Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility Study
March 7, 2019
TABLE OF CONTENTS
Section Description Page 1 Introduction 3 1.1 Existing Facility 3 1.2 Existing Anaerobic Digesters 4 1.2.1 Digester Condition Assessment 5 1.2.2 Historical Biogas Utilization Facilities 7 1.2.3 Existing Flares 7 1.3 Existing Biogas Characteristics 7 1.3.1 Existing Biogas Quality 7 1.3.2 Existing Biogas Quantity 8 1.3.3 Raw Biogas Sampling 10 1.3.4 Existing Peaking Factor 10 1.4 Biosolids Utilization 11 1.5 Projected Biosolids and Biogas Quantities 11 1.5.1 Scenario 1 Future Gas Production 12 1.5.2 Scenario 2 Future Gas Production 12 2 Digester Improvements 15 2.1.1 Maintenance Needs 15 2.1.2 Process Improvements 16 2.1.3 Improvement Phasing 19 3 Renewable Natural Gas 21 3.1 Renewable Fuel Standard 21 3.1.1 The Producers Role 21 3.2 Gas Utility Connection 22 3.2.1 Gas Utility Requirements 23 3.2.2 Interconnect Station 24 3.3 Biogas Upgrading General Design Criteria 24 3.3.1 Biogas Purification Technologies 24 3.3.2 Phasing Strategy Approach 24 3.3.3 Raw Biogas Quantity Design Criteria 25 3.3.4 Raw Biogas Quality Design Criteria 25 3.3.5 Finished Biogas Quality Design Criteria 26 3.4 Facility Impacts to Cost 27 3.4.1 General Type of Construction 27 3.4.2 Area Classification Requirements 27 3.4.3 Budgetary Constraints 28 3.5 Potential Gas Treatment Train 28 3.5.1 Engineers Proforma Evaluation 33 4 Implementation Steps 37 5 Schedule 38
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1 INTRODUCTION
The Western Virginia Water Authority (Authority) owns and operates the Roanoke Regional Water Pollution Control Facility. The Authority formed in 2004 consolidating water and wastewater operations for the City and County of Roanoke. Following inception of the Authority the counties of Franklin and Botetourt were added and the Authority contract operates the Town of Fincastle facilities. There are more than 51,000 wastewater service accounts managed by the Authority. Roanoke Regional Water Pollution Control (WPC) Facility is located in Roanoke, Virginia and was originally constructed in 1951. The facility is located within the Roanoke City limits and treats municipal wastewater from throughout the Roanoke Valley. This memorandum will evaluate certain processes within the existing treatment facility and provide recommendations for possibly upgrading biogas produced during anerobic digestion to pipeline quality renewable natural gas. EXISTING FACILITY
The Roanoke Regional WPC Facility is permitted to treat 55 million gallons a day (MGD) with an average daily flow of 37 MGD. Facility flow is highly variable depending on the amount of infiltration and inflow during wet weather. During dry weather conditions the facility generally treats in the mid 20 MGD, and during wet weather conditions the facility influent increases significantly with a historical peak flow of 137.4 MGD in November 2009. The treatment facility processes include the screening of debris and collection of grit from the influent, primary clarification, aeration, secondary clarification, coagulation, filtration, disinfection and discharge to the Roanoke River. Solids collected during primary and secondary clarification are digested with the digested biosolids stored in lagoons on site prior to land application. Land application is free to farm fields producing non-food crops. As a byproduct of the treatment process, biogas is produced during anaerobic digestion of biosolids in the onsite digesters. More in-depth discussion of the treatment plant facilities along with a liquid and solids schematic are included in Appendix A . The facility currently has a septage receiving station with possible plans to replace the station. The receiving station is a rock box combined with a step screen. Currently septage is received upstream of the raw influent screening. Acceptance at the front end of the plant results in grease accumulating in the thickener. Accumulated grease requires skimming and landfilling of the grease, approximately 2 tons per day. Septage received by the facility is broken into the source with values for 2018 shown in Figure 1 . Historically, the term “Septage” has been used generically by Western Virginia Water Authority staff to represent any waste material that is hauled to the site from outside sources. Through the remainder of the report, septage is delineated as septage (meaning only domestic septage), fats-oils-grease (FOG) from grease traps, and industrial (food) waste. The term "high strength waste (HSW)" generically refers to FOG or industrial waste.
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Roanoke WPC Plant 2018 Septage Receiving 1,600,000
1,400,000
1,200,000
1,000,000 Total Volume 800,000 Domestic Volume 600,000 Grease Volume (gallons) Industrial 400,000
200,000
0
Figure 1: 2018 Septage Receiving Volumes
EXISTING ANAEROBIC DIGESTERS
Biosolids collected in the primary clarifies are processed by a gravity thickener prior to being fed to the digesters. The Roanoke WPC Plant has two gravity thickeners, but operations staff have indicated one thickener is used at a time. Biosolids collected in the secondary clarifiers are thickened using a dissolved air flotation (DAF) system to concentrate the solids prior to digestion. The existing anerobic digestion facilities consist of seven primary digesters and three secondary digesters, as shown in Figure 2. The four original digesters were constructed in the 1950’s with two more added in 1965 and four more digesters constructed in the 1970’s. Concrete primaries have reinforced concrete mat foundations and reinforced concrete roof slabs supported by interior columns. The primary digesters are all buried or partially buried with concrete roofs. Buried digesters are accessible via service hatches. Currently two of the secondary digesters have floating steel covers and the third secondary is equipped with a Dystor® cover. All of the digesters have a diameter of 90 ft and a 1 million-gallon (MG) capacity. Currently, only four primary digesters are operated at a time (Digesters 1, 2, 4 and 5). The primary digesters are fed through a common header on rotating 2-hour intervals to balance the amount fed to each digester. Operations staff indicated the digesters generally operate with at least 30-days of solids retention time (SRT) and an operating pressure of 8 to 8.5 pounds per square inch (psi). Currently the digesters are loaded at an average rate of 53 lbs VS/1,000 ft 3, the total feed concentration averages 3.9% and the feed volatile solids concentration averages
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2.9%. The digesters are designed to operate at mesophilic temperature range. However, digester temperatures have fallen as low as mid 80 degrees Fahrenheit. None of the digesters are currently mixed, which has decreased the WPC Plant capacity to 55 MGD from 62 MGD as designed for mixed digester operation. Initially the digesters were mixed utilizing the Pearth® gas mixing system which discharges compressed biogas into the digester on a sequential basis to form mixing plumes. Due to challenges the facility faced with the discharge lances plugging, the digester mixing system was abandoned.
Figure 2: Anerobic digestion facilities at Roanoke WPC Plant. Currently operated primary digesters are shown in green and red are not being utilized.
1.2.1 DIGESTER CONDITION ASSESSMENT
In 2009 Hazen & Sawyer conducted inspections of Primary Digesters 1, 2 and 3. According to the inspection memo; the inspections included visual assessment of the interior walls, floor slab and roof slab. Visual inspection of the bottom of the roof slabs was limited due to the inspections having to be conducted from the digester floor. Exterior surfaces were not inspected due to the digesters being buried. To inspect Digesters 1, 2 and 3:
• Backfill on the roof was removed • Digester was partially dewatered • Interior walls pressure washed According to the memos, inspections found similar conditions in the three digesters, therefore the following condition description is applicable to Primary Digesters 1-3. The following interior conditions were noted for the primary digesters:
• Sound condition of internal support columns • Delamination of the topping placed over the bottom slab • Spalling of the floor slab • No significant seepage through the walls was noticed even though minor cracking was present • Spalling of the walls where the roof slab meets the walls • Significant deterioration of the top slab including:
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o Extensive spalling o Exposure of rebar and welded wire fabric at all interior column supports and where the slab meets the walls o Spalled concrete hanging from the roof by the exposed reinforcing o Extensive cracking, through the entire slab depth in many locations, on the bottom of the roof slab Based on the condition noticed in 2009 it was determined the floor slabs, walls and columns were in satisfactory condition for operation. Based on the deteriorated state of the roof slab, the cause of damage was determined to be from over pressuring of the digesters rather than concrete chemical corrosion, as was expected by the Authority. During a high flow event the digester overflow was plugged resulting in the digesters becoming pressurized, causing an uplifting pressure on the digester roof slabs. Facility staff have indicated the over pressure event more substantially damaged Primary Digesters 6 & 7. Digesters 6 & 7 have, therefore, been taken offline. Primary Digester 3 is not in service, due to concerns of gas leakage, but could be operated if needed. While inspections of Primary Digesters 4-7 are not available, it is assumed the conditions are comparable to Primary Digesters 1-3 other than the items noted above for Primaries 3, 6 & 7. The existing gas storage (approximately 186,000 cf total assuming equal storage in 3 steel covered secondary digesters) within the system is included in the three secondary digesters. The storage acts to attenuate the gas utilized by the CHP system. A SCADA system monitors the digester gas pressure and cover levels. Upgrades to the digestion facilities were completed in 2004 as part of Roanoke Regional Water Pollution Control Plant Contract B – Process Train Improvements project and additional upgrades were completed in 2010 as part of the Anaerobic Digester Modifications and Biogas Energy Recovery Facilities Project. None of the upgrades addressed the structural issues with the digesters. Upgrades included:
• Installation of gas dryer and compressor facility • Replacement of gas safety equipment and gas collector dome on Primary Digester 1 • Completed digester steam-to-hot water heat exchanger and heat recovery modifications • Installation of a siloxane removal system • Replacement of circulation water pumps • Sludge feed pumps • Condenser and hot water pumps • Replacement of Dystor® cover on Secondary Digester A with a floating steel cover • Digester gas and hot water piping modifications
Siloxanes are removed via two carbon adsorption vessels. According to plant staff, the carbon has not been changed since installation in 2010. Siloxane removal was installed to protect the combustion equipment in the combined heat & power (CHP) system. 6
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1.2.2 HISTORICAL BIOGAS UTILIZATION FACILITIES
The gas produced during anerobic digestion has been used on site beneficially or flared. The Authority installed two 596 kW generators in 2012 to run on the gas produced at the facility as part of a combined heat and power (CHP) system. When both generators are running about 1 megawatt (MW) is produced. The CHP system was designed to heat the digesters and power a portion of the treatment plant. Heat produced, but not used in the digesters was used in on-site in absorption chillers for facility heating and chilling. The CHP system has offset the power purchased by the Authority by approximately 7,577 megawatt hours (MWh) a year. This is 30-50% of the facility energy requirement based on the facility treatment flow and energy needs. One of the generators recently was damaged and as a result all of the digester produced gas has been sent to the boilers or flared. Prior to onsite utilization the raw biogas passes through iron sponge and carbon adsorption vessels to remove hydrogen sulfide (H2S) and Siloxanes. Facility staff indicated that in 2017 the WPC Plant burned a total of 135,200,000 cubic feet of gas in total, including gas to all equipment and flares. A BTU analyzer would be needed to run the engines on dual fuel. 1.2.3 EXISTING FLARES
The facility is equipped with two (2) 8-inch candle stick flares that were installed as part of the Roanoke Regional Water Pollution Control Plant Contract B – Process Train Improvements project in 2004. Gas can be fed to the flares from the boiler building or from the digesters based on an over pressure system. The flares are equipped with a gas meter. An 8-inch candlestick flare would require a minimum gas flow of 120 standard cubic feet per minute (scfm) and a maximum flow of 1,360 scfm. Based on this, the two existing flares should provide more than enough capacity for the anticipated facility gas production rates. EXISTING BIOGAS CHARACTERISTICS
To design the biogas upgrading system, an accurate and historical record of biogas composition is essential. The gas produced at the Roanoke Regional WPC Plant is a product of anaerobic digestion of municipal wastewater collected by the Western Virginia Water Authority collection system. If the facility were to start accepting high-strength waste (HSW), the raw gas characteristics would be anticipated to change as a result. 1.3.1 EXISTING BIOGAS QUALITY
To determine the composition of the raw biogas, a sample of the gas must be taken and sent to a laboratory for analysis, with the report following a week after receipt by the lab. A sample was obtained from the Roanoke WPC Plant using a Tedlar bag. The sample was obtained downstream of the existing hydrogen sulfide removal system. Sampling was done for evaluation of a CHP system at that time. For this memo, data collected on December 29, 2008 is presented below in Table 1. This is one sample and additional data is needed to more fully characterize the raw gas. The value of nitrogen in this sample is higher than expected, typically anerobic digester gas contains 0.1-2.5%. Due to the sample being obtained via Tedlar bag there may have been nitrogen diffusion into the sample. If
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Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility Study
March 7, 2019 additional sampling confirms the elevated level of nitrogen upgrading of the gas may not be practical. The maximum heating value expected from the below gas sample, assuming 100% removal of the CO2, would be 895 BTU.
Table 1: Raw gas samples from 2008 Constituent Value Units Hydrogen Sulfide ND ppb Total Sulfur 160 ppb Oxygen 0.57% % Vol Nitrogen 7.42% % Vol Methane 61.2% % Vol Carbon Dioxide 30.8% % Vol C6 + Hydrocarbons 23,407 ppb Siloxanes 3,477 ppb Total BTU 621 BTU/ft3 Specific Gravity 0.885 relative to air 1.3.2 EXISTING BIOGAS QUANTITY
Existing biogas production was determined based on volatile solids destruction data and metered gas flow recorded in 2017-2018. The information from each is summarized below, and recommended design criteria is included in Section 3.4 of this technical memorandum. Estimates from Actual Volatile Solids Reduction in 2018: Monthly samples are obtained from the digesters and analyzed. The data obtained includes information required to determine the volatile solids destruction efficiency in the primary and secondary digesters. Utilizing the digester data and the estimated amount of biogas production using an assumed 12 to 15 cubic feet per pound of volatile solids reduced is shown in Table 2. Data in the table is from January to May 2018 to allow for comparison to the data obtained in 2018 from the onsite meters. Table 2: Expected biogas production based on volatile solids destruction data for 2018 Average Production Minimum Maximum Production Year (SCFM) Production (SCFM) (SCFM) January 143 72 218 February 98 11 244 March 147 93 182 April 197 67 287 May 129 30 272
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Metered Biogas Production 2017-2018: Metered biogas production for 2017 and 2018 indicated the following data in Table 3. Biogas is metered on each of the four primary digesters currently in use. The meters have a measurement range of 0 to 7,500 scfm according to the rating listed on the meter. Each meter records a gas flow rate in standard cubic feet per hour (scfh) at 1:20 AM and 1:20 PM daily. According to conversations with the plant staff, the meters are not calibrated regularly. Data is included for the entire year of 2017, but only data from January – May of 2018 is included. In the latter half of 2018, the facility received a load of high strength waste that disrupted the treatment process and impacted the data. Therefore, only data from the first 5 months of 2018 has been included.
Table 3: Historical metered biogas production from 2017 and 2018 Production Average Minimum Maximum Standard Year Production Production Production Deviation (SCFM) (SCFM) (SCFM) (SCFM) 2017 335 40 615 58 2018 358 185 610 25
Monthly metered data for January – May 2018 is included in Table 4 for comparison to the vales determined based on volatile solids destruction in the digesters shown in Table 2 .
Table 4: Biogas Production Recorded by On-Site Meters for 2018 Average Production Minimum Maximum Year (SCFM) Production (SCFM) Production (SCFM) January 354 185 610 February 357 234 484 March 376 277 476 April 386 236 494 May 320 190 416
Gas Production Comparison: There is a significant difference between the expected biogas produced due to volatile solids destruction and the metered biogas volumes recorded. At the upper expected gas production based on 15 cf of biogas produced per pound the difference is 0.6-1.2 times greater. Based on this information it is recommended that the meter installation configuration be reviewed. It is possible there is not enough straight pipe upstream and downstream of the meters to allow for accurate readings. Following that, it is recommended that the meters be calibrated.
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1.3.3 RAW BIOGAS SAMPLING
Additional raw gas sampling is suggested for the facility. The following suggested continued sampling for 2019 has been developed based on current sample values, unknowns, potential gas tariff requirements, and experience from other projects. The proposed protocol, included in Appendix B , includes the constituents to be monitored prior to design along with the recommended methods, and frequency as indicated in Table 5. All samples should be obtained via summa containers to minimize nitrogen and oxygen diffusion in the sample during transport. Raw gas samples should be obtained prior to any gas treatment, specifically H 2S, VOC and siloxane removal. The samples should be spread out with 1 sample taken a week, on differing days, to capture raw gas characteristic variations with time. The siloxane and VOC testing results need to indicate individual speciation and parts per million or parts per billion. It is important to note that the suggested sampling plan is proposed for the immediate future and necessary for the design of an upgrading system. A post-project sampling plan would need to be developed in partnership with the biogas upgrading equipment manufacture and receiving gas pipeline for facility operation. Table 5: Proposed raw gas sampling for 2019 Unit of Samples Constituent Recommended Method Measurement Needed Oxygen Mol % Grab Sample/Lab Analysis 10 Samples Nitrogen Mol % Grab Sample/Lab Analysis 10 Samples Methane Mol % Grab Sample/Lab Analysis 10 Samples Carbon Dioxide Mol % Grab Sample/Lab Analysis 10 Samples Gross Heating BTU/DSCF Grab Sample/Lab Analysis 10 Samples Value Carbon Mol % Grab Sample/Lab Analysis 10 Samples Monoxide Hydrogen Sulfide PPMV Grab Sample/Lab Analysis 10 Samples VOC's PPBV Grab Sample/Lab Analysis 10 Samples Siloxanes PPBV Grab Sample/Lab Analysis 10 Samples Ammonia PPMV Grab Sample/Lab Analysis 5 Samples Total Sulfur PPMV Grab Sample/Lab Analysis 10 Samples Moisture Content PPMV Grab Sample/Lab Analysis 5 Samples
1.3.4 EXISTING PEAKING FACTOR
The raw gas produced by the digesters is not always consistent and can be subject to variations. These variations can be caused by feed inconsistency, biosolids composition, mixing consistency, and temperatures in the digesters. To account for the variations, peaking factors must be taken into consideration.
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According to historical metered data, the minimum hourly flow is roughly 10% of the annual average flow rate, while the maximum hourly flow is roughly 170% of the annual average flow rate. Based on recorded volatile solids loading and destruction rates, the minimum and maximum flows are approximately 10 and 200% of the annual average flow rate respectively. BIOSOLIDS DISPOSAL
The existing biosolids produced by the WPC Plant originate from the septage received and the sewage that flows into the facility through the headworks. After digestion, the biosolids are sent to one of four biosolids lagoons as the final step in solids treatment. While in the lagoons, organic matter that was not fully digested in the digesters, are consumed by anaerobic and aerobic bacteria allowing for further stabilization. Biosolids remain in the lagoons for approximately nine months stabilizing. Following this timeframe, the lagoon is decanted and the biosolids are removed and land applied to crops not destined for human consumption. Farmers are not charged for the nutrient rich biosolids applied to their field. The land application of biosolids is done in accordance with EPA standards, 40 CFR Part 503. The quality and quantity of biosolids produced by the facility impacts the possible use of the biosolids. Currently the facility is meeting all regulations for land application of biosolids and does not want to appreciably change current practices. Therefore, the sludge characteristics are not determined further. Due to not having a dewatering system, the facility does not have high volume side stream return flow. Due to this, nutrient recovery of phosphorus was not pursued when previously investigated. The facility has not historically had struvite issues. PROJECTED BIOGAS QUANTITIES
Projections are made to estimate the range of potential biogas quantities that could be produced. Future growth in biogas quantities could be the result of facility upgrades, additional domestic waste or due to the addition of Fats, Oils and Grease (FOG), High-Strength Waste (HSW) or Municipal Biosolids from other plants within the region. External sources of waste could be received at the plant from multiple sources: Domestic Septage: Waste from septic tanks, temporary sanitary waste facilities (porta-potties), manhole cleaning and pumper trucks. Domestic could be introduced at the head of the plant because of inert materials that would likely be included. Inert materials will be removed by the headworks processes of screening, and grit removal. Grease Trap/Interceptor Waste: Fats, oils, grease (FOG). Clean grease could be introduced to anaerobic digesters following grinding and screening. Industrial: Any other HSW which would likely contain high organic loadings or high levels of easily digestible volatile solids. Industrial should be reclassified according to cellulosic content so that integrity of D3 RINs as described in Section 3.1.1 can be maintained.
