Patent Application
20160304344
Distributed hydrogen production has been sought to address a range of applications where pipeline hydrogen is not available. For example, it is not feasible to build conventional large-scale H2 gas plants for distributed and/or on-demand power generation using, e.g., H2 fuel cells, various H2-needed hydrogenation and/or hydrotreating processes, etc.
Hydrogen production typically involves steam-reforming, water-gas-shift reactions, and gas purification steps. However, the reforming reaction is an energy-intensive process step because it is an endothermic reaction and involves a large capital cost. In addition, process flow diagrams involving these conversion steps are too complicated to economically build a small-scale H2 plant, because economy of scale problems exist with conventional unit operation approaches. Another problem is that catalysts often deactivate rapidly due to use of biomass-derived fuel sources. For example, in the literature, catalysts used for steam-reforming of bio-oils can deactivate within a few hours.
Renewable hydrocarbon resources such as alcohols and bio-oil are attractive fuels for hydrogen production. Bio-oil can be derived from various biomass raw materials. However, unique challenges remain to be addressed for production of hydrogen from bio-oil compared to conventional hydrocarbon steam reforming processes. For example, bio-oil contains large amounts of water and oxygen, and exhibits a relatively low heat of combustion. Thus, bio-oil is suitable for steam reforming reactions that require much less stoichiometric amounts of oxygen (from either steam or oxygen gas) than pure hydrocarbons. A significant fraction of the bio-oil is non-volatile.
Various catalysts containing transition metals and precious metals, testing conditions, and various feedstocks have been reviewed in the literature by Chattanathan et al. and studied with model compounds and actual bio-oil.
Some literature reports on Bio-oil reforming at low temperatures have been published. Some catalyst materials have been compared to Ni/alumina, a common reforming catalyst system. Newer catalyst materials show an ability to conduct the bio-oil reforming reaction at a much lower temperature compared to steam reforming of methane. However, stability of these newer catalysts has not been reported. Further, reported steam/oil ratios in the literature are yet too high for practical application.
Reaction temperatures used in the literature are typically above 500° C. Catalyst deactivation is a prominent problem reported in most references. And, long-term catalyst stability has not been achieved or reported in the literature, which is attributed to coke formation as a major cause.
Another feature of the present invention involves CO2 sorbent systems. Capturing CO2 at temperatures of 200 to 700° C. is useful to both existing process improvements and future CO2 emission control. Removing CO2 under reaction conditions can shift the reaction equilibrium toward completion1-7, such as steam reforming and water-gas-shift reaction processes that are widely used for syngas and hydrogen production in today's industries. Hot flue gas CO2 capture for CO2 sequestration purposes is viewed beneficial to efficient thermal energy utilization.8-10 CO2 sorbents and sorption processes have become a very active research field for a recent decade. The materials studied so far are grouped into three categories: CaO-based,8-19 MgO-based,20-26 metal oxide compounds27-31 (such as lithium silicate and dolomite). The MgO-based sorbent works at moderate temperatures (200-400° C.), while the CaO-based works at relatively high temperatures (600-800° C.). Sorption of CO2 on these sorbents is a gas/solid reaction process. One problem with this approach is degradation or de-activation of the sorbent in the presence of steam and over repeated sorption/regeneration cycles. Long-time cyclic stability is necessary to reduce the material consumption and costs for practical applications. For example, about 24,000 sorption/regeneration cycles in 20-min cycling time are need in order for a sorbent to have 8000-h service time. Another problem is about CO2 sorption and sorbent regeneration rates. The slow kinetics of existing sorbent materials requires a large reactor vessel.
Significant efforts are made by the research community toward enhancing the sorbent stability and kinetics. CaO-based sorbents are the most extensively studied one, because of in-expensive and abundant resources. The sorbent performances are improved by dispersing CaO grains into or onto more durable ceramic matrices15-17, coating the CaO grains with more stable ceramic nano-particles18 such as ZrO2, or making CaO into hollow spheres.19
Many publications have reported various CO2 sorbents used for CO2 capture and sequestration. A shortened list of CO2 sorbents that might work under bio-oil steam reforming conditions published during the most recent five years is summarized in TABLE 3. Several articles have reported integrating in-situ CO2 capture with steam reforming of methane. Others have dealt with bio-oil steam reforming reactions with in-situ CO2 capture.
Steam reforming of methane with in situ CO2 capture on a CeZrOx—CaO sorbent produced >95% pure H2. CaCeZrOx (10:1:1) appears to be a promising sorbent but requires a regeneration temperature of 700° C.
Li2CO3—CaO sorbent enhances yield from methane reforming and produces pure hydrogen during a first cycle as high as 99.8 vol %. However, capacity decreases with carbonation-regeneration cycles because the Li dopant escapes from the sorbent and the aged sorbent behaves like CaO, which is known to be unstable during the regeneration cycle.
The working temperature for CaO-based sorbents is also above 700° C., which is too high for most practical applications. Reaction conditions and reactor bed structures reported in the literature are also different from those described herein. Furthermore, a stable CO2 sorbent has not yet been demonstrated in a sorption/regeneration cycle.
The present invention presents the material, bed structure, and reactor innovations that provide advantages over systems described in the prior literature.
In a first aspect, the invention provides a reactor for producing hydrogen, comprising: a reaction channel comprising a steam reforming catalyst or water gas shift catalyst, the reaction channel comprising one or more channel walls; wherein at least one reaction channel wall is porous to carbon dioxide; a sorbent chamber adjacent the reaction channel, separated from the reaction channel by one or more channel walls; and comprising a CO2 sorbent; and wherein the reaction channel comprising the steam reforming catalyst has an average diameter of between 0.1 mm and 10 mm.
