ) from the V characteristics of the system and determining the lowest voltage at which sufficient, non-zero current densities are obtained in the production of the product. Changing the physical dimensions, electric field generation devices, sizes of the electrolyzer, composition of materials, and technological conditions to optimize this characteristic indicator (kA ² / m ² in) is the main method for optimizing these systems, and typical industrial plants are used to optimize the industrial process experiment planning methods.

Leaving aside the practical limitations, there is one fundamental limitation that will be applicable to all systems with a given H / Na ratio (degree of protonation), namely:

(1) No device can operate if it produces more energy as a result of the extraction of hydrogen energy than is introduced into the electrolysis system. For those who are familiar with the principles of thermodynamics, this will mean "violation of the second law"

(2) As a result of this fact, the fundamental thermodynamic limit can serve as a boundary, determined by the choice of the H / Na ratio used in the flow of anolyte:

a. For this limit, it was assumed that H / Na was 0.10.

b. The efficiency of hydrogen energy return was set to 100%.

c. The lowest voltage was calculated at which a process can occur in which the resulting energy consumption of the system is zero ("Vmintheo") (that is, the point at which the electrolysis costs are equal to the estimated 100% hydrogen return efficiency).

d. In this example, this low voltage is 1.5526 V. This number is highly dependent on the Na / C ratio, the H / C ratio, and the hydrogen energy return efficiency. In this optimal case, Na / C is 1.0 and H / C is 1.0.

e. Bringing it directly to the calculation of ecological efficiency, the only solution is approximately 7% power for 93% absorption / conversion of CO _2.

f. Operation at more efficient operating points than this theoretical minimum is possible due to:

i. replenishment of energy consumption due to the utilization of waste heat; and

ii. process energy absorption / conversion part or entirely from the energy production of which does not result in the release of CO ₂ (hydroelectric, solar, wind, nuclear, etc.).

Similarly, assuming the ideal efficiency of hydrogen energy recovery, and so on, as in the above case, we calculated the maximum voltage at which the operation can be "environmental"("Vmaxeco") (that is, at which the process of absorption / conversion of CO ₂ removes more CO ₂ than he creates it):

a. The H / Na, Na / C ratios, and hydrogen energy return efficiency were set to 1.0, 1.0, and 100%, respectively, as in the above case.

b. The stress was calculated at which the removal of CO ₂ would be 50%.

c. In this example, this Vmaxeco value is 4.457 in. Under this stress and condition, the process proceeds along the line ∂CO ₂ = ∂E, the boundary between environmentally friendly and environmentally hazardous operation.

Consequently, environmentally friendly operation will be carried out when the electrolysis system is operated between Vmintheo (1.5526 V) and Vmaxeco (4.457 V). Operation between these two points can be reproduced in many typical electrolysis systems. Laboratory results at voltages equal to or lower than 2.1 V can be easily reproduced by changing the size, concentration, temperature, pressure, flow rate, and so on, of electrolyzers designed in this way.

C. Impact of non-greenhouse energy on environmental performance

When additional energy (which supplies energy to the method) is provided by energy produced without greenhouse gas (GHG) emissions (for example, wind, hydroelectric, solar, nuclear and so on), then the emissions of additional CO ₂ are zero, and the environmental efficiency of the present invention is greatly improved. For this example, the value of 3101209 tons of CO ₂ in table 2 is removed during all subsequent iterations, and so on, obtaining this simplified result: all CO _{2 is} absorbed / converted (7446069 tons), and the total energy required is just 8760000000 basis plus 3648448267 kWh required to perform the work necessary to absorb / convert this basic amount of CO ₂ , while requiring only 29% of the total energy to provide 100% CO ₂ emissions in a method supplied with energy with the emission of non-greenhouse gases, compared with 41% of the total en WGIG to ensure 100% CO ₂ emissions in the process, is supplied with energy with the release of greenhouse gases. This means that embodiments of the present invention provide a significant leverage factor when the process is supplied with energy generated without greenhouse gas emissions. Instead of using energy generated without greenhouse gas emissions to replace energy generating greenhouse gas emissions on a 1%: 1% basis, if energy generated without greenhouse gas emissions is used instead to power processes that are some embodiments of the present of the invention, 1% of the energy obtained without greenhouse gas emissions then prevents the generation of greenhouse gases in an amount raised to a power even greater than in the examples of energy obtained with the formation of greenhouse gases. It is easy to imagine cases in which a given nation, state, or economic entity, using a certain part of the energy obtained without emitting greenhouse gases, when used in such a wonderful way, could more effectively achieve any goal of reducing CO ₂ ; that is, in a very efficient manner, “clean” energy can be used to clean up emissions of other “dirty energy”.