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Potential wastes received will need to be classified differently in the future based on how EPA would classify the gas produced from the particular waste sources. More detail on EPA classifications is given in Section 3.1.1 of this report. Establishing potential gas production needs to be viable based on the plant infrastructure and available waste to be received. To determine the potential digester loading, it is important to verify the digesters are not overloaded with volatile solids and that there is sufficient digester volume to provide enough residence time to keep proper digester operation. The following scenarios were calculated based on digester solids reports, and 15 cf gas produced per pound (15 cf/lb) of volatile solids destroyed due to anerobic digestion. Two potential gas production scenarios were developed based on digester mixing, heating and receiving additional FOG/HSW. Mixed Digestion: Adding mixing to the digesters would allow for decreasing the SRT and increasing the solids loading of the digesters. Currently the unmixed facility is limited by Virginia Department of Environmental Quality Sewage Collection and Treatment Regulations to a minimum 30-day SRT or 40 lbs/1000 ft 3 per day. The regulations for completely mixed systems allow for average digester loading rates up to 200 lbs/1000 ft 3 per day and a minimum of 15 days SRT. If completely mixed, only two (2) primary digesters would be needed to handle current average digester loading. Heated Digestion: If the digesters are adequately heated the volatile solids destruction year around should be similar to what is currently experienced in the summer months. An evaluation of July and August 2018 data indicated 65% volatile solids destruction. The high VS destruction rate is due to the solids retention time (SRT) of greater than 30 days. If the SRT were decreased the VS destruction rate would also decrease. Applying this to the current volatile solids loading and digester feed rate, the digesters have the potential to produce 200 scfm of biogas year around. This is a very low gas production value for a biogas upgrading project. Receiving FOG/HSW: There have been discussions about the facility potentially receiving additional FOG/HSW. The Authority has identified multiple potential sources of FOG/HSW to evaluate acceptance of. For gas production analysis it was assumed the FOG/HSW would have a volatile solids concentration of 72,000 mg/L based on values observed in Des Moines, IA; Sheboygan, WI; Johnson County, KS; Fort Worth, TX and elsewhere. 1.5.1 SCENARIO 1 FUTURE GAS PRODUCTION
Scenario 1 is based on the plant not taking in additional waste beyond the service area growth. In this scenario two fully mixed and uniformly heated primary digesters would be required. Anticipated gas production would be 200 scfm in the first year of production. Over 20 years the gas production would be anticipated to increase to 250 scfm. All of the gas in this scenario is expected to be classified as D3 gas 1.5.2 SCENARIO 2 FUTURE GAS PRODUCTION
Scenario 2 would include the same D3 gas production as Scenario 1 (up to 250 scfm in 2 primary digesters), but would include the additional of FOG/HSW after 5 years of operation. The
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Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility Study
March 7, 2019 establishment of a robust FOG/HSW receiving program would take time, public education, an investment into receiving and storage facilities, and additional labor. As a first step, it has been assumed that the Authority could maximize the use of two digesters with the addition of FOG/HSW. This would require consideration of the feedstock recipe because buffering capacity would be required from one or more sources, such as alkalinity chemical addition, wastes with high TS content, pretreated septage, etc. Because FOG/HSW facilities can typically operate at higher loading rates than facilities with municipal biosolids, it has conservatively been assumed that these two FOG/HSW digesters could be loaded with up to 225 lbs. VS/day / 1,000 cf or up to 320 lbs. COD/day / 1,000 cf. Based on these assumptions, two maximized FOG/HSW digesters could produce as much as 520 scfm consistently. This would require the addition of approximately 86,000 gpd of FOG/HSW, or roughly 15 trucks per day. While it would be possible to mix FOG/HSW with municipal sludge and co-digest this material in more than 2 digesters, keeping the sources (municipal sludge and FOG/HSW) separate would be advantageous to maintain D3 RIN credit for the gas generated exclusively from municipal sludge. The gas generated from the separate FOG/HSW digesters would most likely be classified as a D5 (non-cellulosic) RIN. Gas production associated with Scenario 2 is shown in Table 6.
Table 6: Potential gas production accepting new FOG/HSW (2 digesters for municipal waste and 2 digesters for FOG/HSW) D3 digester Volume 699,800 ft3 Flow to digesters 122,818 gpd Volatile solids to digesters 29,660 lbs/day Solids loading 106 lb VS/1,000 ft3 Solids Retention Time 17 days 288,000 CFD D3 Gas Production 200 SCFM
Buffered HSW gas production
D5 digester Volume 279,700 ft3 HSW to digesters 85,600 gpd HSW volatile solids concentration 72,000 mg/L Solids loading 184 lb/1000 ft3 Solids Retention Time 24 days 748,800 CFD D5 Gas Production 520 SCFM
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Based on the potential increased gas production resulting from starting a FOG/HSW receiving program it is recommended the Authority consider this in more detail. When a program is being developed it is important to consider available feedstock quantity, quality and EPA classification and the resulting impact on existing digestion facilities. Additional system modifications that could increase gas production would be increased sludge thickening prior to introduction to the digesters and heating the digesters consistently, while these items are outside the scope of this memo they are included on a high level. Thickening of the sludge would decrease the volume of water in the digesters and could increase the effectiveness of the current heating system. It also would increase the available digester volume to be loaded with volatile solids. Consistent heating of the digesters might increase the volatile destruction rates resulting in increased gas production.
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2 DIGESTER IMPROVEMENTS
Information in Section 1 indicated some limitations associated with the current equipment and operation. Upgrades to the existing anerobic digestion facilities would protect the long-term operation and increase solids throughput and biogas production of the facility. Proposed facility upgrades are discussed in this section. A high-level engineer’s opinion of probable construction cost for each proposed upgrade and improvement project is included herein. 2.1.1 PRIMARY DIGESTER MAINTENANCE NEEDS
Upgrades to the existing anerobic digestion facilities are necessary to improve the integrity of the structure, improve long-term operation, and increase the useful life of the facility. As discussed in Section 1, Primary Digesters 3, 6 and 7 are not in operation. The other four Primary Digesters are believed by the Authority to need maintenance to protect the structural integrity of the covers. The last structural inspection occurred in 2009. Structural reinforcing was exposed in the roof slabs during that inspection, and additional corrosion since then is likely. It is recommended another inspection be performed before implementing rehabilitation improvements. A new inspection may reveal the need to replace covers on Primary Digesters 3, 6 and 7 in lieu of rehabilitation. The rehabilitation of covers is assumed to include digester cleaning, sealing of cracks, surface preparation and painting of the cover and top 1/3 of interior digester wall height, and replacement of faulty biogas safety equipment, where necessary. Additional digester rehabilitation assumed for all seven digesters would include the addition of structural coating to floor slab, replacement of miscellaneous piping, and replacement of miscellaneous biogas safety equipment and instrumentation. A high level cost estimate for rehabilitation of all seven primary digesters, at one time, is included in Table 7.
Table 7: Digestor Rehabilitation (7 Digesters) – Engineer’s Opinion of Probable Construction Costs
Cost per Units to Upgrade Unit Total Cost Unit Complete Remove Contents, Clean Digester $ 100,000 EA 7 $ 700,000 & Replace Sludge Install and Remove Scaffolding $ 25,000 EA 7 $ 175,000 Seal Digester Cracks (roof slab and $ 100 LF 5,600 $ 560,000 support columns) Paint Digester Interior $ 20 SF 60,366 $ 1,207,314 Low Strength Grout $ 600 CY 126 $ 75,600 Test Digester $ 9,000 EA 7 $ 63,000
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Digester Gas Handling & Safety $ 93,000 EA 7 $ 651,000 Equipment Sub-Total Construction - - - $ 3,431,914 General Contractor Overhead & 12 % - $ 411,829.69 Profit Contingency 30 % - $ 1,029,574 Engineering 12 % - $ 411,830 Total - - - $ 5,285,148
2.1.2 PROCESS IMPROVEMENTS
Addition of digester mixing would allow for the facility to increase capacity and allow for better operation of the digesters. Mixing of digesters reduces scum formation, results in less settling of solids, can increase volatile solids destruction and can increase production of biogas. One option, observed by the operations staff at the Pepper’s Ferry Regional WWTP, would be a cannon mixing system. This system utilizes a "shear tube" that creates shear by generating high upward velocities in tube using a gas principal. The cannon system does not include small diameter lances or orifices that can clog, like some other gas mixing based systems. Biogas produced in the anerobic digesters is compressed and recirculated. The cannon system is mounted to the floor and would not be supported by the digester lid. Cannon mixing systems are suited for installation in round digesters. Each cannon mixer covers an area of approximately 16 to 18 feet in diameter, therefore the 90-foot diameters at Roanoke would be equipped with 10 floor mounted cannon mixers per digester. Each mixer would be 24-inches in diameter. Final mixer design would need to be completed in association with the cannon mixer supplier. Table 8 includes an engineer’s opinion of probable construction cost for the installation of a cannon mixing system to the seven (7) primary digesters. A cannon mixing system proposal is included in Appendix C, note the proposal is for upgrading a single digester, but communication with the supplier indicated the cost saving per digester for installing multiple systems at once. It is assumed the mixing system will be installed at the same time as primary digester rehabilitation therefore cleaning and sludge removal is not included in the below cost estimate. Due to the mixing system requiring compression of the gas, there is a yearly operations cost. Suez, the cannon mixing system supplier, indicated each digester would require 271,560 KWh per year. For all seven (7) of the digesters the total electrical requirement would be 1,900,920 KWh per year. Based on historical electrical cost at the Roanoke facility the yearly cost for mixing a single digester is expected to be $16,294 and $114,055 for all seven (7) digesters. Costs are based on the historically paid cost for electricity of $0.06/KWh. It is expected that the existing electrical system will be used for the new mixers.
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Table 7: Installation of Mixing System – Engineer’s Opinion of Probable Construction Costs Cost per Units to Upgrade Unit Total Cost Unit Complete Digester Mixing System $485,714 EA 7 $ 3,400,000.00 Remove Existing System $ 15,000 EA 7 $ 105,000.00 Miscellaneous Piping $ 5,000 EA 7 $ 35,000.00 and Controls Install Digester Mixing 12 % - $ 408,000.00 Systems Sub-Total Construction - - - $ 3,948,000.00 General Contractor 12 % - $ 473,760.00 Overhead & Profit Contingency 30 % - $ 1,184,400.00 Engineering 12 % - $ 473,760.00 Total - - - $ 6,079,920.00 The Dystor® cover on Secondary Digester “A” was replaced in 2009 by a floating steel cover. The remaining Dystor® cover on Secondary “B” is at the end of its useful life and operational experience indicates a preference to the floating steel covers. The Dystor® cover should be replaced with a floating steel cover, matching the other two secondary digesters. Installation of a floating steel cover would require capital cost expenditure, as indicated in Table 8, but minimal or no operational cost.
Table 8: Installation of Floating Steel Cover – Engineer’s Opinion of Probable Construction Costs Cost per Units to Upgrade Unit Total Cost Unit Complete Remove Dystor Cover $ 50,000 EA 1 $ 50,000 Remove Contents, Clean Digester & $ 100,000 EA 1 $ 100,000 Replace Sludge Install Floating Steel $ 425,000 EA 1 $ 425,000 Cover $ Test Digester $ 9,000 EA 1 9,000 Digester Gas Handling $ 93,000 EA 1 $ 93,000 & Safety Equipment Miscellaneous Piping $ $ 5,000 EA 1 and Controls 5,000 Equipment Installation 12 % $ 81,240 Sub-Total Construction - - - $ 763,240 General Contractor 12 % - $ 91,589 Overhead & Profit Contingency 30 % - $ 228,972
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The Authority is interested in upgrading the current septage receiving station and possibly adding a second receiving station for FOG/HSW waste receiving. Initial receiving station layout, as designed by the Authority, is shown in Figure 3. The receiving station is set up to keep the domestic septage and FOG/HSW separate and feed digesters independently. The planned 5,000 gal storage tanks would allow for buffering of the waste prior to injecting to the digesters. Increasing the FOG/HSW buffer basin volume to 80,000-100,000 gallons would allow for 24-hour buffer time when accepting FOG/HSW in the gas production Scenario 2 described in Section 1.5.2. The Authority is interested in keeping received waste separate from the facility liquid train, as would occur due to feeding o the headworks rather than the digesters, to protect the liquids treatment from immediate impacts due to waste received. From a marketing of biogas standpoint, isolation of the two feed substrates would be exceedingly important. It is assumed the United States Environmental Protection Agency (USEPA) will classify domestic septage as D3 RIN and the FOG/HSW as a D5 RIN. RINS are described in Section 3.1 of this report. EPA has not made a declaration one way or another and final determination will require the facility applying for classification. Separation of the waste would require the addition of backflow valves on the downstream side of the buffer tanks or two pumping stations dedicated to each waste source. Due to the Authority already planning on completing these facility upgrades as a separate project a cost estimate is not included in this report.
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Figure 3: Possible future domestic and FOG/HSW receiving station (Provided by the Authority ) 2.1.3 IMPROVEMENT PHASING
It is recommended mixing be installed in the digesters at the same time structural rehabilitation is completed and that the Dystor® cover be replaced as part of the same project. This would allow for decreasing contractor mobilization costs and result in fully upgraded digesters in one project. Due to the facility not requiring all digesters to be in operation it is suggested the digester upgrades be completed in two phases. Possible phasing would be to complete rehabilitation of the three empty digesters in Phase 1 and completing 1-4 additional digesters when the volume is needed. Anticipated construction costs for structural rehabilitation of the digesters and installing mixing in two phases are
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Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility Study
March 7, 2019 indicated in Table 9. The overall construction cost is increased, but spreads the cost out over a longer time period. The anticipated OM cost for Phase I would have a yearly electrical draw of $48,882 per year. Installation of an upgraded septage handling station is anticipated to be a separate project and occur following digester Phase I work and possibly in conjunction with digester Phase II work. The number of digesters to be rehabilitated in Phase II is directly tied to the volume of septage received to the facility and therefore the two items should be considered together.
Table 9: Anticipated Phasing Costs – Engineer’s Opinion of Probable Construction Costs Phase I Phase II Cost per Units to Units to Unit Total Cost Total Cost Unit Complete Complete Primary Digester Rehab $ 504,982 EA 3 $ 1,514,945 4 $ 2,019,927 Install Mixing $ 580,920 EA 3 $1,742,760.0 4 $ 2,323,680 Replace Dystor® Cover $ 1,175,390 EA 1 $1,175,389.6 - - Sub-Total Construction - - - $ 4,433,095 - $ 4,343,607 General Contractor 12 % - $ 531,971.34 - $ 521,232.79 Overhead & Profit Contingency 30 % - $ 1,329,928 - $ 1,303,082 Engineering 12 % - $ 531,971 - $ 521,233 Total - - - $ 6,294,994 - $ 6,167,921
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3 RENEWABLE NATURAL GAS
The following section includes information on the producing and selling of renewable natural gas (RNG). Also included is a potential gas upgrading train to take the anaerobically produced digester gas and upgrade to pipeline quality natural gas at the Roanoke Water Pollution Control Plant. High level cost estimates associated with the upgrading process also is included. RENEWABLE FUEL STANDARD
The first Renewable Fuel Standard (RFS1) was created under the Energy Policy Act of 2005. The United States Environmental Protection Agency (EPA) released major revisions to the RFS in 2010. The new rule (RFS2) incorporated changes mandated by the 2007 Energy Independence and Security Act (EISA). RFS2 regulates renewable fuels (biofuels) in the US and relies on life cycle analysis (LCA) as a tool to regulate fuels, incorporating greenhouse gas (GHG) emissions from indirect land use change (ILUC) in the calculation. Essentially, the RFS mandates that increasing percentages of U.S. vehicle fuel comes from renewable sources. Additional major modifications were made in 2014 allowing for biogas produced from the anaerobic digestion of municipal wastewater treatment biosolids and refined for use as vehicle fuel to have a preapproved pathway. This has created a very strong incentive for vehicle fuel created from purified biogas. Further information concerning the RFS is included in Appendix D. 3.1.1 THE PRODUCER’S ROLE
The fuel producer, the Water Authority, has flexibility in the roles and responsibilities they can play to participate in markets. Each producer must assess how much direct involvement they desire to reduce costs/maximize profitability. Producer’s responsibilities can be grouped in order to achieve the desired level of involvement. Three simple scenarios are as follows: Scenario 1: (Anticipated cost = 20-25% of Gross Revenue)
• Contract out securing end use • Contract out RIN generation, and RIN sales. • Use intermediaries to handle everything beyond RNG production. Scenario 2: (Anticipated cost = 15-20% of Gross Revenue)
• Contract out securing end use. • Control RIN generation and sales. • Use intermediary for securing offtake. • Responsible for credit generation and compliance.
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Scenario 3: (Anticipated cost = 12-15% of Gross Revenue)
• Control securing end use • Control RIN generation and sales • No intermediaries GAS UTILITY CONNECTION
The Renewable Fuel Standard requires that the RNG be used for transportation fueling to realize its full benefits. The robust United States natural gas “grid” is used to transport the fuels to transportation fueling facilities throughout the nation. The first step in the gas transport process is injection into existing gas pipelines. For the Roanoke RNG facility, this injection will likely be into a pipeline owned by Roanoke Gas Company. Roanoke Gas has a natural gas service line in the vicinity of the treatment facility. Figure 4 indicates the anticipate pipeline for connection and injection of the renewable natural gas. It is anticipated that a finished gas pipeline of approximately 1,600 feet will be needed to convey the RNG from the upgrading site to the finished gas pipeline. There are natural gas service lines closer to the proposed upgrading facility location, but to receive credit in the RFS or LCFS the interconnect must be on a distribution line. During facility design, the interconnect location should be verified with the EPA for acceptance.
Figure 4: Roanoke Gas Co gas pipeline, shown in magenta, in association with the Roanoke WPC Plant
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3.2.1 GAS UTILITY REQUIREMENTS
The gas injected into the existing pipeline will be required to meet certain quality, flow rate, and pressure requirements set in the gas tariff. The interconnecting pipeline system currently operates an 8-inch steel pipe at a maximum pressure of 210 psig. Initial conversations with Roanoke Gas indicate the pipeline has sufficient pipeline capacity to accept the gas produced, but this should be included in future discussions to confirm. The specific gas quality requirements for Roanoke Gas are listed below in Table 10. Two distinctive gas tariffs are included in Table 10 due to Roanoke Gas being fed by both companies. Columbia Gas Transmission feeds the pipeline from the north and East Tennessee Natural Gas feeds the system from the southwest therefore, the gas in the receiving pipeline could be from either gas supply. The upgrading system would need to meet the more restrictive of the two tariffs to guarantee conformance and acceptance. If any of the quality requirements are not met, Roanoke Gas could reserve the right to shut-in, or reject the interconnecting finished gas. It should be noted that the oxygen requirement of 0.02% is very low and has the potential to the challenging to meet. This requirement should be discussed in further detail with Roanoke Gas during additional interconnect discussions.
