US 20090071064 A1
Embodiments of the present invention concern methods, compositions, and apparatus for the continuous conversion of algal lipids into biodiesel. In some embodiments, the biodiesel is formed in a multi-step sequence, the first steps occurring in the presence of water and a strong acid wherein the lipids are released from the algae by means of mechanical and chemical action and are then hydrolyzed to free fatty acids. In a subsequent step, this free fatty acid mixture is reacted with methanol to generate fatty acid methyl esters (also known as biodiesel). Such methods produce biodiesel from algal lipids without the requirement for separate algal cell lysis or lipid extraction or purification prior to the acid catalysis sequence. In other embodiments, the multi-step acid catalysis sequence occurs at 100° C. at two atmospheres of pressure.
1. A method for continuous production of biodiesel from algae comprising:
a. continuously feeding an aqueous suspension comprising algae into a biodiesel production plant; and
b. converting lipids from the algae into biodiesel without an initial purification or extraction step.
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This application claims the benefit of U.S. Provisional Patent Application Ser. No. 60/952,443, filed on Jul. 27, 2007, and entitled, “Continuous Algal Biodiesel Production Facility,” which is incorporated by reference herein in its entirety for all purposes.
Embodiments of the present invention concern methods, compositions and apparatus for continuous algal biodiesel production. In particular, methods are provided for continuous biodiesel production from algal biomass, without the need for extraction or separation of algal lipids for esterification into biodiesel. More particularly, such embodiments concern an in-situ hydrolysis and esterification process for production of fatty acid methyl esters (FAME), commonly referred to as biodiesel.
2. Discussion of Related Art
Reducing the reliance of the United States on imported fossil fuels is of concern for the future. Two-thirds of the oil consumed in the U.S. is imported, and that number will continue to rise unless reasonably priced, domestically produced alternatives can be created. At the same time, the long-term supply of oil is limited, and a large percentage of proven oil reserves are in countries with unstable or hostile governments. Cutoffs in foreign oil supply have been used as a political tool in attempts to coerce U.S. policy. Combustion of petroleum based products releases enormous amounts of CO2 into the atmosphere, contributing to global warming. Thus, foreign petroleum use is a national concern, as well as an environmental concern.
Various attempts have been made to develop biofuels from non-petroleum sources. A large scale effort has been made to develop ethanol from plant materials, primarily from corn grain. Although substantial progress has been made, with almost five billion gallons of corn ethanol produced in 2006, the resulting impact on corn and food prices suggests that there are limits to how much further production is feasible. Present corn ethanol production is only sufficient to provide about 1.5% of U.S. fuel needs and consumes over 20% of total U.S. corn production.
Other technologies have been developed to produce biodiesel from plant sources. Many different irrigated crops, such as soybean, rapeseed, jathropa, palm and sunflower, are currently being used to produce biodiesel. Current biodiesel production often utilize some form of transesterification process, wherein triglycerides or other starting materials undergo an alkali or acid catalyzed transesterification reaction between the fatty acid component of the triglyceride and a low molecular weight alcohol, such as methanol. Glycerol is released as a byproduct of transesterification and fatty acid methyl esters are produced. Such processes may be operated in either a batch or continuous mode. However, it is currently necessary to first separate the triglycerides or other source material from the bulk plant matter before the transesterification reaction can proceed. Using such technologies, current biodiesel production is only a small fraction of corn ethanol production, with present U.S. capacity of less than 100 million gallons per year. In addition, biodiesel production from food crops, such as soybeans, will ultimately encounter the same problems that currently limit corn ethanol production.
Alternatives to increase biofuels production capacity have been proposed, such as conversion to cellulosic ethanol production, utilizing wood, switchgrass or other non-food starting materials. However, cellulosic ethanol technology, while promising, has not yet been developed to the point of full commercial scale production and the time required to reach that point remains uncertain. Other proposals have involved biofuel crop production on marginal or idle land, such as the Conservation Reserve Program (CRP) acreage. Such proposals ignore the practical difficulties of obtaining water supplies to grow such crops, requirements for fertilizer input, low productivity of marginal land, etc.