If in some applications it is possible to obtain energy without the formation of greenhouse gases, sometimes in sporadic forms (for example, solar, "wind power centers", etc.), it is extremely beneficial to utilize this energy to produce large amounts of absorbent material during periods of partial load.

Example 6

Environmental performance of various simulated power plants

In FIG. 10 shows the environmental performance of various simulated power plants, including embodiments of the present invention, and presents various conditions that are major factors in determining environmental performance, (∂CO ₂ / ∂E).

From these calculations, we can generally conclude that:

(1) In the case of the formation of sodium carbonate and the use of standard conditions for the production of chlorine and alkali, the method would have environmental efficiency> 1, and although such operation may be economically acceptable, it produces more CO ₂ than it absorbs.

(2) Changing the food balance in favor of obtaining sodium bicarbonate increases the environmental efficiency of the method. In the event of a change in conditions to produce virtually pure sodium bicarbonate, this advantage is fully optimized.

(3) The introduction of low voltage electrolysis technology greatly shifts the process to a work area characterized by environmental efficiencies lower than 1.0 (that is, environmentally friendly CO ₂ absorption and conversion processes). Optimizing each physical model of the electrolysis system for optimum (kA / m ²⁾ and maximum hydrogen energy production leads to further improvement in ecological efficiency.

(4) The combination of the absorption / conversion process of the embodiments of the present invention with any number of available or specially designed units that convert waste heat to direct current, initial energy investment in electrolysis at direct current and pumping at alternating current, and so on may be additional or fully implemented through the utilization of waste heat.

It should be noted that providing the invention with energy from energy sources obtained without greenhouse gas emissions allows the method to directly achieve 100% absorption of CO ₂ (see discussion in Example 5).

Example 7

Determination of Voptlve (optimal operation at low voltage relative to the capacity of the cell or its area) and Ioptlve (current at operating voltage), for a given V / I characteristic working line for the optimized NVE of the chlorine-alkali cell

It has been shown here that low voltage operation reduces the energy required to produce sodium hydroxide, which is used as a liquid absorbent. Table 3 shows the calculations made for the VI working line in the case of anolyte with a pH of 1.0 / 90 ° C in FIG. 5 (discussed above in Example 4).

Several points should be noted regarding the contents of table 3:

(1) The third column, current output (dimensionless), represents part of the generated current used in the production of chemical products; Losses, such as I2R losses and waste heat from electrolytic fluids, are the main cause of energy loss. The current output drops with a voltage drop.

(2) The electrolytic cell area is normalized for the case of a voltage of 3.975 V (where the current density, and therefore the need for the process in the area, is identical to the standard 5 in electrolysis operating with a characteristic of 3.0 kA / m ² ). Calculate A2 / A1 (dimensionless).

The latter indicator,% of energy saved per dimensionless area, is plotted in FIG. 11 . For such a function, the point at which the maximum slope (change in energy use for voltage change) represents the optimum; that is, at low voltage (for example, 2.1-2.5 V) the slope (Δ energy / Δ m ² ) is relatively small, then at higher voltages (for example, from 2.5 to about 3), the slope ( Δ energy / Δ m ² ) increases to a larger value, and then drops to smaller slopes as the voltage continues to increase. This illustrates that there is a region of large slopes bounded by a region of small slopes on each side; that is, on either side of the Voptlve point, a change in energy use for voltage change is less effective.