Table 10: Roanoke Gas Quality Requirements
Limit Most Restrictive Constituent Columbia Gas East Tennessee Limit Transmission, LLC Natural Gas, LLC Hydrocarbon Dew <25 deg F <25 deg F Point NA Hydrogen Sulfide 0.25 grains/100 CF 0.25 grains/100 CF 0.25 grains/100 CF Total Sulfur 2 grains/100 CF 20 grains/100 CF 2 grains/100 CF Combined Total Inert 4% 4% 4% Gases CO 2, N, etc. Oxygen 0.02% 0.10% 0.02% Nitrogen NA 3.00% 3.00% Carbon Dioxide 1.25% 3% 1.25% < 0.05 < 0.05 C6 + Hydrocarbons NA gallons/1000CF gallons/1000CF Wobbe Number 1350 +/- 4% -1400 1298-1400 1350 +/- 4% -1400 Min Heating Value 967 967 967 Max Heating Value 1110 1110 1110 Max Temperature 120 deg F 100 deg F 100 deg F Water Vapor 7 lbs/mmcf 7 lbs/mmcf 7 lbs/mmcf
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To ensure that the quality requirements are being met the finished gas could be monitored using analyzing equipment at the Roanoke Renewable Natural Gas Facility. Roanoke Gas has indicated they do not have standard equipment for monitoring the gas quality prior to injection into the grid. To reach an agreement with Roanoke Gas with a clear understanding of rules and responsibilities would be required prior to completion of the project. 3.2.2 INTERCONNECT STATION
Pipeline and compression facilities prior to the injection point will likely be owned and operated by the biogas upgrade facility owner/operator. The requirements of the interconnect station will be set by Roanoke Gas and may include a measurement facility, pressure valve, telemetry, odorization, remote shut-off valve, and piping. Odorization may or may not be required based on the existing flow rate in the pipe and the proposed gas addition. Currently the gas in the pipeline is odorized at 1 ppm, with odorization done at interstate connection point. As long as the addition of RNG does not substantially decrease the pipeline odor level, odorization would not be needed. The Authority will be required to fund design and installation of the interconnect station. It is anticipated the cost for the interconnect station to be close to $1 million. It is expected that Roanoke Gas will require easements for all facilities located on Authority property. BIOGAS UPGRADING GENERAL DESIGN CRITERIA
Based on a combination of the information presented previously, the facility design criteria are set as indicated in the following sections. The design criteria and described treatment process will be used to determine a high-level probable project cost. 3.3.1 BIOGAS PURIFICATION TECHNOLOGIES
Biogas purification technology alternatives target the separation of CO 2, hydrogen sulfide, VOC and siloxane removal, and/or oxygen and nitrogen removal. The technology alternatives available to remove each type of contaminant are presented in Appendix E. 3.3.2 PHASING STRATEGY APPROACH
If Scenario 1 gas production is pursued it is proposed to install all of the gas purification equipment in a single phase to process up to 250 scfm. Equipment installation associated with gas production scenario 2 it is recommended the equipment be installed in two phases. Phase I equipment would be sized to process 250 scfm of gas with Phase II sized to increase the production to 720 scfm. Construction of additional biogas upgrading capacity beyond Phase I would require the development of a robust FOG/HSW program. An investigation into the locally feedstocks would be helpful in estimating the potential quantities available, as well as the tipping fee range possible in the region. Once the FOG/HSW program feedstock investigation has been performed, a more accurate phasing strategy can be developed. It would be recommended to size pretreatment vessels for Phase II flow. During lower flow periods the vessel bed life will be extended and during high flow times the bed life would be decreased.
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3.3.3 RAW BIOGAS QUANTITY DESIGN CRITERIA
The facility flow is denoted as an installed (Phase I) and future (Phase II and Phase III) flows to allow for the facility to be upsized as gas production increases in the future. Flow design criteria is presented in Table 11. Due to the large difference between the gas produced as indicated by the gas meters and the volatile solids destruction values it is recommended that the design criteria be reevaluated prior to equipment bidding, following meter installation verification and calibration.
Table 11: Facility flow design criteria Installed Minimum Flow 150 Installed Average Operational 2000 and Flow Standard Phase I Scenario I Scenario II Installed Maximum Flow Cubic Feet 250 Future Minimum Flow per Minute 225 Future Average Operational (SCFM)
ario IIario 450 Flow Phase II
Scen Future Maximum Flow 500
The minimum flow would be achieved with system turn down, and potentially also with gas recirculation. Gas recirculation would be utilized if the raw biogas production falls below the treatment system minimum flow and if there was insufficient gas stored in the digesters. Flow monitoring and SCADA control will be vital to making the system operational at all design gas flows. For gas production greater than the maximum indicated flow for each phase, excess gas would either be stored in the digesters or flared. For design purposes it is assumed that the existing flares will be utilized for flaring raw gas. 3.3.4 RAW BIOGAS QUALITY DESIGN CRITERIA
The data presented in Section 1.6 was used to set the raw biogas design criteria, as indicated in Table 12. Prior to full facility design the raw biogas design criteria should be revisited and updated based on additional raw gas sampling that is anticipated to occur. The design criteria below are based on the continuation of ferric being fed to the treatment process upstream of digestion. It is assumed that pressure and temperature are held constant resulting in % Volume being equal to Mole %. This assumption should be verified during facility design. Nitrogen and oxygen values are assumed to be lower than the raw gas sample indicated, but in line with typical municipal values.
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Table 12: Raw gas design criteria Design Constituent Unit Value
CH 4 MOL% 55-65
CO 2 MOL% 34-44
N2 MOL% 1
H2S PPMV 50-200 Oxygen MOL% 0.2 Heating Value (BTU/ft3) 610-630 Total VOCs PPMV 32 Total Siloxanes PPMV 4
3.3.5 FINISHED BIOGAS QUALITY DESIGN CRITERIA
The finished gas design criteria initially has been set by Roanoke Gas as the gas receiving company. Requirements were previously defined in Section 1.5 and included in Table 13 of this section. These requirements are preliminary and subject to change as further discussions with Roanoke Gas occur and as the planning process continues. Increasing the methane capture would have a direct impact on the project capital and O&M costs, but also would increase the projects revenue potential. If the increased revue potential is desired, the finished gas criteria would need to be updated.
Table 13: Roanoke Gas Quality Requirements Constituent Limit Hydrocarbon Dew Point <25 deg F Hydrogen Sulfide 4 ppm (0.25 grains/100 CF) Total Sulfur 32 ppm (2 grains/100 CF) Oxygen 0.02% Nitrogen 3.00% Carbon Dioxide 1.25% Combined Total Inert 4% Gases (CO 2, N, etc.) C6 + Hydrocarbons < 0.05 gallons/1000CF Wobbe Number 1350 +/- 4% -1400 Heating Value 967-1110 BTU/cf Temperature <100 deg F Water Vapor 7 lbs/mmcf
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FACILITY IMPACTS TO COST
Many choices impact the overall RNG project costs. When evaluating a project, it is necessary to consider both the short and long term expenditure impacts. 3.4.1 GENERAL TYPE OF CONSTRUCTION
Many unit process components within biogas upgrading facilities do not require buildings, and equipment manufactures often supply treatment skids with weather proof and/or sound attenuation enclosures. Some facilities house everything in a building, while others use a containerized approach for as much as possible. In Roanoke VA it would be possible to use a containerized system, but heating would be required to keep items from freezing. The use of containers offers significant capital savings and improved constructability. It is anticipated that the bulk of the equipment supplied on this project will be specified to be housed on containerized skids, but that a fraction of the equipment will be shipped loose for installation in a building. Initially, it is assumed that the following types of equipment will be installed in a building: motor control center, control panels, SCADA equipment, chemical feed and storage, tailgas blower equipment (if tailgas below the lower explosive limits and declassified), condensate pumping, chiller glycol piping/valve skids, etc. This equipment would all be “declassified” and installed at least 30-feet away from classified (NEC Class 1 / Div 1 or 2) equipment enclosure skids. Further information on equipment classification is included in the following section. 3.4.2 AREA CLASSIFICATION REQUIREMENTS
Biogas upgrading facility design must consider the potentially explosive and flammable nature of compressed gas in accordance with applicable codes and safety standards. Electrical systems installed as part of the biogas upgrading project must be properly classified based on proximity to biogas piping and equipment. The supply and designation of electrical equipment shall be based on NEC classification as Class 1 Division 1, Class 1 Division 2, or declassified. 1. NEC Class 1 Division 1
• Equipment located within a 5-foot horizontal and 10-foot vertical radius extending from equipment or piping containing biogas shall be designed and supplied based on NEC Class 1 Division 1 requirements. 2. NEC Class 1 Division 2
• Equipment located within a 10-foot horizontal and 15-foot vertical radius extending from equipment or piping containing biogas (either outside or in a building continuously ventilated to 12 air changes per hour, or as required by over-riding local codes) shall be designed and supplied based on NEC Class I Division 2 requirements.
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3. Declassified
• Equipment (not containing or processing biogas) located more than 30 feet from equipment or piping containing biogas shall be declassified. 3.4.3 BUDGETARY CONSTRAINTS
All of the decisions defining the design criteria directly impact the project budget, and the project budget impacts the long-term facility revenue. This project is unique in that it is one of the only projects the Authority will invest in that has a rapid return period. If a more detailed alternative treatment analysis is completed in the future, including analysis from a business perspective, it would be important to consider both life cycle costs AND the associated revenue enhancement associated with each alternative. POTENTIAL GAS TREATMENT TRAIN
It is proposed to utilize a 3-stage membrane, venting the tailgas, to upgrade the anaerobic digester gas to pipeline quality for injection into the natural gas grid.
Membrane treatment requires pretreatment for H2S, VOC and siloxanes. The higher the pretreatment performance targeted, the higher the methane recovery, and the longer the life of the membranes. Figure 5 shows a high-level process flow diagram associated with the proposed upgrading approach.
The pretreatment indicated in the process flow diagram is iron sponge for bulk H 2S removal and carbon adsorption vessels for VOC and siloxane removal. If possible the existing iron sponge and carbon vessels would be used in the facility decreasing the required capital cost. Further evaluation of this possibility would need to be done for full facility design. It is not proposed to include the addition of oxygen due to the low oxygen content in the finished gas tariff, even though the addition of oxygen increases the media capacity for H 2S. Spent media, iron sponge and carbon, would require removal from the vessels and disposal in a landfill. The specific carbon used would be determined based on the amount and speciation of VOC and siloxanes in the raw gas as well as the anticipated H 2S concentration downstream of bulk desulfurization. Due to proposing a non-regenerative pretreatment system, it is suggested to utilize a lead-lag set up. Being a lead lag set up would allow for replacement of the lead vessel media while using the lag vessel to maintain operation through media changes.
Membrane CO 2 separation requires high feed pressures, requiring feed compression. For Phase I it is recommended compression equipment be installed for 250 scfm. When Phase II is installed feed gas compression equipment could be added in the future to increase facility redundancy. In a 3-stage membrane system the methane recovery rate can be designed in excess of 99%, resulting in very low tailgas methane concentrations. The low methane content in the tail gas can be vented directly to the atmosphere in lieu of using a thermal oxidizer, which is required with systems targeting lower methane recoveries. The treatment process would require integration of the 3-stage membrane upgrading process control to the existing facility SCADA controls and condensate captured would be conveyed to the plant.
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At this level of analysis, the upgrading facility is expected to require an area of 0.41 acres. This would include all equipment required for upgrading and the required classification setbacks. Figure 4 in Section 3.2 indicates a potential location for the upgrading facility. The area identified is 0.41 acres and is anticipated to be large enough when reusing the existing H2S, VOC and siloxane removal equipment. If the Authority would like to move the location of the existing equipment additional area will be needed at the upgrading facility.
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1. ANTICIPATE UTILIXING EXISTING COMPRESSION EQUIPMENT AND MAY REQUIRE A PRV.
Figure 5: Proposed upgrading facility process flow diagram
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Capital Costs Based on price quotes provided by vendors, ancillary equipment suppliers bids and quotes for similar projects the capital cost for a 3-stage membrane upgrading process was estimated as $5,626,880 and $9,097,760 for Scenario 1 and 2 respectively, as indicated in Table 14 . All items indicated in Figure 5 are included in the cost estimate, unless noted in Figure 5 as existing or by others. The capital cost estimate also includes sound enclosures on compression equipment skids, installation of equipment, integrated SCADA control, motor control center and site work required for the project. Table 14: Roanoke Regional WRF RNG Production Feasibility Preliminary Opinion of Probable Construction Costs Three Stage Membrane Without Co 2 Reuse Scenario 1 Scenario 2 250 SCFM 720 SCFM 1 Improvements to Existing Facilities $50,000 $200,000 2 Bulk Desulfurization System - Replace Existing Media $50,600 $50,600 3 VOC/Siloxane Carbon Adsorber System - Replace Existing Media $38,000 $38,000 4 Feed Compression W/Enclosures, Chilling & Dehydration $550,000 $750,000 5 Three Stage Membrane System $1,250,000 $2,440,000 6 Product Compression W/Enclosures $300,000 $455,000 7 Gas Quality Check Equipment Skid $140,000 $140,000 8 Equipment Installation (12%) $279,432 $464,832 9 Concrete Foundations for Skids $27,500 $40,000 10 Condensate Handling (Assumes Routing to Existing) $40,000 $40,000 11 Raw Gas Piping (250 LF / 700 LF Assumed) $15,000 $42,000 12 Off-Spec Gas Pipeline Connection to Exist Flares (300 LF Assumed) $18,000 $18,000 13 Finished Gas Pipeline (1,400 LF Assumed) $105,000 $105,000 14 Building (Declassified; Electrical and Control; 350 SF) $70,000 $70,000 15 Fire Protection Pipeline and System (Assume 300 LF) $19,000 $19,000 16 Electrical & Controls $150,000 $250,000 17 Mechanical Piping, Insulation, Appurtenances $150,000 $250,000 18 Motor Control Center, VFDs, Starters, Soft Starters $150,000 $225,000 19 Control Monitoring Package $80,000 $80,000 20 SCADA Integration into Existing System $45,000 $45,000 21 Miscellaneous Site Work $60,000 $80,000 22 General Contractor Overhead and Profit (12%) $431,000 $696,000 Sub-Total $4,019,000 $6,498,000 Contingency (25%) $1,005,000 $1,625,000 Sub-Total Construction $5,024,000 $8,123,000 Engineering (12%) $602,880 $974,760 Total Project Costs $5,626,880 $9,097,760
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Operation and Maintenance Most of the operational costs associated with a 3-stage membrane upgrading facility are associated with pretreatment, dehydration, compression and replacement of the membranes every 7-10 years. The yearly operational cost for the process is estimated as $1,377,591 in the first year and $1,431,293 in the 20 th year of operation, as indicated in Table 15 with additional information in Appendix F . This cost does not include ferric addition to the treatment process, as that is currently occurring and would not change due to the installation of an upgrading facility. The cost does include the addition of one full time equivalent operations position at an hourly cost of $50. The operation and maintenance costs include replacement natural gas for the existing biogas for heating the digesters. Yearly gas costs were based on the equivalent of 120 scfm at a heating value of 650 btu/scfm and a cost of $3/dekatherm. Remote monitoring of the 3-stage membrane treatment process would be possible through a service contract with the membrane equipment supplier. The service contract can be 24/7 service with phone support regardless of the time. Onsite service visits also are possible with DMT and with many other membrane treatment process vendors. The cost associated with a service plan depends on the level of service desired and can be discussed with the chosen equipment vendor.
Table 15: Roanoke Regional WPC RNG Production Feasibility Opinion of Probable O,M&R Cost Summary Scenario 1 Scenario 2 2012 $519,113 $802,182 2039 $572,815 $855,884 PW O,M&R $7,349,956 $11,197,797
Operator maintenance and operations tasks primarily pertain to operation of the pretreatment processes, which the facility operators are currently familiar with, as the membranes generally require little maintenance. Table 16 indicates operation and maintenance requirements associated with the equipment associated with the treatment steps that comprise the 3-stage membrane upgrading process.
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Table 16: DMT 3 Stage Membrane with Vent O&M Requirements Treatment Step Daily Weekly Monthly Quarterly Annually Media Water Return replacement Iron Sponge SCADA Verification – Visual Inspection Visual Inspection 2x per year H2S Removal Checks Adjust Water of Equipment of Equipment (Approx 32 Timer as Needed hrs) Media replacement VOC Carbon SCADA Visual Inspection Visual Inspection Visual Inspection 2x per year Adsorption Checks of Equipment of Equipment of Equipment (Approx 32 hrs) Replacement Check Gas of Gas Filters, SCADA Visual Inspection Filters, Oil Filters Visual Inspection Compressors Oil Filters, Checks of Equipment and Control of Equipment And Oil Calibration Change Visual 3 Stage 2 Stage Visual Inspection Visual Inspection SCADA Checks Inspection of Membranes Membranes of Equipment of Equipment Equipment
3.5.1 ENGINEER’S PROFORMA EVALUATION
A high level proforma evaluation was completed to evaluate potential project revenue. Revenue projections, over a 20-year period, were based on two scenarios. Option 1 is for producing all D3 gas due to only receiving municipal biosolids at the facility and producing a maximum of 250 scfm of gas. Option 2 is based on the facility accepting FOG and HSW resulting in a mixture of D3 and D5 gas produced with a maximum production of 720 scfm within 5 years of operation. A consistent rate of inflation, for both revenue and costs, throughout the project life was applied. RIN management was set to 15% of the environmental credits to reflect the cost of participating in the RFS. Distribution costs were set to $0.28/DTH delivered to capture the costs associated with transporting the gas to partake in the RFS. The proformas have been completed assuming a high level of risk, selling RINs on a daily basis, to maintain or capture a larger portion of the gross revenue. Figure 6 shows a comparison of net present worth for a facility producing 250 scfm of D3 gas vs a facility producing 720 scfm of a mixture of D3 and D5 gas. Appendix G includes proforma evaluations for both scenarios.
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Comparison of Net Present Worth 250 scfm vs 720 scfm $100,000,000 $80,000,000 $60,000,000 $40,000,000 $20,000,000 $0 250 SCFM 720 SCFM -$20,000,000 Capital Cost OM PW Revenue PW Net PW
Figure 6: Anticipated net present worth for two options for gas production, all D3 gas vs mixed D3 and D5 gas. The cumulative net present value associated with the above described 3 stage membrane with vent upgrading process was estimated as $13,246,230 for Scenario 1 and $56,614,938 for Scenario 2. The estimation is based on mid-range RFS revenue, not selling gas into the LCFS and a 20-year period. However, the Authority may elect to use different revenue projections based on several factors including:
• Future projections given historical data. (High, Mid, Low)
• Appetite for risk.
• Roles and responsibilities required to participate in environmental credit programs. The following sections discuss the various factors the Authority must consider establishing a realistic revenue projection that can be used to further evaluate the upgrading process. Future Projections Given Historical Data The general trendline for the D-6 RIN has been on the increase since the inception of the program and can be seen in Figure 7. However, this graph clearly indicates the volatility of the market over the time period shown in the figure.
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Figure 7: General trendline associated with D6 RIN (from EcoEngineers) Figure 8 represents market price of all RIN D-Codes over the last 3 ½ years. As can be seen, the D3 RIN sold between $1.10 and $3.10 during that period. While D3 RINs have had a relatively short market history compared to D6 RINs, the value generally parallels the D6 RINs, apart from notable large upswings in price during the first quarter of 2016 and again in the first quarter of 2017. The upswings are the result of the Cellulosic Waiver Credit (CWC) being set at much higher prices than the year prior due to significant shortfalls in the actual volume of fuel produced compared to the established mandates for cellulosic-based fuels and given the average fuel price at the time that the CWC was established. The price of a D3 RIN equals the price of D5 RIN plus the value of the CWC. Revenue projections were evaluated given Mid RIN price assumptions, but could be estimated based on high or low RIN value estimations based on the current market. Anticipated High, Mid and Low D3 RIN prices are plotted on Figure 8 to indicate where those price assumptions compare to historical data.