Another alternative source of biofuels production has been proposed for algal culture systems. The National Renewable Energy Laboratory (NREL) in Golden, Colorado spent 10 years and $25 million on an Aquatic Species program that focused on extracting biodiesel from unusually productive species of algae. Before losing funding, the government scientists demonstrated oil production rates two hundred times greater per acre than achievable with fuel production from soybean farming. However, the open pond system utilized by NREL was susceptible to invasion by contaminating algae, bacteria or algal-consuming organisms and algal productivity was adversely impacted by fluctuating environmental temperature and solar radiation. Further, in a pond type of system the light penetration depth into dense algal cultures results in only a limited band of photosynthetic productivity, with the majority of algae being shaded by overlying organisms.
Closed system bioreactors for algal culture have been proposed to provide better control of algal growth conditions, limit invasion by undesired species, and enhance algal growth by supplementing environmental gases with, for example, CO2 emissions from power plants or other fixed carbon sources. While closed systems represent an improvement over open system alternatives, such designs are often expensive to construct and maintain, poorly scalable, and not optimized for maximal algal lipid production under continuous culture conditions. Thus, a need exists in the field for closed system, inexpensive, scalable bioreactors capable of growing algae and producing biodiesel or other products, and methods of use of such bioreactors for continuous production of biodiesel from algae. A further need exists for methods of biodiesel production from algae that does not require separation or extraction of lipids from algal cells, prior to conversion of algal lipids into biodiesel.
Embodiments of the present invention include methods, compositions and apparatus for continuous production of FAME from algae, without the need for purification and/or extraction of lipids from algal cells. Certain embodiments concern methods involving a novel in-situ hydrolysis-esterification reaction process. Lipids, or triglycerides, are first hydrolyzed to create fatty acids at a 99 percent conversion rate. These are then esterified to fatty acid methyl esters (FAME), commonly referred to as biodiesel.
In some embodiments, in order to drive this reaction to completion, a multiplicity of continuously stirred tank reactors are used and are split into three parallel streams. Following the reaction section, continuous centrifuge and decanter units may be used to separate the solid, liquid and oil phases. As methanol is non-renewable and somewhat expensive, about 98 percent of the feed methanol may be recycled. Additional washing and reaction steps may be used to create FAME with 99 percent purity.
Various embodiments concern use of one or more closed system bioreactors for algal culture, to provide starting material for the methods of continuous biodiesel production. Additional details of such closed system bioreactors for algal culture are disclosed in U.S. patent application Ser. Nos. 11/510,148 and 11/510,442; PCT Patent Application Ser. No. PCT/US2006/033252; and provisional U.S. Patent Application Ser. Nos. 60/894,082, filed Mar. 9, 2007; 60/877,907, filed Dec. 20, 1996, and 60/878,506, filed Jan. 3, 2007; the text of each of which is incorporated herein by reference.
Certain embodiments may concern methods, compositions and apparatus for fermentation of glycerin or glycerol into ethanol. In some embodiments, the fermentation may be accomplished using bacterial fermenters, such as E. coli.
While the invention is amenable to various modifications and alternative forms, specific embodiments have been shown by way of example in the drawings and are described in detail below. The intention, however, is not to limit the invention to the particular embodiments described. On the contrary, the invention is intended to cover all modifications, equivalents, and alternatives falling within the scope of the invention as defined by the appended claims.
Biofuel Production from Algal Culture
Algae are a very diverse and simple group of aquatic plant that are widespread across the world. Algae can vary in form from Eukaryote to Bacteria, and are spread across the kingdoms Plantae, Protista, and Protozoa. All forms contain biomass, which can be converted to various renewable fuels. In some embodiments, the types used for culture are photosynthetic Plantae algae, although the skilled artisan will realize that alternative algal types may be utilized in the practice of the disclosed methods. The algae are selected for their high lipid content and efficient growth under a variety of conditions.
Algae is a beneficial feed stock because it has the potential to provide over one hundred twenty times the fuel output (per acre) of soybeans, the primary crop used for present biodiesel production. Even palm oil, which has gained popularity recently, has a lower order of magnitude for production per acre. Algae farms are also advantageous in their ability to provide growth within a contained environment, drastically reducing water usage compared to conventional agriculture. Furthermore, livestock feed and general food supply (a current concern for corn based ethanol, soybean oil, and other biofuels ) will not be impacted by such a system.
An optimal location to grow algae quickly is next to a coal or natural gas fired power plant, where large amounts of CO2 are released through combustion of fuel. The carbon dioxide, a greenhouse gas, can be sequestered by the algae and some of it can be converted into the carbon-containing lipids. Additional biomass may be created, resulting in a reduction of atmospheric carbon dioxide, even after the biodiesel is burned in a combustion engine. At the same time, the CO2 enrichment provides for enhanced algal growth.