The invention, in any of its aspects, can be characterized by one or more of the following: wherein the reaction channel comprising the steam reforming catalyst has an average diameter of 10 mm or less, or between 0.1 mm and 10 mm; preferably 0.5 mm and 5 mm; and wherein the reactor comprises a plurality of parallel reaction channels; preferably at least 10; more preferably at least 100; wherein the catalyst is a steam reforming catalyst and wherein the steam reforming catalyst is a wall coating and wherein an open bulk flow path is present in reaction channel; wherein the steam reforming catalyst comprises Ni, Cu, Mg, Ce, and Al; wherein the reactor comprises a honeycomb monolith; wherein the reactor comprises a plurality of parallel reaction channels alternating with plurality of parallel sorbent chambers; preferably at least 10 parallel reaction channels alternating with at least 10 parallel sorbent chambers; wherein sorbent fills a cross-section of the sorbent chamber; wherein the reactor channel or channels each have an inlet and an outlet; and wherein the sorbent chamber or sorbent chambers having no ports or closed ports such that gaseous flow into and out of sorbent chamber is only through the porous reaction channel wall(s); wherein the cross-section of the reaction channel (and/or sorbent chamber), in the direction orthogonal to flow through reaction channel, is square, rectangular, or circular; wherein the reactor comprises a honeycomb monolith; and wherein the monolith comprises: a macro- and meso-porous material (macroporous refers to pores in the range of 50 nm-100 μm; mesoporous 2 to 50 nm); a metal oxide; titania, alumina, silica; or a dense material such as metal with apertures for gas flow; wherein the sorbent surrounds the reaction channel; wherein a single sorbent chamber surrounds plurality of (preferably parallel) reaction channels; wherein the porous channel wall has a volume average pore size that in the range of about 1 nm to 100 μm, in some embodiments about 2 nm to 1000 nm. The sorbent in the reactor can be any of the sorbent compositions described herein. In some preferred embodiments, the reaction channel wall is not a membrane or a sorbent (in some preferred embodiments the reaction channel wall is inert) but is porous to all gases; this configuration provides for faster operation and greater stability.
The invention includes methods of steam reforming, comprising passing a hydrocarbon and steam into any of the reactors described herein.
In another aspect, the invention provides a CO2 sorbent system is a composite material comprising: Mg-based phase (MgO, Mg(OH)2, MgCO3, or mixture thereof), Ca-based phase (CaO, Ca(OH)2, CaCO3, or mixture thereof), and a promoter phase comprising a mixed alkali metal carbonate having at least 2 alkali atoms selected from the group consisting of Li, Na, and K; and wherein the atomic ratio of Mg:Ca is in range of 0.3 to 3.3; preferably 0.5 to 2.
Preferably, the carbonate is stable to 800° C. Preferably, the sorbent system comprises 3 essential phases: structural stabilizer; active metal oxide phase; and promoter. Following the techniques described herein, the sorbent system can be characterized by elemental analysis, x-ray diffraction (XRD) analysis and scanning electron microscopy (SEM) analysis. Preferably, the alkali in the promoter phase comprises at most 80 atom % of any one of the 3 elements, and in some preferred embodiments, includes at least 10 atom % of each of the three alkali elements Li, Na, and K.
In some preferred embodiments, the sorbent can be further characterized by one or any combination of the following characteristics: wherein the promoter is preferentially dispersed in the composite uniformly and is substantially free of distinctive crystallites or segregated particles; wherein the sorbent comprises 2-30 wt % or 5-20 wt % or about 5-10 wt % of the promoter; comprising essentially no nitrate, or less than 0.5 wt % N, 0.3, 0.2, 0.1% N; or wherein the CO2 sorbent system consists essentially of MgO or CaO and a promoter comprising a mixed-metal alkali metal carbonate having at least 2 alkali atoms selected from the group consisting of Li, Na, and K; wherein the promoter comprises Li, Na, and K, in an atomic ratio of 1-5 Li:1-5 Na: and 1-5 K; wherein the sorbent system is at a sorption temperature of from 300 to 450° C. and a regeneration temperature from 300 to 500° C.; the sorbent system is at a sorption temperature of from 350 to 700° C. and a regeneration temperature from 650 to 800° C.; wherein the sorbent comprises a porous, solid matrix and a liquid phase (molten carbonate comprising at least 2 of Li, Na, K); comprising dolomite (CaMg—CO3) as the Mg and Ca-phases (the dolomite may or may not be pre-calcined); comprising an inert solid support phase, an active solid working phase, and a liquid-phase promoter; wherein the Mg-based phase and Ca-based phases are identifiable by X-ray diffraction (XRD) measurements, and the promoter phase is not be identifiable by XRD or its XRD peaks are substantially smaller than the Mg and Ca-based phases if its crystalline phase is detectable; wherein the promoter is molten and the sorbent system is at a temperature below the melting temperature of pure Na or K carbonate.
In a further aspect, the invention provides a method of converting hydrocarbons to H2, comprising: passing a carbon fuel and water into the reactor; wherein the water is in the form of steam or wherein the water is converted to steam within the reactor; wherein the reaction channel is at a temperature of at least 350° C.; wherein the steam and carbon fuel react in the reaction channel at a temperature of at least 350° C. to produce at least H2 and CO2; wherein at least a portion of the CO2 flows into the sorbent chamber where it is sorbed by the CO2 sorbent; and regenerating the CO2 sorbent. The removal of CO2 has the net effect of driving the equilibrium of the water gas shift reaction toward the synthesis of additional H2.