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March 7, 2019
Figure 8: RIN values 2015-2018 (from EcoEngineers) Appetite for Risk
Additionally, Roanoke Authority must evaluate the proforma while considering the Authority’s own appetite for risk. Many different options are available when it comes to marketing and selling environmental credits. Entities that are willing to take on more risk also have the potential for more upside. If the Authority is willing to sell at market prices monthly, it will be able to realize a higher percentage of the gross revenue generated. However, this subjects the project to potential large market swings as seen in Figure 8. Long-term contracts of 3 to 10 years are available. Such contracts would ensure the sale of environmental attributes at a constant or near constant price. Entering into a long-term contract may insulate the County from occasional market downturns or even catastrophic market collapse. However, these contracts typically demand that the environmental attributes be sold at a very modest price; a significant discount from the current market value.
36
Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility Study
March 7, 2019
4 IMPLEMENTATION STEPS
Considerable work remains in the development of this project. Some of the next steps anticipated at this time are outlined below.
• Rehab the digesters; • Install mixing in the primary digesters; • Address digester low heating times; • Verify existing gas meter installation, calibrate meters and verify gas production values; • Identify thickening adjustments and or upgrades; • Complete raw gas sampling; • Revisit and verify proposed design criteria following additional data gathering; • Determine future plans to receive additional domestic waste, municipal biosolids or FOG/HSW, including amounts and timing to accept waste; • For a biogas project to move forward: o Begin utility connection agreement negotiation; o Procurement of equipment through competitive bidding; o Selection of direction on gas offtake, environmental attribute offtake, and responsibilities and risks deemed palatable for the Authority; o Initial regulatory meetings; o Carbon intensity score modeling;
37
Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility Study
March 7, 2019
5 SCHEDULE
38
Appendix A Roanoke WPC Plant General Information
Appendix A – Roanoke WPC Plant General Information Page 1
There are three distinct aeration treatment trains at the facility, as shown in Figure 1. Each aeration train operates with different biology and the trains are referred to as train “A”, “B” or “C” as described below:
• Train A – 10 square clarifiers with corner sweeps and 10-foot side wall depth • Train B – six square clarifiers with corner sweeps and 15-foot side wall depth • Train C - dosed with ferric chloride and consists of two circular clarifiers
Figure 1 - Plant Layout Aeration initially was provided to the trains via two Roots 1,750 HP centrifugal electric blowers. Each blower is capable of producing 33,788 standard cubic feet per minute (scfm). The blowers operate on 4,160 V, the voltage is stepped up from 480V at the facility. Only one blower is needed at a time to provide the facility aeration needs at normal facility flow rates. The WPC Plant also has a 1,750 HP backup blower, that runs on diesel fuel. There have been operational challenges with the 1,750 HP blowers providing too much aeration resulting in dissolved oxygen (DO) values of 5-6 milligrams/Liter (mg/L) leaving the aeration basins. To address this, two Neuros 700 HP turbo blowers were installed to take the place of the 1,750 HP blowers. Each Neuros blower delivers 14, 750 scfm. A 250 HP blower is now used as a backup blower rather than the 1,750 HP diesel blower. The blower replacement was required because there are not automated valves to control air flow to each basin and zone of the basin. Wet weather improvements to the collection system and treatment process were completed in 2007 to protect the facility and environment from system overflows due to increased system flow. Further facility improvements were completed in 2015 including a new chlorine contact basin, new effluent Appendix A – Roanoke WPC Plant General Information Page 2
screw pumps and modifications to the Biological Aerated Filter (BAF). The effluent screw pumps allow the facility to discharge to the Roanoke River during high river flow events. The modifications to the BAF allow for operation as a one or two stage treatment process. A liquid process flow diagram for the Roanoke WCP Plant is shown in Error! Reference source not found. and a solids process flow diagram is shown in Error! Reference source not found.. Both of these figures are from facility upgrades completed in 2004.
Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility
March 4, 2019
Figure 2: Roanoke WPC Plant Liquids Process Flow Diagram
6
Western Virginia Water Authority Roanoke Regional WPC Renewable Natural Gas Production Feasibility
March 4, 2019
Figure 3: Roanoke WPC Plant Solids Process Flow Diagram
7
Appendix B Gas Sampling Protocol
Appendix B – Gas Sampling Protocol Page 1
Recommendations for Roanoke Digester Gas Characterization
The below table summarizes the proposed analytical testing methods and sampling frequency to determine design values for the gas upgrading process. Most anaerobic digesters have SRTs of 15 to 30 days, hence spacing the samples out over a period of weeks is ideal. The below testing regimen would provide design values for hydrogen sulfide, volatile organic compounds, siloxanes, and ammonia.
Cost per Analysis Constituent Quantity Sampling Frequency Total Analysis Method BTU, H2s, Total Sulfur, Nitrogen, Carbon Dioxide, 350 10 1 to 2 samples per week 3,500 ASTM 3588 Carbon Monoxide, Methane, Oxygen VOCs, TICs, Siloxanes 300 10 1 to 3 samples per week 3,000 EPA TO-15M Ammonia 50 5 1 to 2 samples per week 250 Draeger Tube Summa Cannisters 35 10 NA 350 NA 1 Estimated Lab Costs $7,100 1. Cost estimate based on previous work with AAC labs, actual costs may deviate based on quote from AAC.
AAC labs provides the below manifold and summa container for the customer to take the sample. Sampling Manifold
Appendix B – Gas Sampling Protocol Page 2
Summa Container
Appendix B – Gas Sampling Protocol Page 3
Notes on sampling method
1. Instruction provided by AAC start on the next page of this document.
2. Source gas must be below 15 PSIG.
3. The sampling manifold (first picture above) provides a 3 way valve that allows the tubing to be evacuated of air. It is critical to insure no air enters the samples as it will elevate critical process design criteria (Nitrogen and Oxygen).
4. The sample manifold has an orifice plate located on the cannister side of the 3 way valve in order to avoid a vacuum condition in the sample tubing. A gauge is provided on the sampling manifold to show that the container is under vacuum, and as an indication when the container is full.
5. Nitrogen purge connections on digester gas piping with FNPT threads are ideal for sample connections. End user will need to provide threaded bushing(s) to increase the provide 3/8” MNPT threaded fitting to match the size of the purge connection.
Appendix B – Gas Sampling Protocol Page 4
Sampling Instructions: (Positive Pressure Source Gas)
Equipment Summary: You have been provided with six mini Summa canisters and a gas sampling manifold.
Sampling Summary: To prepare the sampling system, first remove the brass cap on top of the summa canister valve. Connect the blue taped end of the manifold to this valve, hand tighten at first then finish with a 9/16 wrench. Make sure the gas pressure does not exceed 15psi.
Record the sample information on the tag attached to the canister. Connect the provided tubing coming from the source sampling port to the green taped port of the manifold.
Turn on the gas flow from the source sampling port and open the valve on the red taped end of the manifold. Gas will flow through the line and manifold and out the red taped valve. It will not begin to sample until the canister valve is opened.
Purge the line and manifold for 1 minute and then close the valve on the manifold. The system will then start building pressure which will show on the gauge. Allow the system to equilibrate with the source gas pressure for 1 minute. Next, close the source sampling port valve, the system is now isolated and under pressure. Carefully watch the gauge to ensure that the pressure does not drop over the next minute. If it does, there is a leak in one of the connections, in which case, check/tighten all connections and repaeat the process until it passes the leak check.
Once the system passes leak check, open the source sampling port valve again and flush the line and manifold for 1 more minute by slightly opening the manifold valve. Once finished, close the manifold valve and open the valve on the Summa canister (the gauge on the manifold will read -30”Hg once the canister is opened). The canister will then start to fill and you will be able to monitor the progress on the gauge as it does.
To stop the flow at any time simply close the source sampling port valve and canister valve. The gauge at this point will indicate the remaining vacuum in the canister. The canister can be filled up to 1 or 2 psi without being considered pressurized.
When the sampling is completed, disconnect the manifold and tubing from the canister. Replace the brass cap on top of the canister and tighten.
Please call Marcus at 805-650-1642 if there are any questions.
Appendix C Introduction to the Renewable Fuel Standard (RFS)
Appendix C – Introduction to the RFS Page 1
INTRODUCTION TO RFS
When a volume of renewable fuel is produced, it is assigned a Renewable Identification Number (RIN) is assigned to track production, use and trade. According to the Environmental Protection Agency (EPA), compliance in the program is achieved when Obligated Parties, refiners or importers of petroleum, either blend renewable fuels into transportation fuel or obtain RINs to meet their EPA- specified obligation and avoid penalties. Multiple RIN types are produced, depending on percent reduction of GHG produced in the creation and consumption of the fuel as compared to that of traditional petroleum-based fuels. By year 2022, the RFS mandates 20 percent, or 36 billion, Figure 1 from the Environmental Protection Agency (EPA), indicates the mandated amounts on a yearly basis.
Figure 1: Mandated increase in renewable fuels based on D5, D3 and D6 feedstock material. Figure from the EPA.
The bottom portion of Figure 1, in dark blue, represents corn-based ethanol. On the right side of the graph, a D6 credit is associated with the fuel representing a 20-percent reduction in GHG emissions from traditional petroleum-based fuels. The upper section, in red, indicates advanced biofuels and is associated with the production of biogas from fats, oils and grease (FOG) and sugars. This fuel qualifies for a D5 credit, representing a 50-percent reduction in GHG. The light blue portion of the graph represents cellulosic fuels, such as RNG produced from landfill, biogas produced from the anaerobic digestion of manures, and municipal wastewater biosolids. These fuels qualify for a D3 RIN credit which represents a 60-percent reduction in GHG emissions. The two types of RIN Classifications that could apply to wastewater biogas include D3 and D5 RINs. A cellulosic biofuel, identified as a D3 RIN, represents renewable natural gas produced from cellulose, hemicellulose, or Appendix C – Introduction to the RFS Page 2
lignin. An advanced biofuel, identified as a D5 RIN, represents renewable natural gas produced from digestion of any type of renewable biomass except cornstarch ethanol. By EPA regulation (40 CFR part 80 subpart M), renewable natural gas derived from municipal facility digesters are automatically given the D3 fuel pathway and qualify as such. The value of a D3 RIN is roughly three times the value of a D5 RIN, so strategies used to optimize the D3 RIN classification are important when attempting to optimize revenue. The EPA continues to evolve feed stock interpretation of what constitutes D3 vs. D5 within a wastewater treatment plant biogas application. All municipal biosolids-generated biogas is considered D3 by the EPA. If a given digester or digestion system includes the addition of HSW, it becomes more difficult for the EPA to delineate RIN classifications. The EPA would prefer to consider any digester fed with both cellulosic and non-cellulosic materials as D5 only. Figure 2 shows how gas and RINs typically flow. This depiction is very general with the roles of Energy Service Provider, End User, RIN Manager, Broker and Obligated Party often combined, eliminating the number of transactions required in order to maximize efficiencies and profit. Biogas from an anaerobic digester is purified, compressed and injected into the natural gas pipeline, which could be virtually anywhere in the country, as shown in Figure 3. The Energy Service Provider takes control of the fuel and delivers it to the End User. The Energy Service Provider would typically separate the RINs and market those independently from the fuel. The Energy Service Provider may sell the RINs to a RIN Manager which would verify them as valid RINs and sell the RINs to a Broker. The Broker would typically buy RINs from multiple sources from across the country, aggregate those RINs and sell in large bundles to the Obligated Party.
Figure 2: Typical pathway for biogas upgrading and sale to obligated parties
Appendix C – Introduction to the RFS Page 3
Figure 3: Interstate pipelines within the United States
Figure 3 indicates the vast network of Interstate pipelines within the US. This figure does not include intrastate pipelines or local distribution lines. As long as physical connection can be verified through the natural gas pipeline network, a contractual arrangement can connect a source with virtually any end user in the country. A mass balance is used at each pipeline connection point along the route to indicate final usage, while the physical gas molecules are consumed locally. 1.1.1 INTRODUCTION TO LCFS
In addition to participating in the RFS, municipalities can sell gas into the state of California, Oregon or British Columbia to take advantage of a state-based credit system. California’s Low Carbon Fuel Standard (LCFS) was approved by the state in 2009 and implementation began on January 1, 2011. The regulation is one of the nine discrete early action measures to reduce California’s greenhouse gas emissions that cause climate change. This program is part of a comprehensive plan to help California meet its goal of reducing greenhouse emissions to 1990 levels by the year 2020. The LCFS program is based on the principle that each fuel has "lifecycle" greenhouse gas emissions that include CO 2, N 2O, and other greenhouse gas contributors. The lifecycle assessment includes direct emissions associated with producing, transporting, and using the fuels, as well as significant indirect effects on greenhouse gas emissions, such as changes in land use for some biofuels. Participation in the LCFS requires the producer to contract for end use within the State of California. Credits are determined by comparing the renewable fuel’s carbon intensity (CI) to that of petroleum- Appendix C – Introduction to the RFS Page 4
based fuels. An extensive audit of the project is required in order to determine the fuel’s CI and ultimately the value of the credits associated with the production of the fuel. While the program is administered completely separately from the RFS, the credits are stackable. A fuel producer can participate and generate revenue with both programs simultaneously. Other jurisdictions are following California’s footsteps, which is evident in the Pacific Coast Collaborative, a regional agreement between California, Oregon and British Columbia, to strategically align policies to reduce greenhouse gases and promote clean energy. The California Air Resource Board has been routinely working with these jurisdictions. Over time, these LCFS programs will build an integrated West Coast market for low-carbon fuels that will create greater market pull, increase confidence for investors of low-carbon alternative fuels, synergistic implementation and enforcement programs. Washington is currently considering adopting a similar state-based credit system as is a consortium of eight states in the northeast.
Appendix D Suez Cannon Mixing System Quotation
BUDGET PROPOSAL Project : Roanoke WWTP Engineer : Bartlett & West Proposal No : 355744 Date : February 21, 2019
Regional Business Manager Local Representative Mr. Paul Ravelli Mr. Jon Casarotti SUEZ Sherwood-Logan & Associates 8007 Discovery Drive 9710 Farrar Court, Suite O Richmond, VA 23229 Richmond, VA 23236 Ph: 856-761-2407 804-357-2292 [email protected] [email protected] TABLE OF CONTENTS
• PROCESS DESCRIPTION
• DESIGN SUMMARY
• SCOPE OF SUPPLY – BY SUEZ
• SCOPE OF SUPPLY – BY OTHERS
• BUDGET PRICING
• TERMS & CONDITIONS OF SALE
• REFERENCES
• BROCHURE
1
PROCESS DESCRIPTION
Cannon® Mixer
The Cannon® Mixer system is a process in which an anaerobic digester is completely mixed by recirculating the biogas through a cannon. The Cannon® Mixer is a proven anaerobic digester mixer. Using large distinct bubbles, the Cannon® Mixer combines integral mixing and heating within one tank. The result is a reduction of volatile solids and sludge volume, greater overall gas production, and improved sludge dewatering.
Peak digester performance and reduced operating costs are directly associated with Suez Cannon® Mixer. Cannon® Mixer uses large piston bubbles - proven as the most efficient use of energy for fluid displacement - to agitate digester contents thoroughly and economically. Bubble generation every three to four seconds per mixer guarantees better than 90 percent active volume in the digester. Ideal for deep-tank mixing, the system is not only superior in performance but is also easy to install, operate, and maintain.
The operating principles behind Cannon® Mixer have been consistently proven in hundreds of installations. Cannon Mixer fits well within a variety of applications, from retrofitting existing digesters to optimizing two-phase anaerobic digestion systems.
Cannon’s proven mixing process: • A vertical stack pipe, open at each end and varying in length according to digester depth, is the central component of the Cannon Mixer. Based on computerized modeling, multiple units are strategically arranged to optimize mixing zones across the entire floor area - achieving more than 90 percent total active volume. • Recirculated gas is continuously fed to the bubble generator and intermittently discharged into the stack pipe as a large piston bubble. • The piston bubble fills the entire cross section of the pipe, driving out liquid as it rises and creating a siphon. As one bubble leaves the stack pipe at the top, another enters from the generator for both continuous mixing and prevention of solids settling. • Large bubbles burst as they leave the liquid surface, creating substantial turbulence that prevents scum buildup.
Advantages to Cannon® Mixing
Cost effective – reduction in O&M and solids handling costs Environmentally friendly – low energy consumption and increased biogas production Easy to install, operate and maintain Complete mixing (90% active volume) Less scum formation, less solids settling, superior volatile destruction, and more biogas production
2
DESIGN SUMMARY
The proposed system is based on the following design conditions:
GENERAL INFORMATION†
Total Flow to Digesters 134,000 gpd Sludge Type to Digesters Combined Primary & WAS Sludge Concentration 3.00% Estimated Viscosity of Sludge in Digester 300 cps at shear rate of 6 seconds-1 Average Retention Time Please Advise Mixing Requirement Minimum active volume of 90% Digester Temperature 95OF ± 1OF at any point in tank
† The proposed design is preliminary and based on the above information. Final design conditions must be defined for SUEZ to confirm the proposed design. Please advise SUEZ with any changes to the design criteria.
ANAEROBIC DIGESTER INFORMATION
Number of Digesters 1 existing Diameter 90 ft Maximum Sidewater Depth 22 ft Minimum Sidewater Depth 18 ft* Depth of Cone 11.25 ft*
*Assumed. Used height of cone ratio of 1:4.
DESIGN OF CANNON® MIXING SYSTEM
Number of Cannon Mixers 10 Diameter of Mixer 24 in. Mounting Floor Hydraulic Length of Mixer 14.29 ft Pumping Rate through each unit (at max liquid 3,890 gpm level and estimated viscosity) Total Mixing per digester 38,900 gpm
3
DESIGN SUMMARY
Turnover Rate (at max liquid level) 31.5 min Compressor Capacity 364 scfm Discharge Pressure at max liquid level 11.2 psig
4
SCOPE OF SUPPLY – BY SUEZ
SUEZ proposes to furnish the following equipment and services for the Cannon® Mixing System:
QTY ITEM
Cannon® Mixers, each consisting of a non-clog piston bubble generator, draft tube, heating jacket, floor support bracket and fittings. Each mixer will be fabricated of mild steel, 1/4” 10 minimum thickness, except the bubble generator which is fabricated of Type 304 stainless steel, 1/8” thickness. Carbon steel will be prepared by sandblasting followed by two shop coats of epoxy paint. Compressor assembly to be located in a digester building. Compressor assembly will include the following components shipped loose for field assembly by others: liquid ring rotary 1 compressor with explosion proof motor, guard, baseplate, discharge moisture separator, inlet flame arrester, inlet sediment trap seal water line accessories, discharge check valve, inlet and outlet pressure gauges and high/low pressure safety controls. Wall mounted control panel, NEMA 7, consisting of combination starter and circuit breaker, 1 and system operating lights and switches. Gas flow balancing system consisting of one gas flow meter, one pressure gauge, one 1 balancing valve and isolation valves for each mixer. The components for the gas flow balancing system will be shipped loose for field assembly and piping by others. Thermostat to be inserted into the digester to control temperature and heating cycles of the 1 digester. LOT Freight to the jobsite LOT Electronic O&M Manuals
Field Service Days of a SUEZ field service representative for installation inspection, 8 commissioning and training in no more than two (2) trip(s). Additional days of field service are available on a per diem basis at $1,500/Day, plus travel expenses
5
SCOPE OF SUPPLY – BY OTHERS
The following items, but not limited to, shall be provided by Others:
• Installation of any kind and unloading & placement of equipment.