Production of Biodiesel from Triglycerides
Certain embodiments of the present invention concern a continuous process for converting algae to biodiesel (fatty acid methyl esters-FAME). Current methods for production of biodiesel from oil primarily utilize batch processes using a base catalyzed reaction in a dry environment. Harvested algae contain a relatively high percentage of water which obviates use of the current technology in either the batch or continuous mode (soap, not biodiesel, would be produced).
According to some embodiments of the present invention, a process utilizes a two step water based acid catalyzed reaction system to convert the algae lipids to biodiesel. At least one novel feature of such a process is that the conversion of the lipids (embedded in the algae) to fatty acids in the first stage reactors proceeds without the prior extraction of the lipids. Extraction of lipids from algae can be problematic and inevitably introduces at least one additional chemical component which requires subsequent separation from the product, by-products, and other feed chemicals.
Another novel feature is the use of a water-based two stage reaction system to produce biodiesel. Transesterification of the lipids to FAME, the current route to biodiesel, proceeds exceedingly slowly when acid catalyzed. However, if the lipids are converted first to free fatty acids which are then reacted with methanol and an acid catalyst, both reactions occur rapidly and to a high extent. Further, after the biodiesel has been produced in the two stage reactor system, it can be readily, energy-efficiently, and economically separated from the other chemicals in the reactor effluent using equipment common in the chemical industry. All unreacted methanol is recovered and recycled.
The manner in which the by-product, glycerol, is upgraded is also new. Huge quantities of biodiesel will need to be produced to have a significant effect on oil imports. The amount of glycerol accompanying this biodiesel would flood the market, driving prices down dramatically and thus resulting in an almost worthless by-product. Embodiments of the present invention take advantage of a new genetically-engineered strain of E-coli to convert the glycerol primarily to ethanol, which can be sold as automobile fuel. The ethanol itself is purified in the process by employing novel ceramic/molecular sieve membrane technology. This drastically reduces the energy requirement for this step of the process by eliminating the need for high purity distillation.
The residual cell matter from the algae is energy-rich and is burned to generate most of the steam and electricity needed for the overall process (the carbon dioxide from this combustion can be captured and used to grow additional algae), according to embodiments of the present invention. The overall energy balance for the process is exceptionally favorable, estimated at about 6 units of energy produced for every unit consumed, according to embodiments of the present invention. For comparison, ethanol plants are at an overall energy balance of approximately 1.2.
Existing biodiesel technology uses oils from various plants, such as soybeans, as the feedstock. This oil is naturally free of water so a base catalyzed transesterification reaction is feasible. Industrially, the reaction is promoted by the conjugate base of the alcohol (methanol), as shown below:
where the methoxide ion provides a strong nucleophile for the reaction mechanism. The methoxide is generated by dissolving a small amount of base (e. g., potassium hydroxide) in the methanol. The transesterification reaction is reversible and so is carried out in an excess of methanol. The absence of water prevents the formation of soap.
The conversion of algae-based lipids into FAME presents a unique challenge with respect to existing technologies because the algae itself contains water as part of its biological makeup and also due to the harvested algae containing residual, water-based growth media. Consequently, the base promoted technology employed today by the biodiesel industry will not work. As described below, embodiments of the current invention solve this problem. The release of the lipids from the algae takes place in the first reactor, avoiding an expensive extraction step.
Thus, while the lipids are being freed from the somewhat fragile algae structure by means of vigorous mixing in the presence of a strong acid, the first step in the transformation of these lipids to FAME proceeds by means of their acid catalyzed hydrolysis to free fatty acids. For example, using triolein, the major oil component occurring in many plants, as a model lipid, the triolein would be converted to oleic acid as follows:
The general reaction mechanism for this is shown in
The reaction of the oleic acid with methanol to form biodiesel is carried out in a second reactor, according to embodiments of the present invention, as follows:
This type of reaction has been shown by others to occur rapidly at seventy degrees centigrade using an excess of methanol in the presence of a strong acid, with 99% conversion within a one hour residence time. Others have also demonstrated that the acid catalyzed transesterification of lipids directly to biodiesel is very slow (36 hours to achieve 80% conversion) and would be commercially impractical. So the multi-step sequence described above (algae release>Equation 2>Equation 3) significantly improves a design of an algae-to-biodiesel process.