In some preferred embodiments, the method can be further characterized by one or any combination of the following characteristics: wherein the carbon fuel is a hydrocarbon and wherein the reactor comprises a steam reforming catalyst; wherein sorption of CO2 in the sorption chamber generates heat that passes through the reaction chamber wall and adds heat to a steam reforming reaction that occurs in the reaction channel; wherein the CO2 sorbent is regenerated by increasing the bed temperature and/or reducing partial pressure of CO2; wherein the reactor is taken off line (the flow of hydrocarbon is discontinued) while the CO2 sorbent is regenerated; wherein H2 produced in the reaction channel passes out through an outlet; wherein the reforming and sorption are performed at a temperature less than or equal to about 700° C. in a time of less than about 10 minutes; wherein the reforming and sorption are performed iteratively in a plurality of cycles with a delay between any subsequent cycle of less than or equal to about 10 minutes; wherein regenerating the reforming catalyst and the CO2 capture sorbent with an oxygen-containing gas is conducted at a temperature less than or equal to about 700° C. (or 750 C) in a time of less than or equal to about 10 minutes; wherein the steam reforming catalyst is a Ni-oxide based catalyst comprising cerium oxide, titanium dioxide, a zeolite, alumina, or combinations thereof; wherein the CO2 sorbent has any of the characteristics described herein; wherein at least 30% or at least 50% of the heat required for the steam reforming reaction is provided by the CO2 sorption; wherein the steam-reforming reaction is performed in the catalyst bed absent a downstream partial oxidation (PDX) step or a downstream auto-thermal reforming (ATR) step; and/or wherein the hydrocarbon feedstock is a bio-oil.
The invention can also be characterized by any of the features or data provided in the descriptions or examples. For example, the sorbent could be characterized by the cycling test methods described in the examples while maintaining sorption capacity as shown in the examples. An example would be: wherein the CO2 sorbent is subjected to at least 8 regeneration cycles and the CO2 sorbent loses less than 5% of its sorption capacity; or wherein the CO2 sorbent is subjected to at least 8 regeneration cycles and the CO2 sorbent gains at least 5% of its sorption capacity. Furthermore, any aspect of the invention can, in some embodiments, be characterizable by any of the data from the examples ±10% or ±20%; either as a process characteristic or as a physical property when measured under the conditions described in the examples.
Various embodiments of the invention can provide significant advantages. One innovative feature made possible by the inventive system is process equipment elimination and/or minimization compared to conventional hydrogen production. Bio-oil steam reforming (SR) is a highly endothermic reaction requiring significant amounts of heat. Heating the reactor tube in a furnace for SR or using partial oxidation (PDX) or auto-thermal reforming (ATR) are the conventional ways to supply the reaction heat. SR requires a large high-temperature furnace, while PDX and ATR need an air separation unit to supply oxygen. Both methods add complexity and cost. The inventive approach can utilize in-situ heating of the catalyst bed by alternating reforming (endothermic) and regeneration (exothermic) reactions, and by coupling of the endothermic reforming with exothermic carbonation reaction. This can eliminate an external combustor for SR or oxygen separator for PDX/ATR. Produced CO2 is captured in-situ with a solid sorbent that is integrated with the SR catalyst into one reactor.
Another innovative feature is minimization of process steps. In situ CO2 capture shifts thermodynamic equilibrium of SR reactions toward higher one-pass conversion to H2. Thus, a separate water-gas-shift (WGS) process step is avoided. And size of the gas purification unit such as pressure swing adsorption (PSA) can be reduced. The literature reactors typically operate at high temperatures (570-800° C.) presenting a challenging environment for sorbent stability. The invention can, in preferred embodiments, lower the reaction and CO2 capture temperatures to moderate regimes (<700° C.) using new reforming catalysts and CO2 sorbents.
Another innovative feature is to reduce catalyst deactivation, which is known as a major problem for reforming of bio-oils. Coverage of the catalyst pores by stagnant bio-oil liquids results in physical deactivation, while chemical deactivation is mostly due to coking of active sites. The physical deactivation can be mitigated by atomizing the bio-oil into small (<channel size) droplets using steam and/or by providing a smooth catalyst external surface to minimize accumulation of dead liquid pockets. The smooth surface of catalyst-coated monolith channel walls would enable uniform spread and contact of un-vaporized liquid droplets with the catalyst. The chemical deactivation is due to nature of catalysis chemistry. Bio-oils are comprised of some oxygenates that are potent catalyst coking agents. The deactivated catalyst is regenerated by burning off the coke. In the present invention, rapid swing reactor operation allows for catalyst deactivation by enabling frequent in-situ regeneration. As a result, the reforming reaction can be run at high space velocity and moderate temperatures on a highly active catalyst that deactivates. The swing reactor operation enables utilization of highly active catalysts that would not be possible in conventional steady-state reforming operations.
Various preferred embodiments of the present invention may differ from the prior art in several respects:
Different reaction conditions. Particularly, the reaction temperature for the present invention is lower than the most methane reforming and bio-oil reforming reactions reported in the literature. In addition to the temperature, the reactor will be operated by rapid swing between the reaction/sorption and regeneration in the order of minutes.
Different CO2 sorbent and reforming catalyst materials. The CO2 sorbent will be a composite of MgO-based solids and CaO-based solids mixed at micrometer or nano-meter scales, and promoted by molten carbonates (KLiNa— CO3). The catalyst material will be transition metal oxide-alkaline earth/rare earth metal oxide composites such as NiCu—MgCe, and the metal oxide composite combined with acidic catalysts (gamma-alumina, Y, ZSM-5). The catalyst provides redox catalytic activity and hydrocracking activity. The catalyst is in an active state after air calcination.
Different reactor bed structures. In this invention, the sorbent and catalyst inside the reactor can be made in a monolithic structure form, as compared to conventional particles and/or extrudate.
Although various structures have been proposed in the prior art with regard to microchannel apparatus, the inventive systems are characterized by having a number of reaction channels being housed inside a conventional reactor vessel or tube (reactor diameter >10 mm) and generating synergistic effects made possible by the paring of CO2 sorbent with an endothermic H2 forming reaction.
A “porous” sorbent has porosity that aids in providing a high surface area; the porosity is not necessarily through-porosity but is the well-known porosity in macro and/or mesoporous materials or smaller porosity materials such as catalyst supports that provide high surface area. Porosity as measured by N2 BET gas adsorption (for smaller pore sizes) and mercury intrusion porisimetry (for larger sizes). The “porous” channel wall has through-porosity such that CO2 can pass through the channel wall.
A “steam reforming catalyst” can be any steam reforming catalyst known in the literature, and is preferably selected to optimize reaction at the desired temperature for operation. Likewise a “water-gas shift catalyst” can be any material known in the literature for catalyzing the water gas shift reaction.