• All concrete and civil works of any kind.
• All anchor bolts and mounting hardware not specified herein.
• All piping and piping supports.
• All civil, mechanical, electrical and plumbing works.
• It is understood that the existing site has hot water/heat exchangers available for the heating jackets on the mixers. Boiler, heat exchangers and hot water pumps by others.
• All other necessary equipment and services not otherwise listed as specifically supplied by SUEZ.
6
BUDGET PRICING
BUDGET PRICING
Scope of Supply – By Suez $469,000 USD
TERMS & CONDITIONS OF SALE
All budget pricing is based on SUEZ’s standard terms & conditions, which can be provided upon request. This proposal is being provided for preliminary estimating purposes and is a non-binding offer. SUEZ reserves the right to update our scope of supply and cost due to market escalation, changes in the process approach or updated information regarding the design influent/effluent characteristics. SUEZ will not be responsible or liable in any manner for these costs until such time of our mutual execution of a definitive written agreement.
FREIGHT TERMS
FOB Jobsite
PAYMENT TERMS
10% Net Cash, Payable in thirty (30) days from date of submittal of initial drawings for approval
Net Cash, Payable in progress payments thirty (30) days from dates of respective shipments of the 85% Products
Net Cash, Payable in thirty (30) days from Product installation and acceptance or sixty (60) days after 5% date of final Product delivery, whichever occurs first
PRELIMINARY SCHEDULE
Submittals 4-6 weeks following a fully executed agreement
Delivery 14-16 weeks following submittal approval
7
REFERENCES
8
Cannon Mixer Installations
Contract number Contract city Name of Facility Industry Start date Number of units Size Number of digesters Diameter of digesters (in) (ft) Plants Mixers State: Alabama 5 80 2785 Anniston Anniston Municipal 1987 6 18 2 60 2794 Mobile Three Mile Creek Municipal 1988 14 18 2 80 89-448 Ensley Village Creek Municipal 1990 27 24 3 100 94-225 Mobile Clifton C. Williams Municipal 1995 15 30 3 68 98-699 Ensley Village Creek Municipal 1998 18 24 2 100 State: California 17 257 2620 Redding Redding Municipal 1979 10 12 2 70 2674 Roseville Roseville Municipal 1982 5 12 1 90 2691 San Diego San Diego Municipal 1981 9 12 1 125 2763 Tracy Tracy Municipal 1986 10 24 2 75 2769 Burlingame Burlingame Municipal 1986 7 24 2 1 @ 46 - 1 @ 60 2771 San Diego San Diego Municipal 1986 9 18 1 125 2799 Santa Rosa Santa Rosa Municipal 1988 5 24 1 75 90-502 Chino basin Carbon Canyon Municipal 1992 18 24 2 90 91-675 Roseville Roseville Municipal 1992 18 24 2 90 91-711 South San Francisco South San Francisco Municipal 1991 18 30 3 2 @ 70 - 1 @ 68 92-866 Chino basin Chino basin Municipal 1993 21 30 3 2 @ 68 - 1 @ 65 94-184 Rialto Rialto Municipal 1995 7 30 1 75 94-250 Santa Rosa Laguna Subreg WRF Municipal 1996 7 30 1 75 94-258 San Diego Point Loma Municipal 1995 32 30 2 125 96-438 San Diego Point Loma Municipal 1997 17 30 1 125 00-015 San Diego Point Loma Municipal 2000 32 30 2 125 03-382 San Diego San Diego Municipal 2004 32 30 2 125 State: Colorado 7 85 2751 Loveland Loveland Municipal 1985 8 18 2 60 2786 FT Collins FT Collins Municipal 1987 8 24 1 75 91-652 Denver Denver Municipal 1992 24 24 2 100 94-186 Denver Metro Municipal 1995 24 24 2 100 97-541 FT Collins FT Collins Municipal 1998 12 30 2 75 99-848 FT Collins FT Collins Municipal 1999 6 30 1 75 02-273 South Adams County South Adams Municipal 2002 3 24 1 45 State: Connecticut 3 12 90-590 Seymour Seymour Municipal 1991 2 24 2 1 @ 35 - 1 @ 24 98-739 Fairfield Fairfield Municipal 1998 4 30 1 60 06-556 Milford Housatonic Municipal 1998 6 30 1 75 State: Florida 7 143 2762 Orlando Orlando Municipal 1986 12 18 2 75 2789 Boca Raton Boca Raton Municipal 1988 8 30 2 80 2806 Orlando Iron Bridge Municipal 1989 48 24 4 100 89-394 Pompano Beach Broward County Municipal 1990 25 30 5 100 96-453 Boca Raton Boca Raton Municipal 1997 7 30 1 80 98-746 Pompano Beach Broward County Municipal 1998 27 30 3 100 06-535 Glendale Lakeland - 2PAD Municipal 2008 16 24 2 85 State: Idaho 4 14 2656 Caldwell Caldwell Municipal 1980 2 12 1 35 2754 Blackfoot Blackfoot Municipal 1985 4 12 2 40 89-403 Lewinston Lewinston Municipal 1990 4 18 1 50 91-633 Pocatello Pocatello Municipal 1992 4 30 1 65 State: Illinois 11 216 2741 Chicago MSDof Greater Chicago Municipal 1986 120 18 18 12 @ 110 - 6 @ 120 2749 Chicago West Chicago Municipal 1986 6 18 2 55 2767 Addison Addison Municipal 1986 6 18 2 45 2778 Plainfield Plainfield Municipal 1987 4 18 1 50 2788 Charleston Charleston Municipal 1987 3 12 1 35 89-430 Elgin Elgin Municipal 1989 8 5 @ 24 - 3 @ 18 2 1 @ 90 - 1 @ 65 90-559 Batavia Batavia Municipal 1991 6 18 2 40 96-496 Mundelein Mundelein Municipal 1997 7 30 1 85 99-867 Schaumberg Egan Municipal 2000 6 30 1 75 00-069 Rockford Rock River Municipal 2001 24 24 3 75 02-249 Bloomington Bloomington-Normal WRD Municipal 2003 26 30 2 110 State: Indiana 13 69 2519 Brookville Brookville Municipal 1976 1 18 1 32 2592 New Castle New Castle Municipal 1980 8 12 3 2 @ 45 - 1 @ 32 2661 Elkhart Elkhart Municipal 1981 8 12 1 110 2688 Batesville Batesville Municipal 1983 2 12 1 35 Cannon Mixer Installations
Contract number Contract city Name of Facility Industry Start date Number of units Size Number of digesters Diameter of digesters (in) (ft) Plants Mixers 2689 Marion Marion Municipal 1984 8 12 2 65 2713 Valparaiso Valparaiso Municipal 1984 8 12 2 1 @ 65 - 1 @ 60 2725 Decatur Decatur Municipal 1984 3 12 1 50 2746 Fall Creek Fall Creek Municipal 1985 3 16 1 50 2772 Zionsville Zionsville Municipal 1987 2 18 2 20 2780 Vincennese Vincennese Municipal 1987 4 18 1 55 2804 Terre haute Terre haute Municipal 1988 12 24 2 65 89-390 Jasper Jasper Municipal 1990 3 18 1 50 93-034 New Castle New Castle Municipal 1993 7 24 3 2 @ 45 - 1 @ 32 State: Iowa 2 3 2790 Altoona Altoona Municipal 1987 2 18 1 24 90-516 New Hampton New Hampton Municipal 1991 1 30 1 30 State: Kentucky 1 6 2634 Morehead Morehead Municipal 1979 6 12 2 45 State: Louisiana 6 54 2774 Baton Rouge Baton Rouge South Municipal 1988 16 24 4 65 2793 Baton Rouge Baton Rouge North Municipal 1988 8 24 2 65 89-429 Slidell Slidell Municipal 1989 5 24 1 50 89-439 Baton Rouge North CSD Municipal 1989 4 24 1 65 94-226 Baton Rouge South Municipal 1995 9 24 1 90 94-244 Baton Rouge Central Municipal 1995 12 30 3 65 State: Massachusetts 2 27 2653 Northampton Northampton Municipal 1981 3 12 1 45 2722 Boston Deer Island Municipal 1985 24 18 4 108 State: Michigan 2 16 90-523 Midland Midland Municipal 1991 8 24 2 60 07-592 Delhi Township Delhi Municipal 2008 8 18 4 2 @ 19 - 2 @ 40 State: Minnesota 2 32 2601 Lakeville Lakeville - Farmington Municipal 1979 12 12 4 65 90-595 Rochester Rochester WTR Recla Municipal 1992 20 30 2 90 State: Montana 2 10 2734 Laurel Laurel Municipal 1985 6 12 2 45 2753 Lewiston Lewiston Municipal 1986 4 12 2 30 State: Nebraska 1 21 93-099 Omaha Papillon Creek Municipal 1994 21 30 3 75 State: New Jersey 11 137 2796 Vineland Vineland Municipal 1988 16 24 2 70 90-577 Little Ferry Bergin County Municipal 1992 10 24 1 80 92-862 Joint Meeting Joint Meeting Municipal 1993 10 30 1 95 94-240 Ocean County Ocean County Municipal 1995 6 30 1 85 98-645 Little Ferry Bergen County UA Municipal 1999 40 24 4 80 98-648 Joint Meeting Joint Meeting Municipal 1999 10 30 1 95 99-827 Joint Meeting Joint Meeting Municipal 2000 10 30 1 95 00-075 Township of Middletown Middletown Municipal 2001 4 30 1 65 02-259 Rahway Rahway Valley SA Municipal 2003 7 30 1 75 03-380 Elizabeth Joint Meeting Municipal 2004 10 - 1 95 03-381 Rahway Rahway Valley SA Municipal 2004 14 30 2 75 State: New York 30 179 2588 Grand Island Grand Island Municipal 1978 4 12 1 60 2622 Massena Massena Municipal 1979 4 12 2 40 2646 Canandaigua Canandaigua Municipal 1979 8 12 4 35 2649 Phelps Phelps Municipal 1980 2 12 1 30 2659 Warsaw Warsaw Municipal 1981 2 12 2 26 2662 Oneida Oneida Municipal 1981 3 12 3 28 2684 Rotterdam Rotterdam Municipal 1982 2 12 2 20 2687 Lackawanna Lackawanna Municipal 1983 1 18 1 50 2726 Geneva Geneva Municipal 1984 6 12 2 50 2727 Alden Alden Municipal 1985 2 18 2 1 @ 20 - 1 @ 31 2733 Jamestown Jamestown Municipal 1985 10 12 2 75 2736 Attica Attica Municipal 1984 2 18 2 25 2744 Medina Medina Municipal 1984 2 18 1 30 2755 Akron Akron Municipal1985 1 12 1 25 2760 Newark Newark Municipal 1986 4 18 2 40 2761 Hornell Hornell Municipal 1986 12 18 2 40 2779 Hudson Falls Hudson Falls Municipal 1988 6 12 2 45 Cannon Mixer Installations
Contract number Contract city Name of Facility Industry Start date Number of units Size Number of digesters Diameter of digesters (in) (ft) Plants Mixers 2791 Springville Springville Municipal 1987 2 18 1 29 2797 Norwich Norwich Municipal 1988 6 18 2 40 2801 Gowanda Gowanda Municipal 1988 2 24 2 30 2810 Niskayuna Niskayuna Municipal 1989 3 18 1 40 91-697 Macedon Macedon Municipal 1992 2 18 1 35 92-798 Syracuse Onondaga County Municipal 1993 36 30 3 100 93-039 Brookylyn Brookylyn Municipal 1993 4 30 1 60 93-064 New York Bowery Bay Municipal 1995 32 30 4 81 94-176 Amherst Amherst Municipal 1995 8 30 2 50 94-230 Lima Lima Municipal 1995 1 24 1 25 97-550 Geneva Geneva Municipal 1997 2 30 2 50 00-004 Webster Walter W. Bradley Municipal 2001 4 3 @ 18 - 1 @ 30 2 1 @ 30 - 1 @ 35 06-549 Middletown City of Middletown Municipal 2008 6 30 2 41 State: North Carolina 5 102 2759 Durham Durham Municipal 1986 16 24 4 60 89-415 Fayetteville Cross Creek Municipal 1990 18 24 2 90 91-699 Durham Durham Municipal 1992 14 30 2 70 94-194 Wilmington James A. Loughlin Municipal 1995 18 18 2 60 96-472 Durham North & South WRF'S Municipal 1997 36 30 4 85 State: Ohio 11 57 2568 Franklin Franklin Municipal 1977 3 12 1 35 2614 Crestline Crestline Municipal 1977 2 12 1 35 2638 Lancaster Lancaster Municipal 1980 3 12 1 55 2673 Covington Covington Municipal 1982 3 12 1 40 2745 Massilon Massilon Municipal 1986 10 18 2 75 2750 Piqua Piqua Municipal 1985 3 18 1 50 2765 Chillicothe Chillicothe Municipal 1986 6 24 2 55 2775 Girard Girard Municipal 1987 4 18 1 45 94-198 Lancaster Lancaster Municipal 1995 5 30 1 55 01-221 Massilon City of Massilon Municipal 2002 14 30 2 75 State: Oklahoma 2 19 2781 Lawton Lawton Municipal 1987 10 24 2 75 92-803 Muskogee Muskogee Municipal 1993 9 24 3 2 @ 50 - 1 @ 55 State: Oregon 2 19 2757 Roseburg Roseburg Municipal 1986 7 12 2 1 @ 35 - 1 @ 60 2773 Clackamas City Tri Cities Municipal 1986 12 12 2 65 State: Pennsylvania 6 42 2700 Philadelphia Southwest WWTP Municipal 1985 24 12 4 110 2787 Catasaqua Catasaqua Municipal 1987 2 18 1 40 95-281 Slatington Slatington Municipal 1996 1 30 1 35 05-469 Abington Township of Abington Municipal 2006 3 24 1 47 09-658 Hermitage Hermitage, 2PAD Municipal 2013 10 24 4 1 @ 32 - 3 @ 40 13-807 Catasaqua Catasaqua Municipal 2013 2 18 1 40 State: South Carolina 2 41 98-687 Greenville Lower Reedy Municipal 1999 14 30 1 95 98-707 Greenville West.Carolina Regional Municipal 1999 27 30 3 85 State: South Dakota 4 14 2666 Milbank Milbank Municipal 1982 6 18 2 45 2740 hot Springs hot Springs Municipal 1985 1 18 1 25 93-001 Watertown Watertown Municipal 1993 4 24 & 30 2 58 96-394 Watertown Watertown Municipal 1997 3 30 1 58 State: Tennessee 2 37 84-953 Jasper Tennol Energy Co Chemicals 1985 1 20 1 52 01-152 Mocassin Bend Chattanooga - 2PAD Municipal 2006 36 30 6 65 State: Texas 4 60 2548 Odessa Odessa Municipal 1976 4 12 1 80 89-431 El Paso El Paso Municipal 1990 28 24 2 104 90-581 Wichita Falls Wichita Falls Municipal 1991 16 24 2 80 98-631 Austin Hornsby Bend Municipal 1999 12 30 1 105 State: Utah 2 25 2714 Salt Lake City Salt Lake City Municipal 1982 20 18 4 3 @ 95 - 1 @ 100 2770 Price Price Municipal1986 5 18 1 65 State: Vermont 12 40 2655 Middlebury Middlebury Municipal 1981 3 12 1 50 2710 Newport Newport Municipal 1983 1 18 1 30 Cannon Mixer Installations
Contract number Contract city Name of Facility Industry Start date Number of units Size Number of digesters Diameter of digesters (in) (ft) Plants Mixers 2728 Essex Junction Essex Junction Municipal 1984 3 12 1 50 2752 Rutland Rutland Municipal 1986 6 12 3 2 @ 38 - 1 @ 40 2783 Burlington Burlington Municipal 1987 6 12 3 30 92-828 Rutland Rutland Municipal 1993 3 18 1 40 99-753 Montpelier Montpelier Municipal 2000 2 30 2 30 99-963 Montpelier Montpelier Municipal 2000 3 24 1 43 06-519 Newport City of Newport Municipal 2007 1 30 1 35 09-690 South Burlington Airport Parkway - 2PAD Municipal 2012 4 1 @ 18 - 3 @ 24 4 1 @ 18 - 3 @ 30 11-747 Brattleboro Brattleboro - 2PAD Municipal 2013 5 4 @ 24 - 1 @ 30 3 1 @ 26 - 1 @ 40 - 1 @ 35 12-797 Essex Junction Essex Junction Municipal 2013 3 24 1 50 State: Virginia 8 137 2803 Newport News James River Municipal 1989 14 24 2 85 92-733 Fredericksburg Fredericksburg Municipal 1993 9 30 3 40 95-371 Quantico Marine Corp Base Municipal 1997 6 30 2 46 96-490 Centreville UOSA WRF Municipal 1998 21 30 3 80 02-252 Alexandria City of Alexandria Municipal 2003 48 30 4 95 Peppers Ferry Peppers Ferry Municipal 2003 5 30 2 70 07-594 Chesterfield Proctors Creek Municipal 2010 24 30 3 85 07-594 Chesterfield Proctors Creek Municipal 2010 10 30 1 100 State: Washington 1 4 2737 Pullman Pullman Municipal 1985 4 12 2 40 State: Wisconsin 14 131 2665 Cumberland Cumberland Municipal 1982 1 18 1 30 2680 Durand Durand Municipal 1982 1 18 1 28 2729 New Richmond Municipal 1983 1 24 1 40 2742 Milwaukee Milwaukee Municipal 1986 36 18 4 125 89-411 Marinette Marinette Municipal 1991 8 24 2 50 89-412 Port Washington Port Washington Municipal 1990 6 24 2 1 @ 32 - 1 @ 45 90-551 Burlington Burlington Municipal 1992 3 24 1 50 91-684 Algoma Algoma Municipal 1992 3 24 1 40 91-724 Wisconsin Rapids Wisconsin Rapids Municipal 1992 8 24 2 55 92-846 Waukesha Waukesha Municipal 1993 16 30 4 2 @ 90 92-846 Waukesha Waukesha Municipal 1993 8 24 4 2 @ 55 94-148 Stoughton Stoughton Municipal 1994 6 18 2 40 03-358 Madison Madison MSD Municipal 2004 28 30 4 80 08-648 Wisconsin Rapids Wisconsin Rapids Municipal 2009 6 24 1 70 State: Wyoming 1 8 2809 Cheyenne Cheyenne Municipal 1989 8 24 1 70
Country: Canada 3504 Ontario Caledon-Bolton Municipal 1978 2 18 1 40 3511 Ontario Waterloo Municipal 1978 6 12 1 110 3522 Ontario Fort Erie Municipal 1978 2 18 1 30 3531 Ontario Kenora Municipal 1978 2 18 2 1 @ 30 - 1 @ 20 3533 Ontario Duffins Creek Municipal 1977 12 18 2 110 3549 Alberta Red Deep Municipal 1977 6 12 2 60 3557 Ontario Cobourge Municipal 1972 2 18 1 60 3596 Ontario Iroquois Municipal 1984 1 18 1 25 3603 Ontario Kingston Municipal 1977 3 12 1 50 3681 Manitoba City of Winnipeg Municipal 1982 40 18 4 110 99-765 Ontario Inland Empire Utilities Municipal 2000 7 30 1 65 Country: Egypt 93-123 Cairo City of Cairo Municipal 1994 1 24 1 25 Country: Jordan 2669 Amman City of Amman Municipal 1982 12 12 4 56 12-796 Amman As Samra Municipal 2013 42 30 3 111.52 Country: Sweden 96-461 Anniviers Anniviers Municipal 1997 3 30 1 32 Country: Chile 01-203 Santiago La Farfana Municipal 2002 112 30 8 111.5 10-704 Santiago Mapocho Municipal 2010 16 30 1 112 Country: Panama 10-719 Panama City Panama Municipal 2010 10 30 2 61 Country: UAE Cannon Mixer Installations
Contract number Contract city Name of Facility Industry Start date Number of units Size Number of digesters Diameter of digesters (in) (ft) Plants Mixers Palm Jumeirah, Dubai Municipal 2009 Country: Brazil 97-609 Rio De Janeiro Sao Goncalo Municipal 1998 14 30 2 65.6 BROCHURE
Cannon® Mixer enhanced sludge mixing technology
biosolids treatment
COST EFFECTIVE • reduction in O&M and solids handling costs complete sludge mixing
ENVIRONMENTALLY FRIENDLY through biogas re-circulation • low energy consumption and increased biogas production
ready for the resource revolution gas bubble mixing for any size mixer stack anaerobic digestion optional systems heating jacket Cannon Mixer® technology piston bubble Peak digester performance and reduced operating costs are digester liquid being mixed directly associated with the use of Cannon® Mixer. Ideal for deep- tank mixing, the system is not only superior in performance but is also easy to install, operate, and maintain. continous • Based on computerized modeling, multiple units are strategically compressed digester gas arranged to achieve more than 90 percent total active volume. supply • Recirculated gas is continuously fed to the bubble generator and intermittently discharged into the stack pipe as a large piston bubble bubble. generator • The piston bubble fills the entire cross section of the pipe, driving out liquid as it rises and creating a siphon. As one bubble leaves the stack pipe at the top, another enters from the generator for both continuous mixing and prevention of solids settling. • Large bubbles burst as they leave the liquid surface, creating substantial turbulence that prevents scum buildup. advantages Optional Heating Jacket • COMPLETE MIXING (90% ACTIVE VOLUME): Serving as a highly efficient tube-in-tube heat exchanger using hot water recirculation, the Cannon® Mixer’s heating jacket provides the least + less scum formation, less solids settling, superior expensive method to heat sludge, eliminating external heat exchangers volatile destruction, and more biogas production and sludge recirculation equipment. The jacket’s heating actions is uniform and optimally controlled, maintaining temperatures within • COST-EFFECTIVE: 1ºF in even the largest digesters for homo-geneous solids reduction. + low overall energy requirement reduces operating costs by up to 50% + optional heating jackets eliminate the need for KEY FIGURE: expensive external heat exchangers and pump systems Capable of achieving minimum • RELIABILITY: active sludge + no submerged moving parts means easy installation % and low maintenance 90 mixing volume
SUEZ contact 8007 Discovery Drive Richmond, VA 23229 USA Tel. : +1 804 756 7600 Fax : +1 804 756 7643 042016 [email protected] North American Locations:
Appendix E Biogas Purification Technologies
Appendix E – Biogas Purification Technologies Page 1
OVERVIEW OF BIOGAS PURIFICATION TECHNOLOGY ALTERNATIVES
Biogas purification technology alternatives target the separation of CO 2, hydrogen sulfide, VOC and siloxane removal, and/or oxygen and nitrogen removal. The technology alternatives available to remove each type of contaminant is discussed in this section. 1.1.1 CARBON DIOXIDE SEPARATION TECHNOLOGIES
Carbon dioxide must be removed to meet gas line CO 2 tariffs and increase the gas heating value. The most common technologies for CO 2 and CH 4 separation are:
• Water scrubbing • Pressure swing adsorption • Membrane filtration
Amine scrubbing also is utilized for CO 2 removal, but is only financially beneficial for very large systems. Each of these technologies are described below. Water Scrubbing Water scrubbing utilizes a counter-current water shower that scrubs the undesirable gases, such as CO 2 and H2S, out of the biogas stream. The biogas is compressed and sent into the bottom of the scrubber, flowing upward through packing material. Water is flowing downwards through the packing material, collecting the CO 2 and H2S from the biogas. The scrubbed biogas exits the top of the scrubber and goes through a gas dryer. The “dirty” water exits the bottom of the scrubber and is recycled using a stripping column. Many water scrubbing systems utilize a flash tank to increase the system methane recovery. In this scenario, the water from the scrubber vessel flows to the flash vessel, where the pressure is reduced to flash the methane off from the liquid waste stream to be returned to the system feed. Flash vessels operate at a pressure less than the scrubber vessel, but higher than the stripping vessel to keep from allowing backflow. The air stream exiting the stripping column, that contains the impurities removed from the biogas, often must be treated before being released to the atmosphere, as it will contain the H2S that was removed from the biogas. Water scrubbing systems can remove a fraction of the VOCs and siloxanes from the waste stream. A key characteristic of a water scrubber system is that it does not remove nitrogen or oxygen from the biogas. The water scrubbing process can typically only remove up to 2,500 mg/l of H2S. A typical water scrubbing process flow is shown in Figure 1.