Currently, there is no continuous technology of any kind commercially producing biodiesel. For an algae feedstock, processes according to embodiments of the present invention could readily provide for this because their design is based on standard chemical equipment commonly used for continuous processes, according to embodiments of the present invention. Such processes should therefore be scalable to very large capacities while producing biodiesel at an economical price. Acid catalysis was selected for the present methods. In some embodiments, the esterification takes place at 100° C. and at roughly twice atmospheric pressure. The mechanism of this acid catalyzed reaction is shown as
A difference in this synthesis is that a combination of in-vitro hydrolysis and esterification reactions are utilized. The use of a hydrolysis reaction, which converts the triglycerides to free fatty acids, increases the overall rate of conversion to the product. Free fatty acids produced from the hydrolysis readily undergo acid catalyzed esterification with a residence time of about one hour with ninety-nine percent conversion, according to embodiments of the present invention. This can be contrasted with a direct acid catalyzed transesterification of triglycerides, which requires about thirty-six hours to achieve 80 percent conversion. Such a two-step, acid-catalyzed, in vitro reaction sequence drastically simplifies the process, as the lipids do not need to be extracted from the cells, and they can be quickly converted to product. The acid not only catalyzes the esterification reaction but also breaks down the algal cell walls and releases the lipids for conversion. This is considerably less expensive than methods utilizing an extraction step prior to esterification.
Certain embodiments concern design of commercial scale plants that will allow continuous production of 100 million gallons per year of biodiesel from algae. As discussed above, in-situ acid-catalyzed hydrolysis followed by esterification is used as the reaction for the conversion of algae biomass to biodiesel. The design focuses on the downstream processes that will convert an algae stream into biodiesel, according to embodiments of the present invention.
In determining efficient design of a biodiesel production facility, certain assumptions were made according to embodiments of the present invention:
The full process flow diagram (
according to embodiments of the present invention, there are three main feeds to the plant. The first is the algae feed. Its composition is 25 percent water, 30 percent triglycerides (as triolene), and the cell matter which accounts for the other 45 percent. The cell matter contains carbohydrates (as energy storage or as cell wall material), proteins, and other cell mass. Additionally the phospholipids that make up each cell membrane are potential feed stocks for the reaction, but in the treatment below are assumed to be accounted for as part of the lipid content. This algae slurry enters the system at approximately 27° C. and 1 atmosphere. Prior to entering the first reactor sequence, it is combined with the second stream, sulfuric acid at 18 M. This second stream arrives at the same temperature and pressure. These streams are mixed before being split and sent to the three individual reaction chains.
A methanol feed stream for the Fisher Esterification section of the reaction scheme is an additional inlet and is fed at 27° C. and 1 atm.
There is an additional sulfuric acid feed stream entering at the same conditions that is utilized to feed the secondary esterification reaction. It is mixed with a methanol stream.
The reaction scheme is run in parallel to reduce reactor size and individual flow rates, while maintaining the overall flow rates. Prior to reaction, the streams are pressurized by pump P-101 (P-201 and P-301 in the parallel chains) up to 2 atm. This keeps the methanol liquid at the reaction temperature. The pressurized stream is then heated by H-101 (H-201, H-301) to 100° C., the preferred hydrolysis temperature.
The first section of the reaction scheme focuses on the triglyceride hydrolysis. Under appropriate conditions, three water molecules will combine with each triglyceride to form three stable fatty acids and glycerol. Reactors R-101 and R-102 combine to provide the one hour residence time required for ninety-nine percent hydrolysis. They are operated at 100° C. and 2 atm. The output of the reactors reflects the ninety-nine percent conversion of triglycerides to free fatty acids.
The second section of the reaction scheme converts the free fatty acids to form the biodiesel end product. A methanol stream is added, fed at two times the required stoichiometric amount to drive the reaction further to completion. Part of this stream consists of the methanol recycle stream. The combined stream will be around 65° C., but only makes up about six percent of the total mass flow. Therefore, no additional heater is included, as the reactors are jacketed, and the reactions themselves are exothermic. Reactors R-103 and R-104 combine to allow for a one hour residence time at 100° C. and 2 atm, resulting in an 85 percent conversion of the free fatty acids to methyl esters (biodiesel).