An “open bulk flow path” is an unobstructed flow path having a diameter of at least 50 μm.
As is conventional terminology, the invention can be described as “comprising” meaning “including.” In any of the embodiments, described as “comprising,” the invention may alternatively be described more restrictively as “consisting essentially of” or “consisting of” selected features, properties or elements.
The patent or application file contains at least one drawing executed in color. Copies of this patent or patent application publication with color drawing(s) will be provided by the Office upon request and payment of the necessary fee.
One of the advantages of the present invention is that it can reduce the size and complexity of hydrogen production processes and devices. Compared to the conventional hydrogen production processes (
The following are the likely reactions occurring in the bed.
CnHm+nH2O=(n+m/2)H2+nCO steam reforming
CO+H2O=H2+CO2 water gas shift
CnHmOk→CxHy+kH2O Coking & de-oxygenation reaction
CO2+MO=MCO3 CO2 capture
CO2 produced by the reforming reactions is captured with a metal oxide sorbent (MO). The exothermic carbonation reaction supplies heat to the endothermic SR reaction. The coking reactions are also exothermic. In the short term, aromatic-type coke containing some H element is typically formed which is reactive to oxidation. Graphite-type coke (nearly free of H), which tends to occur in the long term, can be avoided by frequent regeneration.
When the CO2 sorbent is saturated and H2 yield becomes low due to decreased temperature and catalyst deactivation, the reactor is switched to regeneration mode. Hot air or O2-containing gas is introduced into the catalyst bed to burn off the coke. As the catalyst bed is heated up, the metal carbonate decomposes to release CO2, thus regenerating the CO2 sorbent. The heat of carbonate decomposition may be partially or completely supplied by the combustion reaction. Residual fuel gas, which is rejected from the downstream PSA purification process, can be fed into the reactor undergoing regeneration and burned to supply additional heat. After regeneration, the reactor is switched back to the SR mode. In this way, the catalyst is maintained in an active state so that the SR reaction can be run at high space velocity.
The compositions and structures of a preferred catalyst system of present invention are illustrated in
The reforming catalyst and CO2 sorbent materials are made into monolithic bed structures in the reactor.
In the monolith bed design, the SR catalyst, WGS catalyst (optional), and CO2 sorbent can be integrated in the same reactor thereby minimizing mass and heat transfer resistance between the SR reaction and carbonation reaction. Essentially, all these sequential reactions can occur at the same spot.
Compared to the packed bed, the monolith bed includes the following advantages:
Straight flow channels minimizes the dead space in the bed, avoids accumulation of residuals and particulates in the bed, and enables rapid changeover of reacting fluids;
Straight flow channels provide a low pressure drop and minimize parasitic power consumption during regeneration;
the reactions are confined inside individual channels and temperature run-away due to auto-thermal reactions can be mitigated; and
adsorbent and catalyst materials are integrated into a fixed structure, and particle attrition and bed upset can be avoided during rapid swing operation.
The monolithic reactor also provides unique performance attributes compared to fluidized beds. Fluidized beds have been studied for both bio-oil SR and biomass gasification. It is a commercially-proven technology for coupling of endothermic and exothermic reactions, and one successful example is the fluid catalytic cracking (FCC) process in refineries. However, the circulating fluidized bed (CFB) system involves complex designs and requires highly specialized technical resources to operate and maintain. For a stand-alone H2 production plant with a capacity that is orders of magnitude smaller than the FCC process, it will be difficult to make such a unit economical. The proposed monolith reactor is compared to CFB as follows:
there is no loss of catalyst and sorbent in the monolith reactor, while significant catalyst loss could occur in the fluidized bed due to severe particle attrition and entrainment in reactor effluents;
the monolith reactor enables the CO2 sorbent coupled with the reforming catalyst in a layer-by-layer structure, while sorbent and catalyst particles can be segregated in CFB; and
the monolith reactor has a simple, modular-type structure that can be scaled up by numbering up of reaction channels. Automotive catalytic converters and diesel particulate filters are examples of how compact a monolith reactor can be made. In contrast, CFB designs are significantly affected by their processing capacity scale and the unit must be large to obtain high productivity.
The monolith reactor makes it possible to develop a compact H2 production plant with a simple process flow diagram (
As illustrated in
After the sorbent was tested through tens of sorption/regeneration cycles, a fully-activated working sorbent containing the three major crystal phases, MgO, CaO, and Ca(OH)2, was formed, as evidenced by XRD analysis. As expected, MgO crystal phase was intact during the sorption and regeneration as a structural stabilizer. The CaCO3 crystal phase was not present in an activated sorbent or its presence is minor. CaO can be converted into Ca(OH)2 in the presence of water vapor. The LiNaK-carbonate phase remained dispersed and was not detectable by the XRD. The following are possible chemical reactions involved in the sorbent preparation and sorption/regeneration processes.
Partial decomposition of dolomite at temperatures lower than 500° C.:
CaMg(CO3)2(s)→MgO(s)+CaCO3(s)+CO2(g) (1)
Fully-decomposed dolomite at temperatures higher than 700° C.:
CaMg(CO3)2(s)→MgO(s)+CaO(s)+2CO2(g) (2)
In the presence of H2O
CaO(s)+H2O(g)→Ca(OH)2(s) (3)
CO2 sorption/regeneration:
CaO(s)/MgO(s)+CO2(g)CaCO3(s)/MgO(s) (4)
Ca(OH)2(s)/MgO(s)+CO2(g)H2O(g)+CaCO3(s)/MgO(s) (5)
As illustrated by SEM images, the as-prepared and activated sorbent have similar micro-structures. Ca, Mg, and alkaline metals are all well mixed at micrometer scale in both fresh and activate sorbents as a composite rather than as segregated particles. The molten carbonate works as a phase transfer catalyst to facilitate the sorption and regeneration reaction process.