Appendix E – Biogas Purification Technologies Page 2
Figure 1: Typical process schematic for a water scrubbing system.
Pressure Swing Adsorption (PSA) Pressure swing adsorption utilizes a selective media to separate CO 2 and other compounds, depending on the media, from raw biogas. First, the biogas is compressed and fed into a pressure vessel containing the media. The CO 2 is adsorbed by the media while the CH 4 passes through. The media has a limited capacity to adsorb CO 2, so at a fixed interval the pressure vessel will go into a “purge” cycle, shutting off the flow of raw biogas, and using a vacuum pump system to decrease the pressure, allowing the CO 2 to desorb and flow into a waste stream. PSA systems continuously run through this “pressure swing” cycle and are designed so that biogas is always being treated, so while one pressure vessel is being purged, another is still treating raw biogas. Depending on the manufacturer and the type of media used, PSA can remove limited amounts of nitrogen and oxygen, but most do not remove H2S. There is one manufacturer that can remove H2S, while the rest require pretreatment of the gas, as the H2S will foul the media used in their PSA systems. Figure 2 shows a typical PSA system including upstream H 2S removal. PSA systems designed to separate CO 2 will typically also capture a significant fraction of VOCs and siloxanes. Appendix E – Biogas Purification Technologies Page 3
Figure 2: Typical Pressure Swing Adsorption Process Diagram.
Membrane Filtration A membrane filtration biogas upgrading system consists of a pretreatment step to remove water, H2S, VOCs and siloxanes followed by compression of the biogas before injection into the membrane filter. Figure 3 shows a typical membrane upgrading process schematic. Membrane operation is based on diffusion through the membrane barrier of CO 2 and other impurities occurring much faster than the diffusion of CH 4, resulting in increased CH 4 concentrations in the upgraded gas. See Figure 4 for a visual representation of a single membrane. Operating membranes as a multi-step treatment process allows for further purification of the gas, resulting in almost pure gas streams. The pretreatment step is necessary to prevent condensation during the compression stage and to lower the concentration of H 2S, VOCs and siloxanes sent to the membrane filter. The membrane will separate H 2S but is not capable of removing enough H 2S to meet pipeline quality requirements.
Figure 3: Typical process schematic of a membrane filtration system including pretreatment
Membrane filtration will remove the CO 2, some H2S and oxygen, as shown in Figure 4 . Manufacturers use recycle streams and multiple stages of membranes, in different configurations, to achieve the Appendix E – Biogas Purification Technologies Page 4
quality of gas and CH 4 recovery required, balanced with the amount of energy required to perform the purification. Many manufacturers sell membrane upgrading processes as modular, containerized products.
Figure 4: Representation of how membrane filters work
Two stage membrane systems have a lower capital cost than three stage membrane systems but have a lower methane recovery rate. The methane recovery rate is the amount of methane within the raw gas that is captured in the finished gas. It is possible to install a two-stage system with the option to increase the recovery later by adding a third stage to minimize initial capital cost. 1.1.2 COMPRESSION
Different biogas upgrading technologies require different feed pressures. As a result, raw biogas must be compressed or boosted prior to removal of CO 2. The decision whether to utilize compressors or blowers is determined by the removal technology and the raw feed gas pressure. Prior to CO 2 separation, pretreatment of the gas will occur. This process will have separate compression requirements and equipment depending on the pretreatment equipment needs, described in other sections of this report. Fans, blowers and compressors are similar, but are utilized in different situations and have different impacts on the gas being processed. The technical delineation is denoted by the American Society of Mechanical Engineers (ASME) based on the specific ratio. Specific ratio is defined as the discharge pressure over the suction pressure. In terms of facility operation all three types of equipment are used to transport gasses through closed pipeline and the type of equipment is selected based on the specific situation. Table 1 compares fans, blowers and compressors for pressure boosting of gas. Table 1: Comparison of fans, blowers and compressors for gas transportation Specific Ratio Pressure Increase Fan < 1.11 <1 atmosphere Blower 1.11 – 1.2 ~1 atmosphere Compressor >1.2 >1 atmosphere
Appendix E – Biogas Purification Technologies Page 5
Fans can transport large volumes of air without increasing the pressure. Blowers transport larger volumes of gas than compressors, but increase the pressure less. The blower increases the kinetic energy in the gas being transported. Compressors are classified as positive displacement or dynamic based on the movement of the gas through the compressor. Positive displacement compressors do not allow the gas to flow backwards through the compressor while the pressure increase and gas intake are intermittent. Dynamic compressors utilize continuous gas flow through successive stages of compression. A dynamic compressor is either considered axial or centrifugal based on the flow direction. Not only do compressors move gas from one point to another, compressors also increase the gas pressure, density and temperature.
Compression of the biogas is required at the beginning of the upgrading process, prior to the CO 2 separation process. Compression of the finished gas is also required before injection into the utility pipeline. All compression stations need to be able to perform at a variety of flow rates. To reach various turndown levels, methods such as multiple compressors, VFDs, slide gates, and recirculation can be implemented. Design for each compression station needs to account for the pressure losses throughout the pipeline and at appurtenances. The finished gas compression needs to be able to operate at different pressures due to seasonal changes in the receiving pipeline. Types of compressors that may be used include oil flooded screw compressor, centrifugal compressors, and reciprocating compressors. Each compressor station is analyzed, and the most appropriate compressor type is selected for the specific application. 1.1.3 HYDROGEN SULFIDE REMOVAL TECHNOLOGY
Hydrogen sulfide (H 2S) is always present in biogas. The primary mechanism for production of this compound is the reduction of sulfur-containing proteins under anaerobic conditions by sulfate- reduction microorganisms. Inorganic sulfur, including sulfates, can also be biochemically converted producing considerable H 2S. Hydrogen sulfide is corrosive to most equipment (compressors, engines, pipelines, valves, storage tanks, etc.) and H 2S combustion leads to sulfur dioxide emissions, which have harmful environmental effects. It is recommended that H 2S be removed early in the biogas upgrading process.
H2S can be removed from digester gas with use of one or more of the following proven approaches:
• Micro-aeration or oxygenation in the digester or digester headspace • Iron salt addition in digester • Biological treatment (bioscrubbers, biofilters or biotrickling filters) • Caustic scrubbing • Iron Sponge Media Absorbers • Iron Granular Media
• Other Bulk H 2S Removal Alternatives • Carbon adsorption
Appendix E – Biogas Purification Technologies Page 6
Due to the Roanoke WPC facility feeding ferric in the treatment process and utilizing iron sponge adsorption vessels the H 2S is greatly reduced in the gas. Other options for removing H2S are included in this section to provide alternatives to the current operation.
Micro-Aeration or Oxygenation in Digester (or Digester Headspace) Dosing small amounts of air or oxygen into an anaerobic digester is a highly efficient and economically feasible technique for H 2S removal from biogas. Micro-Aeration oxidizes sulfide into elemental sulfur by the action of sulfide oxidizing bacteria. The use of micro-aeration is advantageous to desulfurization alternatives downstream of the digester because it limits corrosion within the digester and limits toxicity to methanogens in the digester.
Iron Salt Addition The use of iron salts for mitigation of H2S is widespread in wastewater treatment facilities. Iron salts are typically added to the wastewater collection system for odor control, H 2S control, orthophosphate reduction, and struvite mitigation. Ferric chloride is often used for addition to the headworks and primary clarifiers because of the formation of ferric hydroxide (Fe(OH) 3), which acts as a coagulant. Ferric hydroxide precipitates with sludge and reaches the digesters where it reacts to remove H2S and phosphate. Typical dosages of ferric chloride addition to the raw influent are 5 to 20 milligrams per liter (mg/L). Both ferric chloride and ferrous chloride can be added directly to the anaerobic digesters. Typical dosage rates for ferric chloride and ferrous chloride are 3.2 and 3.7 grams per gram of H2S. Benefits from iron salt addition to anaerobic digesters include H2S control, precipitation of phosphorus and increased degradability of grease. Disadvantages of iron salt addition include increased sludge volumes and the potential for formation of Vivianite scale in piping and heat exchangers. Ferric has been added successfully at the Roanoke WPC facility in multiple applications. Currently ferric is added to the influent prior to the influent pump station. To the secondary clarifiers and prior to the settling basins. Data from 2017-2018 indicate ferric was fed at an average rate of 1,726 gallons per day. It has been assumed ferric will continue to be fed in similar dosages in the future to current day operation.
Biological H2S Removal There are numerous biological H2S removal technologies that rely on the natural biological metabolism of sulfur-oxidizing bacteria to convert H2S into elemental sulfur or sulfate. Biological H2S removal systems can be set up in three basic configurations: A) Bioscrubbers: In a bioscrubber, pollutants are absorbed into liquid flowing counter currently through an absorption column, similar to a water scrubber. The liquid is then sent to a bioreactor for microbes to degrade the contaminants. B) Biofilter: A biofilter consists of a bed of organic material that stimulates biofilm growth through which humidified biogas is conveyed. Contaminants in the biogas absorb and adsorb Appendix E – Biogas Purification Technologies Page 7
into the biofilm and interact with the microbes. Although biofilters are the most commonly used (for odor control applications), H 2S-induced acidification due to the static medium can occur, which hinders microbial activity and can render biofilters ineffective for long term H 2S removal for gas streams with high H 2S inlet concentrations. C) Biotrickling Filters: Biotrickling filters overcome this problem by combining biofilters with bioscrubbers. Biotrickling filters contain a packed bed of chemically inert materials that provide large surface area for gas contact biofilm accumulation. Biogas is injected up through the column while liquid counter-currently flows down, providing contaminant absorption, delivering nutrients to the microbes, and controlling the pH. Biogas is mixed with 4-6% air before entry into the filter bed to supply sulfur-oxidizing microorganisms with O 2 needed for the conversion of H 2S to S 2 and H 2SO 4. A schematic representing the three types of biological H2S removal systems has been included below as Figure 5.
Figure 5: Biological H 2S Removal Options – A) Bioscrubber; B) Biofilter; C) Biotrickling Filter
Bioscrubbers are the only option, of the three, for most pipeline quality biogas upgrading systems because both biotrickling filters and biofilters introduce air into the biogas stream. Bioscrubber (and other biological H 2S removal systems) performance is subject to variations in environmental conditions such as temperature, pH, moisture, nutrient concentrations, and microbial community. The majority of microbes grow and function optimally near 35-degrees C and neutral pH. Wide deviations from these levels will negatively impact the efficiency of the treatment system. Bioscrubber systems are higher in capital cost than granular iron media scrubbers and caustic scrubber systems, but much lower than capital costs than water scrubbers, amine scrubbers, or iron chelating technologies. Bioscrubber technology offers very low O&M costs because only a small amount of caustic is typically required to neutralize the bioreactor. There are no media replacement costs associated with bioscrubbers, and minimal energy consumption.
Appendix E – Biogas Purification Technologies Page 8
However, bioscrubbers offer several drawbacks beyond the higher capital costs:
• Bioscrubbers require a 1 to 3 month startup time before consistent performance is achieved • Unforeseen performance drops result from loading shocks or deficiencies • Media replacement or washing is periodically required • Media clogging is an operational risk • There are many points of potential failure in relative to the other alternatives discussed herein. The components that make up a bioscrubber system are outlined below:
Bioscrubber Nutrient Storage and Ancillaries Bioreactor Nutrient Feed Pump (and Spare) Air Blower (and spare) Bioreactor Settling Vessel Solution Circulation Pump (and spare) Slurry Transfer Pump (and Spare) Caustic Receiving, Storage & Ancillaries Solution Circulation Pump (and spare) Caustic Feed Pump (and spare) Instrumentation, Valves, Controls
Caustic Scrubbers
Caustic scrubbers are recirculating liquid packed towers, which scrub H 2S from the gas stream to a high pH and low temperature water solution by utilizing contact to dissolve the H 2S from the gas into the liquid stream. This type of system utilizes a caustic soda solution recirculated through the packed tower. A continuous feed of caustic soda solution is required as the reactions removing H2S consume the caustic and create acidic conditions. Caustic can either be purchased and handled as a dry chemical or as a liquid chemical. The chemical is hazardous and requires special care for use. In addition, a continuous amount of softened water is also required. A continuous amount of waste is also required to eliminate the resulting spent liquid. The disposal of this liquid waste is relatively simple, accomplished within a wastewater treatment plant. Caustic scrubbing is comparable in capital cost to iron media adsorption technologies, and it has historically resulted in lower replaceable costs than iron sponge or granular iron media replacement costs. So historically, caustic scrubbing has been considered to be more economical than iron media absorbers in applications with high H 2S concentrations, over 1,000 – 1,500 ppm. However, there have been advances in iron media development, so the operational cost savings are comparable at this time. Perhaps the biggest weakness in the caustic scrubber system is the number of potential single points of failure. The components that make up a caustic scrubber system are outlined below: Caustic Scrubber (2) Heat Exchanger (2) Liquid Recirculation Pump (2) Water Softener Chiller (2) Buffer Vessel (2) Glycol Liquid Pump (2) Caustic Receiving, Storage & Ancillaries Instrumentation, Valves, Controls Caustic Feed Pump (and Spare
Appendix E – Biogas Purification Technologies Page 9
Iron Sponge Media Adsorption The oldest commercial process, and the one currently used at the Roanoke WPC facility, for removing H2S from digester gas is “iron sponge” which has been in use for more than 100 years. Iron sponge normally has the lowest initial cost of all commercial processes. The iron sponge concept is quite simple: hydrated iron oxide is impregnated onto redwood chips. The wood chips are placed in a vessel where the gas flows over the wood chips. The H 2S in the gas reacts with the iron oxide to form iron sulfide with removal efficiencies reported up to 99.9% The typical use of anaerobic iron sponge has two potential drawbacks. The first major drawback is that during media change-out, a highly exothermic oxidation reaction could take place causing the media to catch on fire spontaneously. This risk can be reduced with a change to an aerobic sponge with continuous air regeneration. This allows the reaction to go to completion between the iron, sulfur, and cellulose. This method adds about one (1) percent oxygen as outside or compressed air into the feed of the sponge, which results in a higher loading capacity, and elimination of the potential fire issue. The spent wood chip media is nonhazardous and can be used as a soil fertilizer. Servicing typically requires that spent wood chip media be removed with 6,000 pounds per square inch (psi) water jet, shovel and/or be vacuumed out. Iron sponge has historically only been economically utilized for H 2S levels up to approximately 1,500 ppm. At H 2S concentrations above this threshold, the use of biological or caustic scrubber technologies begin to have a life cycle cost advantage. See Figure 6 for a schematic of an iron sponge absorber system. Because Iron Sponge Adsorption is currently used at the facility and the operators are familiar with the process, this technology is anticipated for continued use.