The stream leaving the reaction section is then sent to the separation section. It is first centrifuged in C-101 to remove the algae cell matter. The supernatant includes all liquid components that left the reaction section. The solids are removed at a rate such that the mass flow rate downstream is reduced by almost 50 percent. The liquid stream enters a decanter, which allows for a phase separation with a residence time of two hours. Water, glycerol, acid, and methanol leave in the heavy phase and are sent to the methanol purification section of the plant. The light phase, consisting of fatty acid methyl esters and free fatty acids, is sent to the secondary reaction section of the plant.
The heavy streams from the decanters are combined and preheated for flashing by H-401. This stream is then flashed, resulting in a separation of the glycerol and acid from the water and methanol. Glycerol and acid leave the flash drum as a liquid bottoms stream, which is then sent on for further processing. The vapor overhead stream is combined with the aqueous phase from the secondary esterification reaction, for a total flow into tower T-101. The distillate leaving the tower consists of 99.1 percent methanol, which leaves a bottoms product of mainly water and trace sulfuric acid and glycerol. The methanol distillate stream is recycled back to the esterification portion of the reaction sequence, allowing for a significant reduction in the required methanol input.
The product stream of free fatty acids and fatty acid methyl esters is sent to the biodiesel purification. Acid and methanol is added to reactor R-401 along with the product stream and allowed to react for one hour. The conversion approaches 99 percent and the outlet consists of 96.3 percent methyl ester. This stream is washed in M-101 by water in order to create a two phase system. It then is phase separated in D-401, resulting in a final product stream of 99 percent biodiesel, and a waste stream of methanol and water that goes to T-101.
A large scale view of the overall process flow diagram is presented in
The energy consumption of the plant can be a concern during the development of an overall process flow diagram, according to embodiments of the present invention. In a plant, it may be beneficial to recycle as much of the feed stocks as possible, including methanol and water. Furthermore, it may be advantageous to utilize as many byproducts as possible in order to offset the costs of production. A method used to separate the product biodiesel from the byproducts may be chosen with lower energy requirements. Byproducts may be used in a manner that benefits the overall energy balance and cost of the process. For example, the glycerol produced in the esterification reaction may be processed downstream to produce ethanol, which can either be sold or used as fuel. Methodologies according to embodiments of the present invention can be shown to produce significantly more energy in the form of biodiesel than was required by the process to produce the biodiesel.
Calculation of an overall energy requirement for a particular factory design takes into account each individual process component, according to embodiments of the present invention. However, it may be assumed that the decanters, pumps and centrifuges are perfectly insulated; that is, there was no heat lost from these instruments. Following this method, it was found that this process design according to embodiments of the present invention has an overall energy requirement of 306 MBtu/hr. This energy is used for the heating of incoming streams, preparing a stream for flash, and distillation. Additionally, energy is needed to keep the reactors at the target temperatures. Once the initial feed stream is broken into three streams each one goes through a heat exchanger to reach the target 100° C. Each heat exchanger has a heat duty of 4.4 MBtu/hr. The esterification reactors experience an exothermic reaction, which may require cooling to avoid thermal runaway. Therefore jacketed reactors may be chosen, but cooling water may be useful in such cases. Cooling water is usually a minimal cost in plant design.
In order to separate the water and other by products without using an extremely large distillation column, a flash drum is situated in front of the distillation column. The product stream is heated to 125° C. before entering the flash drum. This heat exchanger has a heat duty of 77 MBtu/hr. The flash drum has a heat duty of 6.5 MBtu/hr in order to keep it at 125° C. The distillation column used to purify methanol has a heat duty of 198 MBtu/hr. The reactors and smaller heat exchangers make up the balance of the total heat duty.
The heat for this process is mainly supplied by steam at 100 psi, which makes up most of the annual utility cost. The 100 psi steam may be used for the feed heat exchangers, distillation column, and the reactors. For the pre-flash heat exchanger, 450 psi steam may be utilized. Additionally, cooling water may be used in a number of applications. It may be used to condense the methanol leaving the distillation column. More cooling water may be necessary for the reactors as mentioned before, as well as cooling any product streams that may require further processing on site. Table 1 shows the costs and amount of each utility that may be used, according to embodiments of the present invention. Electricity is used to power the pumps and centrifuges. Process water is calculated as the overall requirement, and could be potentially supplemented by a recycle later on in the system, but the full requirement is only 0.1 percent of the total annual cost and would not impact the pricing in any significant way, according to embodiments of the present invention.