The dolomite-based composite sorbent was prepared in the following way. The Li2CO3—Na2CO3—K2CO3 eutectics (LiNaKCO3) has a melting point about 397° C. It was prepared by mixing 32.2 wt. % Li2CO3 (99%, Sigma Aldrich), 33.3 wt. % Na2CO3 (99%, Sigma Aldrich) and 34.5 wt. % K2CO3 (99%, Sigma Aldrich). The powder mixture was ground together and then heated to 410° C. for 1 hr. The resulting eutectic was crushed and ground to a fine powder. Dolomite from City Chemical was used as received of after calcination at 850° C. for 4 hours in air. The dolomite was mixed with LiNaK—CO3 at desired weight ratio by ball milling. The solid mixture was added into a Nalgene plastic bottle and mixed with 2-propanol (EMD Chemicals, Canada) and zirconia beads (diameter: 0.3-1.0 cm). The bottle was rotated at a speed of 150-200 rpm for 48-72 hours. The resulting slurry was dried at room temperature (RT) to allow evaporation of 2-propanol. Following drying, the cake was calcined at 350° C. or 725° C. in air in an alumina crucible for 3 hours. The comparative CaO sorbent was prepared by calcining Ca acetate (Alfa Aesar, 97%) at 800° C. for 3 hrs. The comparative MgO sorbent was obtained by calcining Mg5(CO3)4(OH)2.xH2O power (99%, Sigma Aldrich) at 450° C. in air for 3 hours.
Dolomite alone sorbent was prepared by heating an as-received dolomite mineral (City Chemical) was at 850° C. for 3 hours in a furnace to obtain complete decomposition.
MgO was obtained by calcining Mg5(CO3)4(OH)2.xH2O power (99%, Sigma Aldrich) at 450° C. for 3 hours in air. MgO was mixed with Na2CO3 (99%, Sigma Aldrich), Li2CO3 (99%, Sigma Aldrich), and K2CO3 (99%, Sigma Aldrich). The mixture contained 44 wt % MgO, 48 wt % Na2CO3, 3.9 wt % Li2CO3 and 4.2 wt % K2CO3. Typically, 33.5 g of the solid mixture was mixed with 60-100 g of 2-propanol (EMD Chemicals, Canada) and 120 g-200 g of zirconia beads (diameter: 0.3-1 cm). The bottle was rotated for 48-72 hours at a speed of 150-200 rpm. The obtained slurry was dried at RT ° C. in plates to evaporate 2-propanol and form thin cake (0.2-0.5 mm). Following the dry, the cake was calcined at 450° C. for 3 hrs. The calcined cake was crushed and sieved to 40-100 mesh particles for fixed-bed tests.
Performance characteristics of sorbents for CO2 capture under steam-reforming conditions were tested on a thermo gravimetric analyzer (TGA, Netzsch Thermal Analysis, STA 409 cell) at ambient pressure. The sample weight for each test was approximately 20 mg. CO2 absorption was evaluated by heating sample in 10-100% CO2 to 500-600° C. Cyclic CO2 absorption and desorption were evaluated by exposing the sample to the CO2 gas for 50 minutes at 550° C. and then, raising the temperature in N2 flow and holding at 700° C. for 60 minutes. The gas flow rates for the sorption and regeneration processes were maintained at 70 ml/min and 90 ml/min, respectively. The heat of absorption was measured along with the TG tests through differential scanning calorimetry (DSC). The TGA tests help to identify CO2 sorption and regeneration temperatures, measure CO2 sorption capacity, and assess sorption and regeneration kinetics.
For comparison, performances of CaO sorbents are shown in
Furthermore, sorption/regeneration performances of the CaO sorbents with and without the promoter are compared at different sorption temperatures but under constant regeneration temperature at 700° C. in
Because Mg and Ca are cheap alkaline earth metals, their compounds are attractive materials for CO2 capture application. Dolomite is a natural mineral of Mg and Ca carbonate compounds. At lower temperature, MgCO3 in dolomite is decomposed, whereas CaCO3 would decompose at a higher temperature.
MgCa(CO3)2→MgO+CaCO3+CO2 at 400-450° C.
MgCa(CO3)2→MgO+CaO+CO2 at 700-850° C.
To investigate feasibility of using this raw material for present application, the pre-calcined dolomite powder was tested under two sets of sorption/regeneration conditions (
CaO(s)/MgO(s)+CO2(g)CaCO3(s)/MgO(s)
In
The above dolomite composite sorbent can be used at lower working temperatures, in which MgO becomes an active phase while CaCO3 functions as an inert phase.
MgO(s)/CaCO3(s)+CO2(g)MgCO3(s)/CaCO3(s)
The sorbent was tested under three sets of sorption/regeneration conditions. The solid green and red lines in
Performance characteristics of a different composite sorbent are shown in
The experimental results conducted this quarter for CO2 sorbents are summarized in table 4. The sorbents that work at sorption temperatures 400° C. with regeneration temperature <600° C. are identified. More experiments are being conducted to find the sorbents with sorption temperatures from 400 to 550° C.
To confirm performances of the composite sorbent design, a larger number of sorption/regeneration cycling tests were performed with two dolomite sorbents of different promoter loading levels. The sorption was conducted at 550° C. in 100% CO2 and regeneration was performed at 700° C. in N2. Sorbent #1 loaded with 5 wt. % promoter was tested for 16 cycles. Sorbent #2 loaded with 10 wt. % promoter was tested for 8 cycles. The results are plotted in
TGA screening tests revealed strong synergism of the promoted dolomite sorbent. Impacts of the sorbent preparation on its cyclic performances were studied. The results are summarized in table 5. The dolomite can be used as received or pre-calcined at 850° C. in air for hours. The 850° C.-calcination converts the carbonate into Mg and Ca oxides. Addition of the promoter into the dolomite was evaluated at two different loading levels (5 wt. %, 10 wt. %). The resulting dry mixtures were calcined at two different temperatures (350, 725° C.) prior to TGA tests. Comparative pairs of sorbents #1 vs. 2, 3 vs. 4, and 5 vs. 6 show the impact of dolomite starting materials on the CO2 working capacity at the same promoter loading level and at the same calcination temperature. The CO2 sorption capacity as-measured with the as-received dolomite is normalized on the basis of MgO+CaO content, i.e., on the same basis as the sorbent derived from the calcined dolomite. No obvious advantage can be seen by using the pre-calcined dolomite. Comparative pairs, sorbent 1 vs. 3, 2 vs. 4, and 7 vs. 6, show no obvious advantage by increasing the promoter loading from 5 to 10 wt. % with the same starting material and same calcination temperature. Comparative pairs of sorbents 3 vs. 5, 4 vs. 6, and 2 vs. 7 reveal the impact of the composite sorbent calcination temperature with the same dolomite precursor and same promoter loading level. 725° C.-calcination does not show any significant difference from 350° C.-calcination. This group of experiments indicates that the composite sorbent preparation is not so sensitive to the starting material, calcination temperature, and promoter loading. A working sorbent structure is formed through repeated sorption/regeneration test cycles. The promoter is used as an additive to decorate the solid surface and intended not to exist as a separate phase. Thus, its excessive loading is not necessary or may become a negative factor by blocking accessibility of CO2 into the sorbent pores.