First, a function was obtained that approximates operation behavior very closely (pay attention to the formula of the polynomial trend line obtained by the least squares method). In this example: y = -10.164x ³ + 88.256x ² -235.06x + 198.37 is a pretty good approximation. Then, the first derivative of the function was calculated by typical decomposition of polynomials: dy / dx = (3) (- 10,164) x ² + (2) (88,256) x-235,06 = max. The x (Volts) values can be iterated to find the maximum of this first derivative, which can be carried out by various methods, giving as a solution 2,894 c.

Note that voltages lower than 2,894 V can be used, and as a result, additional energy savings are obtained. Some preferred embodiments will optimize low energy operation below this “natural optimum point”. In these cases, the additional membrane area used will result in a “partially optimized” electrolysis system, but low energy operation for the total decarbonization process may be further beneficial by operating below this natural optimum for a given electrolysis subsystem. However, if this is done, then the voltage / energy advantage disappears after that, while the area factor continues to proportionally worsen operating efficiency per unit area.

The electric current and current density that correspond to this Voptlve can be determined, or a similar dependence for V and I can be obtained using the least squares method, or graphically, using the working curve to determine Ioptlve. In this example, the calculated value is 10.419 A (or for 0.01 m ^{2 the} area of the cell, as in this case), 1.042 kA / m ² .

TABLE 3 I
Current (A) Voltage
(at)
operating conditions: anolyte pH 1.0, 90 ° C Current output Δ (kA / m ² ) /
ΔV ΔEnergy /
Δ (kA / m ² ) Energy = V * I * Eff thirty 3,975 97% 1.82 0.55 115.67 25 3,7 96% 1.87 0.53 88.83

twenty 3,433 94% 1.87 0.54 64.62 fifteen 3,165 91% 1.80 0.56 43.34 10 2,887 88% 1.79 0.56 25.30 7.5 2,747 83% 1.74 0.58 17.15 5 2,603 78% 1,69 0.59 10.19 four 2,544 73% 1.49 0.67 7.41 3 2,477 67% 1.43 0.70 4.98 2 2,407 60% 1.19 0.84 2.90 1,5 2,365 54% 0.89 1.12 1.92 one 2,309 49% 0.64 1,57 1.13 0.7 2,262 44% 0.61 1.65 0.70 0.5 2,229 40% 0.62 1,60 0.44 0.4 2,213 36% 0.32 3.10 0.31 0.3 2,182 32% 0.56 1.80 0.21 0.2 2,164 29% 0.23 4.40 0.12 0.1 2.12 26% 0.00 212.00 0.05 I
Current (A) Area
electrolyzer A2 / A1 % of standard energy /
kgmol % Energy Saving at Low Voltage Δ energy savings / m ² thirty 1,000 77% 23% 25 1,302 71% 29% 22.22 twenty 1,790 65% 35% 19.77 fifteen 2,669 58% 42% 15.82 10 4,572 51% 49% 10.80 7.5 6,744 46% 54% 8.05 5 11,357 41% 59% 5.22 four 15,619 37% 63% 4.03 3 23,248 33% 67% 2.87 2 39,873 29% 71% 1.78 1,5 60,121 26% 74% 1.24 one 102,631 23% 77% 0.75 0.7 166,292 twenty% 80% 0.48 0.5 262,506 eighteen% 82% 0.31

0.4 367,228 16% 84% 0.23 0.3 551,770 fourteen% 86% 0.16 0.2 927,267 12% 88% 0.09 0.1 2103,359 eleven% 89% 0.04

All methods and devices described herein and set forth in the claims can be created and implemented in the light of the present description without further experiment. In the process of how the methods and devices of this invention were disclosed through preferred embodiments, it would become apparent to those skilled in the art that the methods and devices described here, and the steps or steps of the steps of the method can be changed without departing from the idea, essence, and scope inventions. More specifically, it would become apparent that similar results could be achieved by replacing the compositions described herein with related chemical compositions. It is assumed that all of these similar replacements and modifications obvious to specialists in this field fall within the essence, scope and idea of the invention defined in the attached claims.