A gravity separator may be used instead of a distillation column or other energy consuming unit, to conserve energy. The biodiesel and water exist in different phases and will separate if given enough time, according to embodiments of the present invention. Decanters help to not only separate the water from the biodiesel, but since the acid, methanol, and glycerol are miscible in water they are also removed from the product stream. This leaves only biodiesel and unreacted free fatty acids in the product stream, and allows for continuous plant operation.
In order to get biodiesel of 99 percent purity, the biodiesel and unreacted free fatty acids may be run through another reactor which converts almost all the free fatty acids to biodiesel. A decanter is once again used to remove the added methanol and sulfuric acid. To make the product as pure as possible it is then washed with water to remove any trace amounts of acid, methanol, and glycerol. If the concentration of free fatty acids in the final product is too high, residence time in the reaction can increase to reduce FFA levels.
Table 2 shows that the disclosed process has a positive energy balance, according to embodiments of the present invention. The amount of sun used to make algae grow was not included in the energy balance. The positive energy balance is an important consideration in the design and construction of a plant, in comparison to alternative methods. Concerns have been expressed about the net energy produced in the production of ethanol and rapeseed derived biofuels. The energy balance for the disclosed plant shows that 5.6 units of energy are produced for every one unit input into the system.
Embodiments of the present invention employ many different types of equipment, including reactors, centrifuges, decanters, a distillation column and many pumps and heat exchangers. Units discussed in this section have specification sheets that may be found at
According to an embodiment of a production process, there are three different types of reactors used. The first is the reactor used for the hydrolysis of the triglyceride into free fatty acids. These reactors are listed as R-101 and R-102, R-201 and R-202, and R-301 and R-302 in the process flow sheet. In these reactors the algae and acid are fed at a pressure of 2 atm and a temperature of 100° C. The conditions in the reactors are the same as the feed conditions, so no heating or pressurization needs to be done in the reactors, just the maintenance of these conditions. The reactors are glass-lined continuous stirred tank reactors (CSTR) and are agitated and jacketed. The glass lined reactors may assist with the acidic nature of the feed. The volumetric flow rates into each reactor are 8,585 gal/hr, according to embodiments of the present invention. The residence time for each reactor is half an hour, according to embodiments of the present invention. A half hour residence time may be chosen to reduce the size of each reactor and help in achieving the 99 percent hydrolysis of triglycerides into free fatty acids in a smaller volume. Once the volumetric flow rate and residence time are known, the size of each reactor may be calculated; according to some embodiments of the present invention, each reactor is calculated to be 4,293 gal. An oversize factor of 1.16 may be used, leading to the choice of a a 5,000 gallon reactor. Such a rector size may allow for head space and help minimize overfilling. Challenges in designing the reactors include selecting a residence time and operating conditions, according to embodiments of the present invention.
Each reactor may be jacketed to help maintain a selected temperature, for example 100° C. Cooling water and 100 psi steam may be used to heat and cool the reactor as necessary. The reactors have very little heat duty overall compared to other units in the process, as the inlets are already at operating conditions, according to embodiments of the present invention. The overall operating costs of the reactor are minimal due to low utility usage and maintenance cost, according to embodiments of the present invention. If one hydrolysis reactor must stop production for maintenance, then there are two other chains that can still produce biodiesel, according to embodiments of the present invention. The overall production of the plant will decrease but will not have to stop if one reactor needs maintenance, according to embodiments of the present invention.
The second type of reactor that may be used is for the esterification of the free fatty acids into the fatty acid methyl esters or biodiesel, according to embodiments of the present invention. These reactors are listed as R-103 and R-104, R-203 and R-204, and R-303 and R-304 in the overall process flow diagram. The feed to these reactors contains the components leaving the hydrolysis reactors which, according to embodiments of the present invention, may include: cell matter, water, glycerol, acid, and free fatty acids. Another feed stream to the reactor is the methanol for the esterification reaction. The total volumetric flow rate of both streams is 9,472 gal/hr, according to embodiments of the present invention. The conditions in these reactors are the same as the hydrolysis reactors with a temperature of 100° C. and a pressure of 2 atm, according to embodiments of the present invention. The esterification reactors may be glass-lined due to the corrosive nature of the feed. Glass-lined may be preferred in some cases over stainless steel due to the high cost of stainless steel. All of these reactors are CSTRs and are jacketed and agitated as well, according to embodiments of the present invention. The residence time of these reactors is also half an hour, according to embodiments of the present invention. The total residence time may be selected to be 30 minutes, according to embodiments of the present invention. In order to ensure the 85 percent conversion a total residence time may be selected to be one hour between the two reactors, according to embodiments of the present invention. The one hour may be divided across two reactors to help reduce the size of each reactor and help in achieving the desired conversion, according to embodiments of the present invention. The calculated volume for each reactor is 9,023 gallons, according to embodiments of the present invention.