Figure 6: Schematic of an iron sponge removal system for H2S removal
Iron Granular Media Adsorption A number of iron based, engineered medias have been developed as alternatives to iron sponge medias for H 2S. SulfaTreat is one of several such medias that has a long history in H 2S removal. Others, like FerraSorp and SulfaTrap are newer on the market and offer significant advantages relative to iron sponge and SulfaTreat medias. Appendix E – Biogas Purification Technologies Page 10
SulfaTreat An alternative form of solid media H 2S removal system termed “SulfaTreat” uses iron oxide-based chemistry, but uses a different substrate media base to address the media changeout issue. SulfaTreat is a dry, free-flowing iron oxide-based media that selectively removes H2S and some light mercaptans from gas and liquid streams. SulfaTreat begins as a safe and stable compound that is environmentally non-hazardous in un-reacted and reacted forms. The solid, clay-like media is an inorganic ceramic material coated with an iron oxide. The iron oxide reacts with the H 2S to form iron pyrite, also a stable compound. The unique media molecular structure allows SulfaTreat to remove more sulfur than iron sponge per unit volume. However, the cost of the media as compared to iron sponge is significantly higher. Additionally, required contact time is higher than that of iron sponge. This results in a larger storage requirement with a higher system capital cost, in addition to a higher media replacement cost while resulting in longer times between media change-out. The spent media is nonhazardous if handled correctly. Servicing typically requires that the spent media be vacuumed out. Occasionally, media composition requires removal with a high-pressure water auger. Similar to iron sponge, SulfaTreat is commonly utilized for H 2S levels up to approximately 1,500 ppm. Several facilities that utilize SulfaTreat have observed challenges in media removal similar to the challenges associated with the removal of iron sponge media, and the media does not appear to be priced to outperform iron sponge economically. FerraSorp Another media in the market today is FerraSorp, a German media. FerraSorp is a pelletized, organic compound based on iron hydroxide and combined with alkalizing compounds (Ca(OH) 2 and others), binding substances, and highly porous. The manufacturer claims the iron hydroxide is primarily located on the inside of the pellet and iron hydroxide is not exposed in high quantities on the surface of the pellet. The claim is that it will not result in clumping or bridging like iron sponge media for this reason, making removal of the media considerably easier than with iron sponge media. A case study from a side-by-side test of medias at Chicago’s Stickney plant showed that FerraSorp media can have a bed life of more than 70% longer than conventional iron sponge media. Based on the pricing of the product and the claimed capacity, this media is very cost competitive as an iron absorption media.
FerraSorp requires continuous regeneration with oxygen to achieve claimed levels of capacity for H 2S. “Offline regeneration” (by removing the media and exposing it to air before reinserting it into the vessel) is not an option with FerraSorp. The system requires the addition of 0.2% - 1.0% oxygen (depending on the H 2S concentration) to continuously regenerate the media. While this operational strategy has been implemented successfully in many pipeline quality gas applications with low finished gas oxygen concentration requirements, it is a challenge that can be avoided with traditional iron sponge technologies that allow for offline regenerations. The continuous regeneration process is required for FerraSorp because following the regeneration reaction, elemental sulphur gets bound onto the surface layers as well as in the pores of the pellets. Thus, the surface of the pellets gets blocked over time, and the gas is no longer exposed to the iron Appendix E – Biogas Purification Technologies Page 11
hydroxide on the interior of the pellet. If done “offline”, the media’s surface can quickly become bound with elemental sulfur.
Another challenge observed when using this media is that during startup, when CO 2 is first introduced into the reactor, the CO 2 reacts with calcium hydroxide within the pellets to form calcium carbonate, and the filter media temperature can quickly rise to 70-80 degrees C. This reaction is finished within a short time after startup, and the media reaches its final durability and is ready to be used.
In the removal of the H 2S with FerraSorp, the gas phase H 2S is first solubilized before removed by microbes. Therefore, a minimum gas moisture of 40% is required (and a maximum of 60% is recommended). Water is also required in the continuous regeneration process, as 3 mols of water are required to regenerate 1 mol of iron sulfide into iron hydroxide. When optimizing the gas moisture, it is recommended that the gas is not condensing in the absorber, or at least the condensate should not stay in contact with the compound for extended periods. Water accumulation should be removed routinely in operation, as pellets can be dissolved and/or turn into iron hydroxide sludge. When filling the vessel with the pellets, rough handling may cause the pellets to fall apart, decreasing the effectiveness of the product. The equipment needed to carefully place the media into the vessel is provided by the manufacturer. Dust is present in the handling of the product, so a dust-mask, gloves, and safety glasses should be worn when filling vessels.
Other Bulk H 2S Removal Alternatives There are several other technologies that are utilized to remove bulk quantities of H 2S, including water scrubbing, amine scrubbing, chelated iron liquids, and the Claus process. Each of these technologies are viable alternatives for significantly higher quantities of H 2S removal than what is commonly encountered in wastewater treatment anaerobic digester biogas applications. However, each is significantly higher in capital cost than can reasonably be considered in this application due to the low gas flows and low H 2S loadings. Carbon Adsorption Carbon adsorption is highly effective in H 2S polishing applications. However relative to the above described H 2S removal technologies, carbon replacement costs make carbon cost prohibitive as a method of bulk desulfurization.
A carbon adsorption system is typically utilized for final H2S polishing downstream of any bulk desulfurization method. The bulk desulfurization alternatives are capable of getting H2S levels down to the 50-100-ppm range or lower. The carbon adsorption system would remove additional H2S down to the 2-ppm range. The carbon chosen is based on the anticipated H2S levels in the carbon system feed gas. 1.1.4 VOC/SILOXANE REMOVAL
Siloxanes are organic compounds containing silicon, hydrogen, oxygen and carbon. Siloxanes in digester gas are the result of silicon monomers in industrial silicon products, including personal hygiene and healthcare products. Pipeline gas tariffs require removal of siloxanes. If siloxanes are present when gas is combusted, silicon dioxide is left behind, reducing performance and increasing maintenance needs. Appendix E – Biogas Purification Technologies Page 12
Silicon dioxide has been linked to health effects including lung disease. Removal in gas upgrading processes most often is done via adsorption, single or two-stage refrigeration. If siloxanes are removed via adsorption, other constituents can be removed in the same process, specifically H2S and VOCs. Adsorption can be accomplished through non-regenerative adsorption vessels, such as activated carbon, or regenerative adsorption vessels, such as temperature/pressure swing adsorption. In small siloxane loading conditions, non-regenerative adsorption vessels are advantageous relative to regenerative systems because of the lower capital costs. VOCs are either manmade or naturally occurring hydrocarbons with low water solubility. When released to the atmosphere, VOCs result in odors and pose environmental or health concerns. VOCs can be removed from gas using thermal oxidation, biological treatment, adsorption, scrubbing or membrane treatment. However, VOCs are typically removed in biogas upgrading systems upstream of the primary biogas upgrading equipment, using adsorption technology to protect the integrity of the upgrading equipment. Pressure Swing Adsorption or Temperature Swing Adsorption Pressure swing adsorption (PSA) utilizes a selective media to separate VOCs and siloxanes, depending on the media, from raw biogas. Figure 2 in Section 1.1.1 shows the treatment process schematic associated with a typical PSA system, the general process is the same for CO 2 separation and VOC/siloxane removal. First, the biogas is compressed and fed into a pressure vessel containing the media. The contaminants are adsorbed by the media while the CH 4 passes through. The media has a limited capacity to adsorb the contaminants, so at a fixed interval the pressure vessel will go into a “purge” cycle, shutting off the flow of raw biogas, and using a vacuum pump system to decrease the pressure, allowing the contaminants to desorb and flow into a waste stream. PSA systems continuously run through this “pressure swing” cycle and are designed so that biogas is always being treated, so while one pressure vessel is being purged, another is still treating raw biogas. A temperature swing adsorption system operates in the same manner, but the swing cycle is based on temperature rather than pressure. TSA is another viable alternative for removal of large loadings of VOC and siloxanes. Carbon Adsorption Similar to H2S polishing, carbon adsorption can also be used to remove VOC and siloxanes from raw biogas. The process is non-regenerative, the carbon eventually must be removed from the vessel with fresh carbon added. To aid in maintenance efforts a free-flowing carbon is often utilized. Specific carbons are chosen for each application based on the quantity and speciation of VOC and siloxanes in the raw gas. Figure 7 shows a two-vessel carbon adsorption system. The system shown is set up for operation in parallel, but systems also can be set up in a lead lag situation, like the iron sponge for H2S removal. The bed life is determined by the volume of media contained in each vessel and the amount of VOC or siloxanes in the raw gas. For design, six months of bed life is often specified. Appendix E – Biogas Purification Technologies Page 13
Figure 7: Example of a two-vessel carbon adsorption system
The Roanoke WPC facility currently used carbon adsorption vessels for removal of siloxanes and VOCs. Due to past successful use it is expected that use of carbon adsorption vessels will continue. 1.1.5 OXYGEN AND NITROGEN REMOVAL
Oxygen in gas increases the rate and effect of corrosion. When water, CO 2 or H2S are present, oxygen in the gas will combine, resulting in corrosion. Oxygen in the gas can allow sulfate-reducing bacteria that cause corrosion. Often, oxygen is removed from gas via adsorption, catalytic treatment or cryogenic condensation. The effective removal process is determined based on the levels of oxygen in the raw gas and the specific finished gas tariff or use. When determining the need for an oxygen removal system it is important to consider that the amount of oxygen in the raw gas will essentially double following CO 2 removal, if there is no oxygen removal in CO 2 separation process. Pressure swing adsorption is the most common method utilized when nitrogen removal is required, as is typically the case in landfill gas upgrading applications. However, nitrogen is rarely a challenge requiring treatment in digester gas applications.
1.1.6 TAILGAS HANDLING
The biogas upgrading process results in a “tailgas” stream that generally consists primarily of CO 2, but can also include trace amounts of CH 4, sulfur compounds, oxygen, nitrogen, volatile organic compounds, or other constituents depending on the upgrading process utilized. The tailgas stream may require further treatment prior to off gassing or require air permitting for discharge. Tailgas may be directly vented, treated by thermal oxidation, or beneficially reused.
Depending on the level of CH 4 and other constituent capture achieved by the upgrading system, it may be possible to discharge the tailgas stream, without further treatment, utilizing a vent. For example, if upgrading is accomplished by utilizing a high performance, three-stage membrane system, the tailgas would contain over 99.52% CO 2 and minor concentrations of other constituents. In other applications using three-stage membrane treatment, regulators have allowed the tailgas to be directly vented. There are no three-stage systems currently utilized in the United States, but several are in design or construction. If a two-stage membrane system is used, the CH 4 slip and other constituent Appendix E – Biogas Purification Technologies Page 14
slip would be greater, and significantly higher CH 4 content and other constituent content would remain in the tailgas, providing additional incentive for treatment of the tailgas prior to final disposal. Often a thermal oxidizer is utilized for treatment and disposal of the tail gas stream. Thermal oxidizers require supplemental natural gas to achieve combustion of the tail gas stream.
The tailgas from a PSA system would include H2S, and depending on the equipment supplier used, VOCs, siloxanes, CO 2 and CH 4. A PSA system would result in a tail gas stream that needs to be treated prior to discharge, and a thermal oxidizer is most often used for this application. Tailgas from a water scrubbing system would contain H2S, VOCs, siloxanes and CO 2 requiring either a thermal oxidizer or other treatment to capture constituents prior to venting to the atmosphere.
The tailgas CO 2 from a membrane system could be reused externally by selling to an external buyer, or reused internally, within the wastewater plant. CO 2 reuse applications can be designed for 100%
CH 4 recovery by sending the impurities removed from a CO 2 recovery process to the feed of a membrane system. If designed in this way, the CO 2 upgrading process doubles as a tailgas management tool removing the need for a thermal oxidizer to decompose any CH 4 in the tailgas stream. Reuse of the CO 2 also greatly reduces the CO 2 emissions to the atmosphere, decreasing a facilities impact on air quality.
Appendix F Operational Costs
ROANOKE REGIONAL WPC RNG PRODUCTION FEASABILITY - SCENARIO 1 OPINION OF PROBABLE O,M&R COST SUMMARY
3 STAGE 3 STAGE MEMBRANE MEMBRANE (YEARLY) (PRESENT WORTH)
1 2020 $519,113 $519,113 2 2021 $521,577 $499,117 3 2022 $524,078 $479,914 4 2023 $526,616 $461,472 5 2024 $529,192 $443,760 6 2025 $531,807 $426,749 7 2026 $534,461 $410,411 8 2027 $537,155 $394,717 9 2028 $539,890 $379,642 10 2029 $542,665 $365,162 11 2030 $545,482 $351,251 12 2031 $548,341 $337,887 13 2032 $551,244 $325,048 14 2033 $554,189 $312,713 15 2034 $557,179 $300,862 16 2035 $560,214 $289,474 17 2036 $563,294 $278,532 18 2037 $566,421 $268,017 19 2038 $569,594 $257,912 20 2039 $572,815 $248,202
PW O,M&R $7,349,956 ROANOKE REGIONAL WRF RNG PRODUCTION FEASABILITY SCENARIO 1 - PRELIMINARY OPINION OF PROBABLE OPERATION & MAINTENANCE COSTS THREE STAGE MEMBRANE WITHOUT CO REUSE 2
YEAR Full Capacity FULL CAPACITY PLANT FLOW RATE (SCFM) = 250 PLANT HOURS OF OPERATION (HRS) = 24
Description Factors Sub-Totals