MATERIALS USED IN THE DESCRIPTION

The following materials, to the extent that they provide reference data or other details in addition to those presented in the description, are specifically provided here by reference to them.

"Annual Energy Outlook 2005 (AEO2005)," prepared by the Energy Information Administration, available through National Energy Information Center, EI-30, Washington, DC.

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"Effects of the Operating Pressure On The Performance of Electrolytic Cell at Elevated Temperatures," by Y. Ogata, M. Yasuda, and F. Hine, Nagoya Institute of Technology, Japan. In Proceedings of the Symposium on Electrochemical Engineering in the Chlor-Alkali and Chlorate Industries, The Electrochemical Society, 1988.

"Electrochemical Hydrogen Technologies - Electrochemical Production and Combustion of Hydrogen," edited by Hartmut Wendt, Institute fur Chemische Technologie, Stuttgart, Elsevier Press 1990.

"Electrochemical Process Engineering - A Guide to the Design of Electrolytic Plant," by F. Goodridge and K. Scott, University of Newcastle upon Tyne, 1995.

"Exergy Study of the Kalina Cycle," by Goran Wall, Chia-Chin Chuang, and Masaru Ishida, presented at 1989 American Society of Mechanical Engineers (ASME), Winter Annual Meeting, San Francisco, California, December 1989, published in "Analysis and Design of Energy Systems: Analysis of Industrial Processes AES Vol. 10-3, pp 73-77 ASME.

"Industrial Electrochemistry," 2 <nd> Edition, edited by Derek Plefcher and Frank Walsh, 1990.

"Modern Chlor-Alkali Technology," edited by M.O. Coultier, Society for the Chemical Industry, London, 1980.

"Modern Chlor-Alkali Technology," Volumes 1-7, The Royal Society of Chemistry Information Services, 1998.

"Some Guidelines for Analysis, Design, and Operational Principles of Bubble Column Reactors," and other selected information contained in "Bubble Column Reactors," by Wolf-Dieter Deckwer, Gesellschaft fur Biotechnologische Forschung mbH, Braunsweig, Germany, translated by Robert Field, 1991 ISBN 0-471-91811-3.

"Transport and Structure in Fuel Cell Proton Exchange Membranes," by Michael Anthony Hickner, dissertation submitted to the faculty of Virginia Polytechnic Institute and State University, 2003.

Figure 8. "CO2 Absorption vs. (Low) Interfacial Gas Velocity at Various Column Heights of 0.18M Sodium Hydroxide Solution in an Unpacked Sparged Bubble Column," from A. Schumpe thesis, University of Hanover, 1978.

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Wie-rong et al., J. Zhejiang University Science, ISSN 1009-3095, 2004.

Claims (25)