The esterification reaction is an exothermic reaction that may result in a need to cool the reactors. The reactors may be cooled with cooling water through their jackets. The overall usage of cooling water to maintain the desired temperature in the reactors may be minimal, which leads to low operating cost due to utilities. Another operating cost is maintenance. If, for some reason, a reactor needs stop production due to maintenance, the other two branches may still be able to continue making biodiesel, according to embodiments of the present invention.
The final reactor is used to complete the esterification of free fatty acids to biodiesel, according to embodiments of the present invention. This reactor is listed as R-401 on the process flow diagram. After the product stream is separated from the byproducts (glycerol, water, etc), the biodiesel product stream is fourteen percent free fatty acids. The final reactor may be added to convert the remaining free fatty acids to biodiesel. The reactor feed may be an acid and methanol mixture, along with the biodiesel product containing fourteen percent free fatty acids. The conditions in the reactor are 100° C. and 2 atm, according to embodiments of the present invention. The volumetric flow rate through the reactor is 11,923 gal/hr, according to embodiments of the present invention. This reactor may also be a jacketed and agitated glass-lined CSTR. The residence time may be chosen to be one hour for reasons stated for the other esterification reactors. The size for the reactor is 11,923 gal, according to embodiments of the present invention. This reactor houses a low concentration of free fatty acids that need to react, according to embodiments of the present invention.
As with the other reactors, the temperature may be maintained using cooling water. Also, the amount of cooling water required is a minimal expense and so the overall yearly maintenance and cost is minimal, according to embodiments of the present invention. According to some embodiments of the present invention, if for some reason this reactor has to be shut down for maintenance, the whole plant would have to either stop production or the biodiesel product would have to be stored until the maintenance is completed. Therefore, two intermediate storage tanks may be included to allow for twenty-four hours of storage. Alternatively, a backup reactor may be provided to allow for maintenance during continuous production.
A process according to embodiments of the present invention utilizes three centrifuges to separate the cell matter from the liquid after the hydrolysis and esterification reactor sequence. After the triglycerides in the algae cells have been extracted and reacted, they may need to be separated out. The centrifuges are listed as C-101, C-201 and C-301 in the process flow diagram. 316 stainless steel may be used for the material due to the acidity of the solution. The centrifuges may be designed based upon the number of tons of solid per hour they were required to separate. The amount of cell matter removed is estimated at 23.85 tons/hr for each centrifuge, according to embodiments of the present invention.
According to embodiments of the present invention, decanters separate the heavy byproducts from the biodiesel product. Decanters are referred to as D-101, D-201, D-301 and D-401, according to embodiments of the present invention. The decanters are designed to allow a sufficient time for the aqueous phase (water, glycerol, acid, and methanol) to separate from the oil phase (biodiesel, free fatty acids). This is accomplished by a gravity separation due to the significant difference in density of the phases, according to embodiments of the present invention. The time taken for a separation is given by Equation 4:
In this equation μ is the viscosity of the light phase, ρH is the density of the heavy phase, and ρL is the density of the light phase. A calculation time for separation of two hours may be used as the residence time of the liquid in the decanter. The volumetric flow rate through D-101, D-201 and D-301 is 7,112 gal/hr, according to embodiments of the present invention. The volume of these decanters may be 14,107 gal, according to embodiments of the present invention. The decanters may be horizontal vessels. The length to diameter ratio may be selected as 6, to allow for proper settling of the heavy phase from the light phase. With such volume, the length of each decanter may be 60 ft. with a diameter of 10 ft., according to embodiments of the present invention. The decanters may be constructed out of 316 SS due to the acidity of the feed, according to embodiments of the present invention. D-401 has a different size, due to a volumetric flow rate of 16,791 gal/hr, according to embodiments of the present invention. With this flow rate the decanter size may be 33,305 gal, according to embodiments of the present invention. This volume with a length to diameter ratio of 6 yields a length of 58.9 fit and a diameter of 9.8 ft. According to some embodiments of the present invention, 100 percent the aqueous phase is removed from the oil phase in the decanters. The decanters require very little utilities to ensure proper operating conditions, resulting in a minimal yearly operating cost. If any of the decanters after the initial reaction scheme have to be repaired there may be others that can function while maintenance is completed. If the decanter after the final reactor were to need maintenance, then the flow from that reactor could be stopped or moved to a storage tank elsewhere, according to embodiments of the present invention. However, decanters are unlikely to require maintenance due to their simple design.