A packed bed sorption test was conducted on a tube tubular reactor (316 stainless steel, 0.5″ OD, 0.049″ wall thickness). The sorbent particles of 40-100 mesh were loaded into the reactor tube by tapping the reactor wall. The resulting packing density was about 0.6 g/cc. The sorbent bed was held by quartz wool on the top and bottom. The gas flow was directed downward through the sorbent bed. Individual gas flow rates were controlled by mass flow controllers (MKS Instruments model 1179A and 1479A), while water flow rate was controlled by HPLC pump (Cole Parmer model DigiSense 89000). The bed temperature was controlled by the furnace heating. The sorbent bed was heated to 700° C. in 0.5 liter/min air to activate prior to lowering to designated sorption temperatures. The water was preheated to vaporize at 150° C. prior to entering the reactor. The reactor effluent was cooled down to 4° C. to condense the majority of the entrained water vapor. The residual gas was analyzed by a quadruple mass spectrometer (Kurt J. Lesker model Quad 200). Typically, 0.5 liter/min dry gas consisting of 33% CO2 and 67% H2 was used to simulate the reforming gas.
To elucidate effects of sorption conditions on the sorbent performances, we conducted packed bed tests with the sorbent #1.
The present composite sorbent exhibits very rapid sorption kinetics. The CO2 sorption on the packed bed is nearly completed within 2 min for all the runs. It is estimated from the sorption breakthrough curve that the first order CO2 sorption rate constant is about 9.4 l/s for the un-saturated sorbent and about 4.0 l/s for the 95% saturated. The rapid sorption is necessary to reduce the reactor size and to combine the CO2 sorption with steam reforming reactions in situ. The CO2 sorption rate obtained in this work is comparable to CO2 adsorption into porous adsorbents. The former is a gas/solid carbonation reaction process, while the latter is mostly a physical adsorption process. The rapid sorption kinetics is attributed to catalyzing functions of the promoter. CaO has a higher density than CaCO3. It is known that solid diffusion often is a rate-limiting step for carbonation reactions. The liquid-phase promoter and presence of H2O may help breakup of CaO crystals into highly-porous structures for rapid CO2 diffusion and reaction.
The same sorbent bed was tested under various conditions over a period of 17 days, while the regeneration conditions were kept the same (700° C. in air). It was noticed that the CO2 sorption capacity increased with the time as shown in
The dolomite sorbents prepared and tested in the previous examples were analyzed by electron scanning electromicroscopy (SEM) to check the micro-structures and compositions. As shown by SEM pictures in
In addition to atomic compositions, the elemental distribution is very important to obtain the synergistic effects of the composite sorbent. Elemental mapping was conducted during SEM analysis. Maps of Mg, Ca, Na and K elements for one fresh and two spent sorbents are presented in
The crystal phases of the composite sorbent were measured by X-ray diffraction (XRD).
After the sorbent was tested under various conditions and unloaded from a fully-activated state, MgO crystal phase was intact (
This example shows that sorbents without a sufficient amount of eutectic carbonate promoter is not active for CO2 sorption. Dolomite from City Chemical was calcined at 850° C. for 4 hours in air. 90 g of 850° C. calcined dolomite (MgO—CaO) was mixed with 3.22 g of Li2CO3 (99%, Sigma Aldrich), 3.33 g of Na2CO3 (99%, Sigma Aldrich) and 3.45 g of K2CO3 (99%). The solid mixtures were added into a 500 ml Nalgene plastic bottle and mixed with 250 g of 2-propanol (EMD Chemicals, Canada) and 500 g of zirconia beads (diameter: 0.3-1 cm). The bottle was rotated for 18-96 hours at a speed of 150-200 rpm. The obtained slurry was dried at RT ° C. in plates to evaporate 2-propanol and form 0.2-0.5 mm thick cake. The cake was crushed and dried at 120 C. Then the sieved particles (<100 mesh) were calcined 725° C. for 3 hours in air.
Both dried and pre-calcined sorbent particles were tested for CO2 sorption in the packed bed at 500° C. sorption temperature in CO2 and 725° C. regeneration in air. The sorbents showed initial CO2 sorption activity but quickly de-activated in the subsequent test cycles. The samples were analyzed to understand the problem. The atomic compositions of as-prepared and tested sorbents are listed in the Table 7 below.
Small grains and porous structures (
This example confirms the necessity to have an eutectic carbonate phase be uniformly distributed in an active composite sorbent with adequate amounts, such as, K and Na ratios to Mg greater than 0.020.
A TiO2 monolith (Cormetech, NC) was used to make the monolith-structured sorption bed.
As shown in
First, alternative rows of channels on the bottom monolith were plugged by cement to hold the sorbent powder. Then, adjacent rows of the channels on the bottom monolith were plugged by cement to prevent the sorbent powder from filling all the channels. The dolomite sorbent was crushed and sieved to the sizes well below the channel opening (1.0 mm), such as 300 μm. The pugged cement was held inside a tube that was mounted onto a vibrator or shaker in vertical orientation. The sorbent powder was poured onto top of the monolith, while the tube was shaken or vibrated. The powder may be added gradually as the channels were being filled. Filling was stopped after the channel was full.