1. A method of removing carbon dioxide contained in a gas stream, comprising the steps of:
(a) obtaining sodium hydroxide in an aqueous mixture;
(b) mixing sodium hydroxide with carbon dioxide obtained from the source to produce bicarbonate products or a combination of carbonate and bicarbonate products in a mixture; and
(c) the selection of these products from the mixture, thereby removing carbon dioxide from the solid mineral product, which contains 70 mass. % sodium bicarbonate;
where the production of hydroxide includes:
obtaining sodium salt;
mixing the salt with acid and water, acid and steam, or acid, water and steam, to obtain a protonated brine; and
electrolysis of protonated brine to produce sodium hydroxide;
where stage (b) includes changing the product equilibrium in favor of obtaining bicarbonate products. 2. The method of claim 1, wherein the amount of carbon dioxide is formed in the steps (a) to (c) and the amount of carbon dioxide is less than the amount of carbon dioxide removed in the steps (a) to (c). 3. The method according to claim 1, wherein the mixing takes place in two separate chambers, wherein one chamber is used to produce carbonate products, and the other chamber is used to produce bicarbonate products. 4. The method according to claim 1, wherein the mixing takes place in a bubble column or series of bubble columns. 5. The method according to p. 1, in which stage (C) includes a process of selection as a result of hot deposition. 6. The method according to p. 5, in which heat for the allocation process is obtained as a result of heat exchange with incoming flue gases. 7. The method according to p. 1, further comprising:
transportation of products to remote places of sequestration;
combining the products with acid in a neutralization reaction to produce pure carbon dioxide; and
carbon dioxide injection into the carbon storage. 8. The method according to p. 1, in which other components of the gas stream are neutralized and / or captured at the stages (a) to (c). 9. The method of claim 1, wherein the electrolysis of the protonated brine comprises burning a carbon source of fuel to produce a constant electric current. 10. The method according to p. 1, in which the amount of acid mixed with salt is based on determining the optimal protonation rate, which reaches the minimum energy to obtain reagents and maximum energy to remove from the products. 11. The method according to p. 1, in which:
protonated brine is electrolyzed in an electrolyzer having a catholyte space and anolyte space; and
products are recycled into the catholyte space of the electrolyzer. 12. The method according to p. 1, in which the energy consumed in the method is supplemented by the heat of the flue gases extracted from the gas stream. 13. The method of claim 1, wherein the salt is sodium chloride and the acid is hydrochloric acid. 14. The method of claim 1, wherein the gas stream is a plant exhaust stream. 15. The method according to p. 14, in which the installation is a power plant that uses as a carbon source of fuel. 16. The method of claim 13, wherein the electrolysis of the protonated brine also produces chlorine gas and hydrogen gas. 17. The method according to p. 16, in which gaseous chlorine and gaseous hydrogen are burned with the formation of hydrochloric acid. 18. The method of claim 16, wherein hydrogen gas is used to generate energy. 19. The method according to p. 17, in which hydrochloric acid obtained from gaseous chlorine and gaseous hydrogen is used to produce additional hydroxide. 20. The method according to p. 16, in which gaseous hydrogen is burned with atmospheric oxygen or oxygen from a warehouse of chemical products to produce water. 21. The method according to p. 18, in which stage (C) includes a hot deposition process, in which heat for the evolution process is obtained from the energy produced by burning hydrogen gas. 22. The method according to p. 18, in which gaseous hydrogen is burned together with coal to improve emissions of coal furnaces. 23. The method according to p. 18, in which gaseous hydrogen is used in the combustion process to obtain a constant electric current in the fuel cell. 24. A method of removing carbon dioxide obtained from a source and contained in a gas stream, comprising the steps of:
(a) obtaining sodium hydroxide in an aqueous mixture;
(b) mixing sodium hydroxide with carbon dioxide obtained from the source to produce solid bicarbonate products or a combination of carbonate and bicarbonate products in a mixture;
(c) isolating said solid products from the mixture, thereby removing carbon dioxide from the mineral product; and
where the production of sodium hydroxide includes:
obtaining sodium salt;
mixing the salt with acid and water, acid and steam, or acid, water and steam to produce a protonated brine; and
electrolysis of protonated brine to produce sodium hydroxide using a diaphragm electrolyzer, a membrane electrolyzer, or a mercury cathode electrolyzer;
where stage (b) involves changing the product equilibrium in favor of producing bicarbonate products, and where the amount of carbon dioxide is formed in steps (a) to (c) and the amount of carbon dioxide is less than the amount of carbon dioxide removed in steps (a) )-(from). 25. A method of removing carbon dioxide obtained from a source and contained in a gas stream, comprising the steps of:
(a) obtaining sodium hydroxide in an aqueous mixture;
(b) mixing sodium hydroxide with carbon dioxide obtained from the source to produce solid bicarbonate products;
(c) isolating said solid products from the mixture, thereby removing carbon dioxide from the mineral product; and
where the production of sodium hydroxide includes:
obtaining sodium salt;
mixing the salt with acid and water, acid and steam, or acid, water and steam to produce a protonated brine; and
electrolysis of protonated brine to produce hydroxide and hydrogen gas;
where stage (b) includes changing the product equilibrium in favor of obtaining bicarbonate products.

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