A flash drum reduce the volume of liquid sent to the distillation column by flashing off about half the water and almost all the methanol, leaving glycerol, acid and some water in the liquid phase, according to embodiments of the present invention. The flash drum is referred to as F-101 on the process flow diagram. The inlet flow is the combination of the three aqueous streams from the decanters, and the total volumetric flow rate is 10,100 gal/hr, according to embodiments of the present invention. Heuristics may be used to determine a residence time in the flash drum of 10 minutes, and that the volume of the flash drum may be twice the value of the volumetric flow rate times the residence time, according to embodiments of the present invention. From this, the volume of the flash drum may be calculated to be 3,972 gal. The conditions inside the column are 125° C. and 1 atm., according to embodiments of the present invention. At this temperature and pressure there is a 95 percent vapor fraction in the column. Flash drums may be vertical columns with a length to diameter ratio of five; such value may also be calculated with heuristics, according to embodiments of the present invention. The height of the flash drum may be 39.8 fit with a diameter of 8 ft, according to embodiments of the present invention.
The unit identification for the distillation column is T-101 on the process flow diagram. A distillation column separates the water and any residual glycerol and acid from the methanol, according to embodiments of the present invention. The purified methanol may then be recycled to become part of the reactor feed. This recycle process greatly reduces chemical costs. The water that leaves the bottoms of the column may be further processed. A software simulation engine, for example ASPEN, may be used to determine the number of trays required for the column, the reflux ratio and the flow rates of liquid and vapor through the column, according to embodiments of the present invention. The total number of trays for the column is 17 and the reflux ratio is 18.9, according to embodiments of the present invention. The methanol product leaving the top of the column may be 99 percent pure. The height of the column may be determined by the number of trays with 2 foot spacing, which would make the column 34 feet tall. The tray type may be selected to be sieve trays, and the tray material may be determined to be 316 SS due to the possibility of acidic conditions in the column. The flow rates for the liquid and vapor may be used to calculate the column diameter, which is 5.45 fit in some embodiments. The overall weight of the column is 8,421 lbs., according to embodiments of the present invention. The column material may be selected to be 316 SS as well due to the acidity of the solution. The column pressure is 2 atm to make for a smaller diameter and a higher temperature of vapor leaving the column, according to embodiments of the present invention.
The main operating cost of the distillation column consists of utilities for the reboiler and the condenser. The reboiler may be supplied with 100 psi steam and the condenser may be supplied with cooling water, according to embodiments of the present invention.
Pumps in the system may operate to raise the pressure in the feed streams from one atmosphere to a pressure of two atmospheres, according to embodiments of the present invention. The pumps are designated as P-101, P-201, P-301, P-401 and P-501. The pumps may be constructed of 316 SS. The sizes of the individual pumps are described in
Heat exchangers H-101, H-201, H-301, and H-501 heat feed streams from their initial temperature to the production temperature of 100° C., according to embodiments of the present invention. H-401 may be used to heat the aqueous solution from the decanters to 125° C. to prepare it for the flash drum, according to embodiments of the present invention. Heat exchangers may be made of 316 SS due to the acidity in the solution. The surface area in the heat exchanger and overall heat transfer coefficients may be calculated using a facility design software program, for example ASPEN, according to embodiments of the present invention. Heat exchangers are not duplicated according to embodiments of the present invention, and due to the fact that three parallel streams exist, if H-101, H-201, or H-301 need maintenance only one of the three branches needs to stop production.
Various modifications and additions can be made to the exemplary embodiments discussed without departing from the scope of the present invention. For example, while the embodiments described above refer to particular features, the scope of this invention also includes embodiments having different combinations of features and embodiments that do not include all of the described features. Accordingly, the scope of the present invention is intended to embrace all such alternatives, modifications, and variations as fall within the scope of the claims, together with all equivalents thereof.