Nine pieces of the monoliths were filled with the composite sorbent powder and associated weight gains are listed in the Table 8 below. The average packing density of the sorbent powder inside the channel is 0.68 g/cc with standard deviation of 0.105 g/cc by assuming 1.0 mm channel width and 0.5 cm plugged channel length.
The NiCuMgCeCeAl—O metal oxide composite catalysts were first prepared as particles to characterize their properties. The nitrate salts used in the preparation were Ni(NO3)2.6H2O 100.0%, Cu(NO3)2.2.5H2O 98%, Mg(NO3)2.6H2O 99%, Ce(NO3)3.6H2O 99%, and Al(NO3)3.9H2O 98%. The nitrate salts were weighed according to required stoichiometric ratios and mixed. De-ionized water was added to obtain a clear solution. Then, a mixture of NaOH (0.88 mmol/g) and Na2CO3 (0.88 mol/g) solution was prepared. The mixed metal nitrates solution was heated in a water bath on a hot plate to 75-80° C., the NaOH/Na2CO3 solution was added drop-wise with stirring to pH=9-10. The temperature was controlled at 75-79° C. For sample 61880-49, about 89 g of NaOH (0.88 mmol/g)/Na2CO3 (0.88 mol/g) solution was added to 85.5 g of metal nitrates solution in 20 min. The mixture was stirred for 1 hr at 75-78 C for 1 hr. The solution compositions used for preparation of the composite catalysts are summarized in table 9. The precipitate was aged for 10 h at RT, centrifuged and washed with DI water for 4 times until pH=7. The precipitate was dried overnight at 40-50° C. and then 120 C for 28 hr. The dried precipitates was heated at 1° C./min to 500° C. and held at 500° C. for 6 h in air.
The BET surface areas and pore sizes of four composite catalysts are listed in Table 10. The first four samples were prepared in the same batch at about 10 g scale. BET surface area and pore volume decreased with increasing Al content in the composite, while the pore size increased. The Ni:Cu:Mg:Ce:Al=1:0.2:0.1:1:2 formulation was chosen for parametric studies of the preparation process. A larger batch of the particle (100 g scale) was prepared using this formulation with quick precipitation and 2 times of water washing. With the sample calcination condition, the resulting catalyst has BET surface area of 140 m2/g, slightly less than the small-scale trial. The catalyst was further calcined at 725° C. for 6 hours.
The same TiO2 monolith as used in the previous sorbent loading was used to make the monolithic reforming catalyst by impregnation. Table 11 lists the four solutions used for impregnation. The impregnation solution was prepared using the same metal precursors as used in the above example. The preparation procedure is described in the following. The 2-cm TiO2 monolith was fully immersed in the solution at room temperature (20° C.) and soaked a few minutes under vacuum until no gas bubbles emerged out of the solution. The monolith was taken out of the solution. The excessive solution was drained first and then spins out by centrifugation. The wet sample was left in the fume hood to dry overnight and further dried in an oven at 120° C. The dried sample was calcined at 500° C. for 6 h with 1° C./min ramp rate in air. The impregnation was repeated to increase the catalyst loading.
After the calcination, the monolith body was cut open to expose the interior channels. All the channels looked uniform in color and were of any slid deposit on the channel wall. Thus, the metal oxide reforming catalyst was deposited into the TiO2 support pores. The monolith was crushed into particles for BET surface area and pore volume measurements. The BET surface area and pore volume decreased with increasing the catalyst loading. This is explained by increased occupancy of the pores on the monolith channel by the catalyst.
With the above success, impregnating TiO2 monoliths was tried with more complex solutions. Different from the above procedure, the impregnation was conducted with hot solution and monolith to increase the catalyst loading. One impregnation solution was prepared by mixing the Ni(NO3)2.6H2O, Cu(NO3)2.2.5H2O, Mg(NO3)2.6H2O, Ce(NO3)3.6H2O, and Al(NO3)3.9H2O salts with an atomic ratio of Ni:Cu:Mg:Ce:Al=1:0.2:0.1:1:2. The weight concentration is shown in table 12. A minimal amount of water was added into the salt mixture, heated, and stirred at 60° C. until complete dissolution. A clear, green solution was obtained and denoted Sol 61880-72. Another solution was prepared by adding KNO3 salt in the above mixture with an atomic ratio of K/M (M=Ni+Cu+Mg+Ce+Al)=0.08, which is denoted as Sol 61880-72-K.
Bare TiO2 monoliths were placed in 75° C. oven to warm up prior to impregnation. The impregnation was conducted by immersing the warm monolith in the 70° C.-solution and degassing under vacuum for 5-10 min until no major gas bubbles emerged. Then, the monolith was taken out of the solution. The excessive solution inside the channel was drained first and then, removed on a centrifuge at 3000 rpm for 2-5 minutes. The wet monolith was left in the fume hood to dry at room temperature, and dried inside an oven at 120° C. overnight. After drying, two monoliths were impregnated second time with respective solution. All the dried monoliths were calcined in a tubular furnace in air at 500° C. for 6 h. Table 13 lists the four monolith catalyst samples prepared.
One-time impregnation resulted in the 16.6 and 18.9% catalyst weight gains with the 61880-72-K and 61880-72 solutions, respectively. Upon autopsy, the channels appeared uniform.
The elemental compositions and uniformity of the monolith catalysts were analyzed by SEM/EDS. Two SEM samples were prepared for each of the first three monoliths in Table 13. First sample has clean, fractured channel walls, while second sample has the flat channel wall being exposed. Three spots along the channel wall thickness were analyzed: two close to the wall surface, one in middle of the wall thickness. The elemental compositions of interior channel walls are compared to the exterior channel wall in Table 14. Presence of all the metallic elements inside the channel wall is confirmed. There are some variations in relative metal element loading at different spots at micrometer-scale. But, co-presence of targeted elements throughout the channel wall is confirmed.