FERRIC ADDITION H2S IRON SPONGE REPLACEMENT LABOR NUMBER OF REPLACEMENTS PER YEAR 0.85 $8,531 CONTRACTOR COST PER REPLACEMENT $10,000.00 MISCELLANEOUS EQUIPMENT PADDED LINER COST PER CHANGE OUT $2,400 $2,048 MEDIA CHANGES PER YEAR 0.85 MEDIA LBS MEDIA PER CHANGE OUT 51,000 $26,106 MEDIA CHANGES PER YEAR 0.85 ADSORBANT COST PER LB $0.60 TOTAL H2S REMOVAL O&M COSTS $36,684 VOC/SILOXANE CARBON REPLACEMENT MEDIA LBS MEDIA PER CHANGE OUT 9,000 $55,000 MEDIA CHANGES PER YEAR 2 ADSORBANT COST PER LB $2.50 ANNUAL MEDIA CHANGE OUT LABOR COSTS $10,000.00 TOTAL CARBON ADSORPTION COSTS $55,000 FEED GAS COMPRESSION COMPRESSOR REBUILD REQUIREMENTS COMPRESSOR REBUILD FREQUENCY (YRS) 10 $2,600 COMPRESSOR REBUILD COST $26,000 MISCELLANEOUS EQUIPMENT FILTER REPLACEMENTS & ELEMENTS, OIL REPLACEMENTS $5,372.00 $5,372 SYSTEM POWER REQUIREMENTS KW USAGE 153 $93,820 HOURS OF OPERATION PER DAY 24 KW-HR PER YEAR 1,340,280 COST PER KW-HR $0.07 TOTAL FEED GAS COMPRESSION COSTS $101,792 MEMBRANES MEMBRANE REPLACEMENTS MEMBRANE REPLACEMENT COST $247,500 $24,750 MEMBRANE LIFETIME (YRS) 10 YEARLY MEMBRANE REPLACEMENT COST $24,750 TOTAL MEMBRANE REPLACEMENT COSTS $24,750 PRODUCT GAS COMPRESSION COMPRESSOR REBUILD REQUIREMENTS COMPRESSOR REBUILD FREQUENCY (YRS) 3 $13,333 COMPRESSOR REBUILD COST $40,000 MISCELLANEOUS EQUIPMENT FILTER REPLACEMENTS & ELEMENTS, OIL REPLACEMENTS $9,000.00 $9,000 SYSTEM POWER REQUIREMENTS KW USAGE 30 $18,396 HOURS OF OPERATION PER DAY 24 KW-HR PER YEAR 262,800 COST PER KW-HR $0.07 TOTAL PRODUCT GAS COMPRESSION COSTS $40,729 GLYCOL CHILLER REFRIGERATION SYSTEM REBUILD REQUIREMENTS REFRIGERATION SYSTEM REBUILD FREQUENCY (YRS) 5 $1,400 REFRIGERATION SYSTEM REBUILD COST $7,000 GLYCOL/WATER CIRCULATION REBUILD REQUIREMENTS REFRIGERATION SYSTEM REBUILD FREQUENCY (YRS) 10 $300 REFRIGERATION SYSTEM REBUILD COST $3,000 MISCELLANEOUS EQUIPMENT FILTER REPLACEMENTS & ELEMENTS, OIL REPLACEMENTS $1,500.00 $1,500 SYSTEM POWER REQUIREMENTS KW USAGE 7 $4,292 HOURS OF OPERATION PER DAY 24 KW-HR PER YEAR 61,320 COST PER KW-HR $0.07 TOTAL GLYCOL CHILLER COSTS $7,492 OPERATIONAL STAFFING OPERATIONAL STAFFING REQUIREMENTS TOTAL EMPLOYEES 1 $104,000 HOURS OF OPERATION PER EMPLOYEE PER WEEK 40 EMPLOYEE COST PER HOUR $50.00 TOTAL OPERATIONAL STAFFING COSTS $104,000
TOTAL O&M COSTS $370,448 ROANOKE REGIONAL WRF RNG PRODUCTION FEASABILITY - SCENARIO 2 OPINION OF PROBABLE O,M&R COST SUMMARY
3 STAGE 3 STAGE MEMBRANE MEMBRANE (YEARLY) (PRESENT WORTH)
1 2020 $802,182 $802,182 2 2021 $804,646 $769,996 3 2022 $807,147 $739,129 4 2023 $809,685 $709,524 5 2024 $812,261 $681,131 6 2025 $814,876 $653,898 7 2026 $817,531 $627,778 8 2027 $820,224 $602,724 9 2028 $822,959 $578,692 10 2029 $825,734 $555,640 11 2030 $828,551 $533,527 12 2031 $831,411 $512,314 13 2032 $834,313 $491,964 14 2033 $837,258 $472,441 15 2034 $840,248 $453,711 16 2035 $843,283 $435,742 17 2036 $846,363 $418,501 18 2037 $849,490 $401,959 19 2038 $852,663 $386,086 20 2039 $855,884 $370,856
PW O,M&R $11,197,797 ROANOKE REGIONAL WPC RNG PRODUCTION FEASABILITY SCENARIO 2 - PRELIMINARY OPINION OF PROBABLE OPERATION & MAINTENANCE COSTS THREE STAGE MEMBRANE WITHOUT CO REUSE 2
YEAR Full Capacity FULL CAPACITY PLANT FLOW RATE (SCFM) = 720 PLANT HOURS OF OPERATION (HRS) = 24
Description Factors Sub-Totals
FERRIC ADDITION H2S IRON SPONGE REPLACEMENT LABOR NUMBER OF REPLACEMENTS PER YEAR 2.46 $24,570 CONTRACTOR COST PER REPLACEMENT $10,000.00 MISCELLANEOUS EQUIPMENT PADDED LINER COST PER CHANGE OUT $2,400 $5,897 MEDIA CHANGES PER YEAR 2.46 MEDIA LBS MEDIA PER CHANGE OUT 51,000 $75,184 MEDIA CHANGES PER YEAR 2.46 ADSORBANT COST PER LB $0.60 TOTAL H2S REMOVAL O&M COSTS $105,651 VOC/SILOXANE CARBON REPLACEMENT MEDIA LBS MEDIA PER CHANGE OUT 9,000 $139,600 MEDIA CHANGES PER YEAR 6 ADSORBANT COST PER LB $2.50 ANNUAL MEDIA CHANGE OUT LABOR COSTS $10,000.00 TOTAL CARBON ADSORPTION COSTS $139,600 FEED GAS COMPRESSION COMPRESSOR REBUILD REQUIREMENTS COMPRESSOR REBUILD FREQUENCY (YRS) 10 $3,500 COMPRESSOR REBUILD COST $35,000 MISCELLANEOUS EQUIPMENT FILTER REPLACEMENTS & ELEMENTS, OIL REPLACEMENTS $7,000.00 $7,000 SYSTEM POWER REQUIREMENTS KW USAGE 440.64 $270,200 HOURS OF OPERATION PER DAY 24 KW-HR PER YEAR 3,860,006 COST PER KW-HR $0.07 TOTAL FEED GAS COMPRESSION COSTS $280,700 MEMBRANES MEMBRANE REPLACEMENTS MEMBRANE REPLACEMENT COST $712,800 $71,280 MEMBRANE LIFETIME (YRS) 10 YEARLY MEMBRANE REPLACEMENT COST $71,280 TOTAL MEMBRANE REPLACEMENT COSTS $71,280 PRODUCT GAS COMPRESSION COMPRESSOR REBUILD REQUIREMENTS COMPRESSOR REBUILD FREQUENCY (YRS) 3 $25,000 COMPRESSOR REBUILD COST $75,000 MISCELLANEOUS EQUIPMENT FILTER REPLACEMENTS & ELEMENTS, OIL REPLACEMENTS $9,000.00 $9,000 SYSTEM POWER REQUIREMENTS KW USAGE 90 $55,188 HOURS OF OPERATION PER DAY 24 KW-HR PER YEAR 788,400 COST PER KW-HR $0.07 TOTAL PRODUCT GAS COMPRESSION COSTS $89,188 GLYCOL CHILLER REFRIGERATION SYSTEM REBUILD REQUIREMENTS REFRIGERATION SYSTEM REBUILD FREQUENCY (YRS) 5 $2,000 REFRIGERATION SYSTEM REBUILD COST $10,000 GLYCOL/WATER CIRCULATION REBUILD REQUIREMENTS REFRIGERATION SYSTEM REBUILD FREQUENCY (YRS) 10 $400 REFRIGERATION SYSTEM REBUILD COST $4,000 MISCELLANEOUS EQUIPMENT FILTER REPLACEMENTS & ELEMENTS, OIL REPLACEMENTS $2,000.00 $2,000 SYSTEM POWER REQUIREMENTS KW USAGE 20 $12,264 HOURS OF OPERATION PER DAY 24 KW-HR PER YEAR 175,200 COST PER KW-HR $0.07 TOTAL GLYCOL CHILLER COSTS $16,664 OPERATIONAL STAFFING OPERATIONAL STAFFING REQUIREMENTS TOTAL EMPLOYEES 1 $104,000 HOURS OF OPERATION PER EMPLOYEE PER WEEK 40 EMPLOYEE COST PER HOUR $50.00 TOTAL OPERATIONAL STAFFING COSTS $104,000
TOTAL O&M COSTS $807,084
Appendix G Proforma Evaluations
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 2021 2022 2023 2024 2025 2026 2027 2028 2029 2030 2031 2032 2033 2034 2035 2036 2037 2038 2039 2040 Days/Year 365 Pure Methane to Pipeline (DTH) 172 283 395 507 619 619 619 619 619 619 619 619 619 619 619 619 619 619 619 619 Avg NG Price $3.10 $ 3.15 $ 3.19 $ 3.24 $ 3.29 $ 3.34 $ 3.39 $ 3.44 $ 3.49 $ 3.54 $ 3.60 $ 3.65 $ 3.71 $ 3.76 $ 3.82 $ 3.88 $ 3.93 $ 3.99 $ 4.05 $ 4.11 D5 RIN Price $0.50 $ 0.51 $ 0.52 $ 0.53 $ 0.54 $ 0.55 $ 0.56 $ 0.57 $ 0.59 $ 0.60 $ 0.61 $ 0.62 $ 0.63 $ 0.65 $ 0.66 $ 0.67 $ 0.69 $ 0.70 $ 0.71 $ 0.73 D3 Waiver Credit $1.80 $ 1.84 $ 1.87 $ 1.91 $ 1.95 $ 1.99 $ 2.03 $ 2.07 $ 2.11 $ 2.15 $ 2.19 $ 2.24 $ 2.28 $ 2.33 $ 2.38 $ 2.42 $ 2.47 $ 2.52 $ 2.57 $ 2.62 D3 Waiver Credit Discount 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% Avg LCFS $/Metric Ton $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00
Revenue Gross NG Revenue $ 194,399 $ 325,570 $ 460,632 $ 599,672 $ 742,781 $ 753,922 $ 765,231 $ 776,710 $ 788,360 $ 800,186 $ 812,188 $ 824,371 $ 836,737 $ 849,288 $ 862,027 $ 874,958 $ 888,082 $ 901,403 $ 914,924 $ 928,648 Gross D5 RIN Credit Revenue $ 367,696 $ 618,833 $ 879,868 $ 1,151,096 $ 1,432,822 $ 1,461,479 $ 1,490,708 $ 1,520,522 $ 1,550,933 $ 1,581,952 $ 1,613,591 $ 1,645,862 $ 1,678,780 $ 1,712,355 $ 1,746,602 $ 1,781,534 $ 1,817,165 $ 1,853,508 $ 1,890,579 $ 1,928,390 Gross D3 Waiver Credit Revenue $ 1,323,706 $ 2,227,798 $ 3,167,523 $ 4,143,947 $ 5,158,160 $ 5,261,323 $ 5,366,550 $ 5,473,881 $ 5,583,358 $ 5,695,026 $ 5,808,926 $ 5,925,105 $ 6,043,607 $ 6,164,479 $ 6,287,768 $ 6,413,524 $ 6,541,794 $ 6,672,630 $ 6,806,083 $ 6,942,204 D3 Waiver Credit Discount $ (264,741) $ (445,560) $ (633,505) $ (828,789) $ (1,031,632) $ (1,052,265) $ (1,073,310) $ (1,094,776) $ (1,116,672) $ (1,139,005) $ (1,161,785) $ (1,185,021) $ (1,208,721) $ (1,232,896) $ (1,257,554) $ (1,282,705) $ (1,308,359) $ (1,334,526) $ (1,361,217) $ (1,388,441) Gross LCFS Credit Revenue $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - Gross Total Revenue $ 1,621,060 $ 2,726,640 $ 3,874,518 $ 5,065,926 $ 6,302,131 $ 6,424,460 $ 6,549,179 $ 6,676,337 $ 6,805,980 $ 6,938,158 $ 7,072,920 $ 7,210,317 $ 7,350,402 $ 7,493,226 $ 7,638,844 $ 7,787,311 $ 7,938,682 $ 8,093,016 $ 8,250,369 $ 8,410,802
Distribution Costs $ (17,559) $ (17,822) $ (18,089) $ (18,361) $ (18,636) $ (18,916) $ (19,199) $ (19,487) $ (19,780) $ (20,076) $ (20,377) $ (20,683) $ (20,993) $ (21,308) $ (21,628) $ (21,952) $ (22,282) $ (22,616) $ (22,955) $ (23,299) RIN Management $ (211,365) $ (357,487) $ (509,370) $ (667,184) $ (831,107) $ (847,743) $ (864,712) $ (882,021) $ (899,676) $ (917,684) $ (936,053) $ (954,789) $ (973,901) $ (993,395) $ (1,013,278) $ (1,033,560) $ (1,054,248) $ (1,075,349) $ (1,096,873) $ (1,118,828) Net Total Revenue $ 1,392,136 $ 2,351,331 $ 3,347,059 $ 4,380,381 $ 5,452,388 $ 5,557,801 $ 5,665,268 $ 5,774,828 $ 5,886,524 $ 6,000,397 $ 6,116,489 $ 6,234,845 $ 6,355,508 $ 6,478,523 $ 6,603,938 $ 6,731,799 $ 6,862,153 $ 6,995,050 $ 7,130,540 $ 7,268,674 Present Worth of Revenue $ 1,392,136 $ 2,268,858 $ 3,116,380 $ 3,935,433 $ 4,726,731 $ 4,649,120 $ 4,572,795 $ 4,497,736 $ 4,423,921 $ 4,351,329 $ 4,279,940 $ 4,209,734 $ 4,140,691 $ 4,072,791 $ 4,006,016 $ 3,940,346 $ 3,875,763 $ 3,812,248 $ 3,749,785 $ 3,688,354 Cumulative PW of Revenue $ 1,392,136 $ 3,660,994 $ 6,777,375 $ 10,712,808 $ 15,439,540 $ 20,088,659 $ 24,661,454 $ 29,159,190 $ 33,583,110 $ 37,934,439 $ 42,214,380 $ 46,424,114 $ 50,564,805 $ 54,637,596 $ 58,643,612 $ 62,583,958 $ 66,459,721 $ 70,271,969 $ 74,021,753 $ 77,710,107
Total Expenses $ 802,182 $ 804,646 $ 807,147 $ 809,685 $ 812,261 $ 814,876 $ 817,531 $ 820,224 $ 822,959 $ 825,734 $ 828,551 $ 831,411 $ 834,313 $ 837,258 $ 840,248 $ 843,283 $ 846,363 $ 849,490 $ 852,663 $ 855,884 Present Worth of O&M Costs $ 802,182 $ 776,423 $ 751,519 $ 727,439 $ 704,158 $ 681,647 $ 659,880 $ 638,833 $ 618,481 $ 598,801 $ 579,769 $ 561,364 $ 543,565 $ 526,351 $ 509,703 $ 493,602 $ 478,028 $ 462,965 $ 448,396 $ 434,303 Cumulative PW of O&M Costs $ 802,182 $ 1,578,606 $ 2,330,124 $ 3,057,564 $ 3,761,721 $ 4,443,368 $ 5,103,249 $ 5,742,082 $ 6,360,563 $ 6,959,364 $ 7,539,133 $ 8,100,497 $ 8,644,062 $ 9,170,413 $ 9,680,116 $ 10,173,718 $ 10,651,746 $ 11,114,711 $ 11,563,107 $ 11,997,410
Net Profit Before Financing $ 589,954 $ 1,546,685 $ 2,539,912 $ 3,570,696 $ 4,640,126 $ 4,742,924 $ 4,847,737 $ 4,954,604 $ 5,063,565 $ 5,174,663 $ 5,287,938 $ 5,403,434 $ 5,521,195 $ 5,641,265 $ 5,763,690 $ 5,888,515 $ 6,015,790 $ 6,145,561 $ 6,277,877 $ 6,412,790 Present Worth of Net Revenue $ 589,954 $ 1,492,435 $ 2,364,862 $ 3,207,994 $ 4,022,574 $ 3,967,473 $ 3,912,914 $ 3,858,902 $ 3,805,439 $ 3,752,528 $ 3,700,171 $ 3,648,370 $ 3,597,126 $ 3,546,440 $ 3,496,313 $ 3,446,744 $ 3,397,734 $ 3,349,283 $ 3,301,389 $ 3,254,052 Cumulative PW of Net Revenue $ 589,954 $ 2,082,389 $ 4,447,251 $ 7,655,245 $ 11,677,818 $ 15,645,291 $ 19,558,205 $ 23,417,108 $ 27,222,547 $ 30,975,075 $ 34,675,247 $ 38,323,617 $ 41,920,743 $ 45,467,183 $ 48,963,496 $ 52,410,240 $ 55,807,974 $ 59,157,257 $ 62,458,646 $ 65,712,698 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 2021 2022 2023 2024 2025 2026 2027 2028 2029 2030 2031 2032 2033 2034 2035 2036 2037 2038 2039 2040 Days/Year 365 Pure Methane to Pipeline (DTH) 172 174 176 179 181 183 185 188 190 192 194 197 199 201 203 206 208 210 212 215 Avg NG Price $3.10 $ 3.15 $ 3.19 $ 3.24 $ 3.29 $ 3.34 $ 3.39 $ 3.44 $ 3.49 $ 3.54 $ 3.60 $ 3.65 $ 3.71 $ 3.76 $ 3.82 $ 3.88 $ 3.93 $ 3.99 $ 4.05 $ 4.11 D5 RIN Price $0.50 $ 0.51 $ 0.52 $ 0.53 $ 0.54 $ 0.55 $ 0.56 $ 0.57 $ 0.59 $ 0.60 $ 0.61 $ 0.62 $ 0.63 $ 0.65 $ 0.66 $ 0.67 $ 0.69 $ 0.70 $ 0.71 $ 0.73 D3 Waiver Credit $1.80 $ 1.84 $ 1.87 $ 1.91 $ 1.95 $ 1.99 $ 2.03 $ 2.07 $ 2.11 $ 2.15 $ 2.19 $ 2.24 $ 2.28 $ 2.33 $ 2.38 $ 2.42 $ 2.47 $ 2.52 $ 2.57 $ 2.62 D3 Waiver Credit Discount 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% 20.00% Avg LCFS $/Metric Ton $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00 $0.00
Revenue Gross NG Revenue $ 194,399 $ 199,911 $ 205,545 $ 211,303 $ 217,187 $ 223,201 $ 229,346 $ 235,625 $ 242,040 $ 248,596 $ 255,293 $ 262,136 $ 269,126 $ 276,267 $ 283,562 $ 291,013 $ 298,624 $ 306,398 $ 314,338 $ 322,447 Gross D5 RIN Credit Revenue $ 367,696 $ 379,985 $ 392,618 $ 405,605 $ 418,954 $ 432,675 $ 446,777 $ 461,270 $ 476,164 $ 491,469 $ 507,196 $ 523,355 $ 539,958 $ 557,016 $ 574,540 $ 592,543 $ 611,035 $ 630,030 $ 649,541 $ 669,580 Gross D3 Waiver Credit Revenue $ 1,323,706 $ 1,367,946 $ 1,413,426 $ 1,460,177 $ 1,508,234 $ 1,557,629 $ 1,608,396 $ 1,660,571 $ 1,714,189 $ 1,769,288 $ 1,825,905 $ 1,884,079 $ 1,943,850 $ 2,005,258 $ 2,068,345 $ 2,133,153 $ 2,199,726 $ 2,268,109 $ 2,338,347 $ 2,410,488 D3 Waiver Credit Discount $ (264,741) $ (273,589) $ (282,685) $ (292,035) $ (301,647) $ (311,526) $ (321,679) $ (332,114) $ (342,838) $ (353,858) $ (365,181) $ (376,816) $ (388,770) $ (401,052) $ (413,669) $ (426,631) $ (439,945) $ (453,622) $ (467,669) $ (482,098) Gross LCFS Credit Revenue $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - $ - Gross Total Revenue $ 1,621,060 $ 1,674,253 $ 1,728,904 $ 1,785,050 $ 1,842,728 $ 1,901,978 $ 1,962,839 $ 2,025,351 $ 2,089,555 $ 2,155,495 $ 2,223,213 $ 2,292,754 $ 2,364,164 $ 2,437,490 $ 2,512,778 $ 2,590,078 $ 2,669,440 $ 2,750,915 $ 2,834,557 $ 2,920,417
Distribution Costs $ (17,559) $ (17,822) $ (18,089) $ (18,361) $ (18,636) $ (18,916) $ (19,199) $ (19,487) $ (19,780) $ (20,076) $ (20,377) $ (20,683) $ (20,993) $ (21,308) $ (21,628) $ (21,952) $ (22,282) $ (22,616) $ (22,955) $ (23,299) RIN Management $ (211,365) $ (218,478) $ (225,790) $ (233,308) $ (241,036) $ (248,979) $ (257,144) $ (265,536) $ (274,160) $ (283,023) $ (292,131) $ (301,490) $ (311,107) $ (320,987) $ (331,138) $ (341,567) $ (352,280) $ (363,285) $ (374,590) $ (386,201) Net Total Revenue $ 1,392,136 $ 1,437,953 $ 1,485,024 $ 1,533,381 $ 1,583,057 $ 1,634,083 $ 1,686,495 $ 1,740,328 $ 1,795,615 $ 1,852,395 $ 1,910,704 $ 1,970,581 $ 2,032,064 $ 2,095,194 $ 2,160,012 $ 2,226,559 $ 2,294,878 $ 2,365,014 $ 2,437,012 $ 2,510,917 Present Worth of Revenue $ 1,392,136 $ 1,387,517 $ 1,382,677 $ 1,377,624 $ 1,372,368 $ 1,366,916 $ 1,361,277 $ 1,355,457 $ 1,349,465 $ 1,343,308 $ 1,336,992 $ 1,330,526 $ 1,323,915 $ 1,317,166 $ 1,310,285 $ 1,303,279 $ 1,296,153 $ 1,288,914 $ 1,281,568 $ 1,274,119 Cumulative PW of Revenue $ 1,392,136 $ 2,779,653 $ 4,162,329 $ 5,539,954 $ 6,912,322 $ 8,279,238 $ 9,640,515 $ 10,995,972 $ 12,345,437 $ 13,688,745 $ 15,025,738 $ 16,356,264 $ 17,680,178 $ 18,997,344 $ 20,307,629 $ 21,610,908 $ 22,907,061 $ 24,195,976 $ 25,477,543 $ 26,751,662
Total Expenses $ 519,113 $ 521,577 $ 524,078 $ 526,616 $ 529,192 $ 531,807 $ 534,461 $ 537,155 $ 539,890 $ 542,665 $ 545,482 $ 548,341 $ 551,244 $ 554,189 $ 557,179 $ 560,214 $ 563,294 $ 566,421 $ 569,594 $ 572,815 Present Worth of O&M Costs $ 519,113 $ 503,283 $ 487,958 $ 473,124 $ 458,762 $ 444,859 $ 431,398 $ 418,364 $ 405,745 $ 393,526 $ 381,695 $ 370,237 $ 359,142 $ 348,397 $ 337,991 $ 327,912 $ 318,150 $ 308,695 $ 299,536 $ 290,664 Cumulative PW of O&M Costs $ 519,113 $ 1,022,396 $ 1,510,355 $ 1,983,478 $ 2,442,241 $ 2,887,099 $ 3,318,497 $ 3,736,861 $ 4,142,606 $ 4,536,133 $ 4,917,827 $ 5,288,065 $ 5,647,207 $ 5,995,603 $ 6,333,594 $ 6,661,506 $ 6,979,656 $ 7,288,351 $ 7,587,887 $ 7,878,551
Net Profit Before Financing $ 873,023 $ 916,376 $ 960,946 $ 1,006,765 $ 1,053,864 $ 1,102,276 $ 1,152,034 $ 1,203,172 $ 1,255,726 $ 1,309,730 $ 1,365,222 $ 1,422,239 $ 1,480,821 $ 1,541,005 $ 1,602,832 $ 1,666,345 $ 1,731,584 $ 1,798,594 $ 1,867,418 $ 1,938,102 Present Worth of Net Revenue $ 873,023 $ 884,234 $ 894,718 $ 904,501 $ 913,606 $ 922,058 $ 929,879 $ 937,093 $ 943,720 $ 949,782 $ 955,298 $ 960,289 $ 964,773 $ 968,769 $ 972,294 $ 975,367 $ 978,003 $ 980,220 $ 982,031 $ 983,454 Cumulative PW of Net Revenue $ 873,023 $ 1,757,257 $ 2,651,975 $ 3,556,475 $ 4,470,081 $ 5,392,139 $ 6,322,018 $ 7,259,111 $ 8,202,831 $ 9,152,613 $ 10,107,910 $ 11,068,199 $ 12,032,972 $ 13,001,740 $ 13,974,035 $ 14,949,402 $ 15,927,405 $ 16,907,625 $ 17,889,656 $ 18,873,110