A new group of monolith catalysts listed in Table 15 were prepared using the NiCuMgCeAl solutions and procedures same as used above. With the repeated procedures, the resulting catalysts showed consistent weight gains. First impregnation resulted in catalyst weight gain 13.2-14.5 wt % among six samples. Different from the above trial with NiCuMgCeAl composite solution, second time impregnation was conducted on the calcined monolith rather than the dried. The second impregnation resulted in final catalyst weight gain of 25.5-30.0 wt % among the four samples.
Preparation of the monolith catalyst by coating was tried with coating slurry made of NiCuMgCeAl—O composite particle, which was prepared using the previous particle preparation procedure. The coating slurry was introduced into the bare TiO2 monolith channel and drained. The excessive solution was further removed from the channel by centrifugation. Resulting coating looked uniform for the three samples. The coated samples were dried in fume hood overnight and then, dried in an oven at 120° C. for 5 days. The dried monolith was calcined at 500° C. for 6 h in air in a tubular furnace. Table 16 lists the weight gains. The catalyst loading was only a few percentage, which is fairly low. More problematically, the coating layer did not have any adhesion on the channel wall, and existed as a loose layer. This example shows that impregnation is a more effective method to load the metal oxide composite catalyst into monoliths.
Bio-oil feed used in this testing was an oak pyrolysis oil (TCPDU) received from NREL on Oct. 29, 2010. The oil remains fluidic at room temperature but was very viscous. Its composition was reported as 44.94% C, 7.29% H, 47.66% 0, 0.01% S, and <0.05% ash. The molecular formula is CH1.95O0.80.
2.0 g of catalyst particles were packed into a ½″ reactor tube with a bed packing height about 3.2 cm. The catalyst names and testing results of five reactor runs are summarized in Table 17. NiCu—MgCe—Al oxide composite was prepared in the catalyst particle form (Cat ID 61880-35 in Table 10). NiCuCe/TiO2 (as-received), Ni/TiO2 (700° C.-calcined), and Ni/TiO2 (H2O-washed and 700° C.-calcined) were prepared in the monolith form (Cat #61880-34, 61880-25-1 in Table 11). The Ni/TiO2 (700° C.-calcined) and Ni/TiO2 (H2O-washed and 700° C.-calcined) were prepared using the same procedure as Ni/TiO2 (61880-25-1 in Table 11) except for using of different TiO2 monolith. The 700° C.-calcined and H2O-washed monolith supports were used to remove potential harmful components in the as-received monolith. The monolith was crushed and sieved into 20-100 mesh, same sizes as the catalyst particle.
For first two runs, top of the catalyst bed was packaged with a layer of quartz wool and a monolith spacer. In 3rd run, the monolith spacer was removed. In runs 4 and 5, nothing was packed on top of the catalyst bed. The reactor was heated to a mid-bed temperature of 400° C. Steam was introduced first and followed with bio-oil delivered by a syringe pump. The reactor effluent was cooled, and the steam and oil were condensed into a gas/liquid separator. The off-gas flow rate was measured by a flow meter and the gas composition was analyzed by a Mass Spec using argon as an internal standard. The gas production rate is calculated based on the off-gas flow rate and composition. All the runs were conducted near atmospheric pressure at GHSV=100,000. After the reaction was held for 10 to 40 min, the reactor was heated to either 450° C. or 500° C. Since the gas production varied with time, initial gas production rates are used for comparison of different reactor runs in Table 17.
Compared to the TiO2 monolith-supported catalysts, the NiCu—MgCe—Al oxide composite catalyst showed a significant reforming activity at 400° C. Thus, this catalyst formulation was used to prepare monolith catalysts.
The NiCu—MgCe—Al oxide composite catalyst before and after the bio-oil reforming test was analyzed by SEM. The elemental compositions are listed in Table 18. The micro-structures and EDS analytical spots are shown in
The TiO2 monolith impregnated with the K—NiCuCeMgAl composite catalyst (Cat 60195-145-3 in table 15) was tested for steam reforming of bio-oil. To mitigate the small flow rate problems with the reacting testing apparatus, ¼″ quartz tube of 6.2 mm OD×4.2 mm ID was mounted onto top of the monolith by cement. The monolith module was housed inside a steel reactor tube of 20 mm ID, and steam/bio-oil line was connected to the quartz tube. In this way, the steam/bio-oil feed was directly delivered to the four reaction channels of the monolith module. The as-received bio-oil was a viscous, dark fluid. The bio-oil was pumped into the reactor at room temperature, while water was vaporized to abut 110° C. into steam. The steam combined with the bio-oil flowed downwards into the monolith channel via the quartz tube. The steel reactor tube was heated by a two-stage furnace with the top and bottom bed temperature being monitored and controlled.
The reactor temperature was raised to about 560° C. in 200 sccm of air. When the reactor temperature stabilized, 200 sccm of Ar purge gas was introduced into the reactor with 0.2 cc/min water and 0.2 cc/min of bio-oil. The reactor pressure was controlled at about 5 psig. The reactor effluent was cooled to about 6° C. The liquid was condensed and collected, while the gas was sampled for composition analysis by Mass Spec. After the steam reforming reaction was carried out about 20 min, the bio-oil and water feeds were stopped and 200 sccm air flow was introduced to conduct the catalyst regeneration. The regeneration was carried out about 20 min.
The liquid products collected from all the reforming runs looked clear. These products are a mixture of water and un-converted (or reformed) bio-oil. A clear solution is very beneficial to re-processing of the remaining hydrocarbon species, and preferred embodiments of the invention can be characterized by the recovery of a clear liquid phase as the primary or only recovered liquid phase and, optionally, recycle of all or part of the liquid phase.
This patent application claims the priority benefit of U.S. Patent Application Ser. No. 62/109,840.
| Number | Date | Country | |
|---|---|---|---|
| 62109840 | Jan 2015 | US |
